investigation of organic osmotic agents forward …epubs.surrey.ac.uk/808886/1/thesis _oct-...

Investigation of Organic Osmotic Agents

Forward Osmosis Desalination Process

Saleh O.M.AlAswad

A thesis submitted in fulfillment of the

Requirements of University of Surrey

For the degree of Doctor of Philosophy

This research program was carried out at the Centre for Osmosis Research and

Applications, Surrey University

September 2015

© Saleh O.M. AlAswad

Abstract

This study investigated the effects of novel membrane-osmotic agent systems on forward

osmosis process efficacy. Glucose, sucrose and NaCl were investigated as osmotic agents,

and nano-filtration, reverse osmosis and hollow fibre flat sheet as membrane types. The

effects of draw solution concentration, flow rate and temperature as well as feed water flow

rate and temperature were investigated for the aforementioned membrane types. The efficacy

of forward osmosis process was measured on the basis of water flux, water recovery rate,

water permeability, specific energy consumption and solute flux where applicable. Single,

binary and ternary systems were considered.

Experimental results showed the NF membrane and the RO membrane performed better at

low and high concentrations of osmotic agent. Higher water flux rates were achieved by

using NF membrane for both types of osmotic agents with changing osmotic agent

concentration rates. The best result was obtained through the combination of NF membrane

and glucose at lower concentrations and with sucrose at higher concentrations. The NF

membrane-sucrose system showed better results for all parameters when changing feed water

flow rate at lower temperatures while the NF-glucose system showed better performance

when increasing temperature of the solution. Specific energy consumption increased in all the

combinations of membrane-osmotic agent with increasing osmotic agent concentration rates.

However, the lowest energy requirements were noted for the combination of NF membrane -

glucose as an osmotic agent. Overall, both NF and RO membranes showed better results at

different osmotic agent flow rates, but glucose proved to be the superior osmotic agent.

For the binary systems, higher FO process efficacy across almost all parameters was noted

for the systems with deionised water (DW). However, systems also used more energy;

because such systems also used more energy they are not necessarily superior to other

systems. Water recovery rate and water flux were considerably higher in a ternary system

involving sucrose + NaCl + DW compared to the ternary system using brackish water as feed

solution in FO process. Overall, FO efficiency for a ternary mixture of glucose osmotic agent

+ NaCl salt + BW was higher in water recovery and water flux than the other ternary

mixtures with DW as feed solution.

Based on the study results, a number of recommendations for future work are provided.

Manipulated Osmosis Desalination (MOD) is a promising desalination approach that should

be further investigated by analysing the factors affecting the process’ efficacy. Such factors

are likely to include various membrane parameters (thickness, porosity and different pore

diameters) and draw solutions as well as the process parameters including temperature, flow

rates and osmotic agent concentrations. Further investigations should be conducted for binary

and especially ternary systems involving different types of membranes and solutes for

refining and optimising the process of selection. The Close loop circulation system needs to

be replaced by the no-circulation systems for better FO performance in future studies.

Importantly, the experimental results should be compared with the developed mathematical

models to further validate the results. Finally, future MOD studies should focus on

identifying optimal regeneration approaches. The selection of osmotic agents based on their

regeneration economics along with other important FO parameters could be undertaken in

future research.

Table of Contents

1. CHAPTER ONE: INTRODUCTION ............................................................................ 1

1.1. Research Background ..................................................................................................... 1

1.2. Overview of Desalination Processes ............................................................................... 2

1.2.1. Thermal Desalination .............................................................................................. 2

1.2.1.1. Multistage Flash ................................................................................................ 3

1.2.1.2. Multi-Effect Evaporation .................................................................................. 3

1.2.1.3. Thermal Vapour Compression .......................................................................... 3

1.2.1.4. Adsorption Vapour Compression ..................................................................... 3

1.2.2. Alternative Approaches .......................................................................................... 4

1.2.3. Membrane Approaches ........................................................................................... 4

1.2.3.1. Electrodialysis and Electrodialysis Reverse ..................................................... 5

1.2.3.2. Reverse Osmosis ............................................................................................... 5

1.2.3.3. Nanofiltration .................................................................................................... 7

1.2.3.4. Forward Osmosis .............................................................................................. 7

1.3. Research Rationale.......................................................................................................... 7

1.3.1. Manipulated Osmosis Desalination ........................................................................ 9

1.4. Research Contribution .................................................................................................. 11

1.5. Research Aims and Objectives ..................................................................................... 12

1.6. Thesis Organisation ...................................................................................................... 14

2. CHAPTER TWO: LITERATURE REVIEW .............................................................. 15

2.1. Water Desalination........................................................................................................ 15

2.1.1. Importance of Water Desalination ........................................................................ 15

2.1.2. Brief History of Water Desalination ..................................................................... 16

2.1.3. Current Status of Desalination Technologies ....................................................... 18

2.2. Types of Desalination ................................................................................................... 21

2.2.1. Thermal Approaches ............................................................................................. 21

2.2.1.1. Multistage Flash (MSF) .................................................................................. 22

2.2.1.2. Single Effect Evaporation (SEE) .................................................................... 22

2.3.1.3. Multi-Effect Desalination ............................................................................... 25

2.2.1.4. Freezing, Humidification and Solar Stills ....................................................... 25

2.2.2. Membrane Approaches ......................................................................................... 26

2.2.2.1. Electrical Driven Approaches ......................................................................... 26

2.2.2.2. Pressure Driven Approaches ........................................................................... 27

2.4. Forward Osmosis .......................................................................................................... 29

2.4.1. Process Description ............................................................................................... 29

2.4.2. Advantages of FO ................................................................................................. 32

2.4.3. FO Applications .................................................................................................... 33

2.4.3.1. Energy Applications........................................................................................ 33

2.4.3.2. Wastewater Treatment .................................................................................... 35

2.4.3.3. Pharmaceutical Applications .......................................................................... 38

2.4.3.4. Food Processing .............................................................................................. 38

2.4.3.5. Other Applications of FO................................................................................ 39

2.4.4. FO Applications for Water Desalination .............................................................. 39

2.4.4.1. Direct FO Desalination ................................................................................... 40

2.4.4.2. Indirect FO Water Desalination ...................................................................... 43

2.4.6. Membranes for Forward Osmosis Desalination ................................................... 45

2.4.6.1. Phase Inversion Cellulosic Membranes (PICM)............................................. 45

2.4.6.2. Thin Film Composite Membranes (TFCM) .................................................... 46

2.4.6.3. Chemically Modified Membranes (CMM) ..................................................... 49

2.4.6.4. Selection and Evaluation of FO Membranes .................................................. 49

2.4.7. Draw Solutions for Forward Osmosis .................................................................. 50

2.5. Concentration Polarisation ............................................................................................ 52

2.5.1. Introduction ........................................................................................................... 52

2.5.2. External Concentration Polarisation ..................................................................... 53

2.5.3. Internal Concentration Polarisation ...................................................................... 56

2.6. Manipulated Osmosis Desalination (MOD) ................................................................. 61

2.6.1. Process Description ............................................................................................... 61

2.6.2. Advantages and Considerations ............................................................................ 62

2.6.3. Existing MOD Projects ......................................................................................... 65

2.6.3.1. University of Surrey Rig ................................................................................. 65

2.6.3.2. Gibraltar MOD Pilot Plant .............................................................................. 66

2.6.3.3. Production Facilities in Oman ........................................................................ 67

2.7. Chapter Summary ......................................................................................................... 69

3. CHAPTER THREE: MATERIALS AND METHODS .............................................. 71

3.1. Introduction ................................................................................................................... 71

3.2. Flat Sheet Membrane Study .......................................................................................... 71

3.2.1. Experiment Outline ............................................................................................... 71

3.2.2. Experiment Set Up ................................................................................................ 72

3.2.3. Process Operation ................................................................................................. 74

3.2.4. Osmotic Agents ..................................................................................................... 76

3.2.5. Osmotic Agent Concentration Effect Study ......................................................... 76

3.2.6. Feed Water Flow Rate Effect Study ..................................................................... 77

3.2.7. Temperature Effect Study ..................................................................................... 78

3.3. Hollow Fine Fibre Membrane Study ............................................................................ 78

3.3.1. Experiment Outline ............................................................................................... 78

3.3.2. Experiment Set Up ................................................................................................ 79

3.3.3. Process Operation ................................................................................................. 81

3.3.4. Osmotic Agents ..................................................................................................... 83

3.3.5. Experimental Measurements ................................................................................. 85

3.3.6. Synthesis of Brackish Water and Seawater .......................................................... 89

3.4. Calculations of FO Performance Parameters ................................................................ 90

3.4.1. Water Flux ............................................................................................................ 90

3.4.2. Water Permeability ............................................................................................... 90

3.4.3. Recovery Percentage ............................................................................................. 91

3.4.4. Salt Flux ................................................................................................................ 91

3.4.5. Specific Energy Consumption .............................................................................. 91

3.4.6. Membrane Flow Rate Factor ................................................................................ 92

3.4.7. Net Driving Pressure ............................................................................................. 92

4. CHAPTER FOUR: RESULTS AND DISCUSSION .................................................. 93

4.1 Introduction .................................................................................................................... 93

4.1 NF Membrane Study ...................................................................................................... 94

4.1.1 The Effects of Osmotic Agents’ Concentrations ................................................... 94

4.1.1.1. Effect on the Water Flux ................................................................................. 94

4.1.1.2 Effect on the Water Recovery Percent ............................................................. 97

4.1.1.3. Effect on the Water Permeability .................................................................... 97

4.1.1.4 Effect on the Specific Energy Consumption .................................................... 98

4.1.1.5 Effect of OA Concentration on the Solute Flux............................................... 98

4.1.1.6. Discussion ....................................................................................................... 98

4.1.2 The Effects of Feed Water Flow Rate ................................................................... 99

4.1.2.1. Effect on the Water Flux ................................................................................. 99

4.1.2.2 Effect on the Water Recovery Percent ............................................................. 99

4.1.2.3 Effect on the Water Permeability ................................................................... 100

4.1.2.4 Effect on the Specific Energy Consumption .................................................. 100

4.1.3.3. Effect on the Solute Flux .......................................................................... 102

4.1.3.4. Discussion ................................................................................................. 102

4.1.4.The Effects of Osmotic Agents’ Flow rate .......................................................... 102

4.1.4.3. Effect on the Water Flux .......................................................................... 102

4.1.4.4. Effect of Recovery Percentage ................................................................. 103

4.1.4.5. Effect on the Water Permeability ............................................................. 103

4.1.3.4. Effect on Osmotic Agent Flow Rate on Specific Energy Consumption (SEC) ..

.................................................................................................................. 105

4.1.3.5 Effect on the Solute Flux ............................................................................... 105

4.1.3.6. Discussion ..................................................................................................... 105

4.1.4 The Effects of Feed Water Temperature .............................................................. 105

4.1.4.1. Effect on Water Flux ..................................................................................... 106

4.1.4.2. Effect on the Water recovery percentage ...................................................... 106




4.1.4.6. Discussion ..................................................................................................... 109

4.1.5 The Effects of Osmotic Agents’ Temperatures ................................................... 109

4.1.5.1. Effect on the Water Flux ............................................................................... 109

4.1.5.2 Effect on the Water Recovery Percent ........................................................... 111




4.1.5.6. Discussion ..................................................................................................... 112

4.1.6. Overview of the NF Membrane System Results ................................................ 112

4.2. RO Flat Sheet Membrane Study ................................................................................. 113

4.2.1. The Effects of Osmotic Agents’ Concentrations ................................................ 113


4.2.1.2. Effect on the Percent of Water Recovery ..................................................... 116

4.2.1.3. Effect on the Water Permeability .................................................................. 116

4.2.1.4. Effect on the Specific Energy Consumption ................................................. 116

4.2.1.5. Effect on the Solute Flux .............................................................................. 117

4.2.1.6. Discussion ..................................................................................................... 117

4.2.2. The Effects of Feed Water Flow rate .................................................................. 117


4.2.2.2. Effect on the Water Recovery Percentage .................................................... 119




4.2.2.6. Discussion ..................................................................................................... 120

4.2.3. The Effects of Osmotic Agents’ Flow rate ......................................................... 121



4.2.3.3. Effect on the Water Recovery Percentage (%R) ........................................... 123


4.2.3.5. Discussion ..................................................................................................... 123

4.2.4. The Effects of Feed Water Temperature ............................................................. 124

4.2.4.1. Effect on Water Flux ..................................................................................... 124





4.2.4.6. Discussion ..................................................................................................... 127

4.2.5. The Effects of Osmotic Agents’ Temperatures .................................................. 127


4.2.5.2. Effect on the Water Recovery Percent .......................................................... 129



4.2.5.5. Discussion ..................................................................................................... 130

4.2.6. RO Membrane Results Overview ....................................................................... 130

4.3. A Comparison Study of Membrane - Osmotic Agent Systems .................................. 131

4.3.1. Osmotic Agents’ Concentration .......................................................................... 131


4.3.1.2. Effect of OAs Concentration on Percent Recovery ...................................... 132

4.3.1.3. Effect on of OAs Concentration on Water Permeability .............................. 133


4.3.1.5. Discussion ..................................................................................................... 135

4.3.2. Feed Water Flow Rate ........................................................................................ 136





4.3.2.5. Discussion ..................................................................................................... 140

4.3.3. Osmotic Agents’ Flow Rate ................................................................................ 141




4.3.3.4. Effect of OAs Flow Rate on Specific Energy Consumption (SEC) ............. 144

4.3.3.5. Discussion ..................................................................................................... 145

4.3.4. Feed Water Temperature .................................................................................... 146

4.3.4.1. Effect of Feed Water Temperature on Water Flux ....................................... 146




4.3.4.5. Discussion ..................................................................................................... 150

4.3.5 Osmotic Agent Temperature ................................................................................ 151

4.3.5.1. Effect of Osmotic Agent Temperature on Water Flux.................................. 151



4.3.5.4. Effect of OA Temperature on Specific Energy Consumption (SEC) ........... 153

4.3.5.5. Discussion ..................................................................................................... 154

4.3.6. Generalised Results for the Membrane-Osmotic Agent Systems ....................... 155

4.4. Binary and Ternary OsmoticSystems ......................................................................... 156

4.4.1. Binary System with Deionised and Brackish Water as the Feeds ...................... 156




4.4.1.4. Effect of OPD on Specific Energy Consumption (SEC) .............................. 159

4.4.1.5. Effect of OPD on Solute Flux in sucrose and glucose Binary System ......... 160

4.4.1.6. Discussion ..................................................................................................... 161

4.4.2. Binary System with Seawater as the Feedwater ................................................. 162



4.4.2.3. Discussion ..................................................................................................... 164

4.4.3. Ternary Systems ................................................................................................. 165

4.4.3.1. Sucrose Ternary Systems and Water Recovery Percent ............................... 165

4.4.3.2. Sucrose Ternary Systems and Water Flux .................................................... 166

4.4.3.3. Glucose Ternary Systems and Water Recovery Percent............................... 168

4.4.3.4. Glucose Ternary Systems and Water Flux .................................................... 169

4.4.3.5. Discussion ..................................................................................................... 170

5. CHAPTER FIVE: CONCLUSIONS ......................................................................... 172

5.1. Introduction ................................................................................................................. 172

5.1. NF Membrane Study Outcomes.................................................................................. 173

5.1.1. Effect of Osmitic Agent Concentration Change ................................................. 173

5.1.2. The Effect of Feed Water Flow Rate Change ..................................................... 173

5.1.3. The Effect of Draw Solution Flow Rate Change ................................................ 173

5.1.4. The Effect of Feed Water Temperature Change ................................................. 173

5.1.5. The Effect of Draw Solution Temperature Change ............................................ 173

5.2. RO Membrane Study Outcomes ................................................................................. 174

5.2.1. The Effect of Draw Solution Concentration Change .......................................... 174

5.2.2. The Effect of Feed Water Flow Rate Change ..................................................... 174

5.2.3. The Effect of Draw Solution Flow Rate Change ................................................ 174

5.2.4. The Effect of Feed Water Temperature Change ................................................. 174

5.2.5. The Effect of Draw Solution Temperature on Flow Rate Change ..................... 174

5.3. Comparative Performance of NF and RO Membranes ............................................... 175

5.4. Outcomes for Binary and Ternary Systems for HFFM Membrane ............................ 176

6. CHAPTER SIX: FUTURE WORK ........................................................................... 179

7. References .................................................................................................................. 180

8. Appendix A1: Osmotic Agents Samples’ Analysis ................................................... 202

9. Appendix A2: Viscosity Meter TC-502 .................................................................... 206

10. Appendix A3: The Concentration of Mixture of Organic Osmotic Agent (Sucrose)

with Inorganic Osmotic Agent Sodium Chloride VersusDW/BW. ....................................... 207

11. Appendix A4:Combination of Organic Asmotic Agents Mixture of Sucrose and NaCl

versus DW/BW as Feed Water .............................................................................................. 209

12. Appendix A5: Combination of Organic Asmotic Agents Mixture of Glucose and NaCl

versus DW/BW as Feed Water .............................................................................................. 210

13. Appendix A6: Flat Sheet Unit Experimental Work for Osmotic Agent Sucrose andRO

Membrane .............................................................................................................................. 212

14. Appendix A7: Flat Sheet Unit Experimental Work for Osmotic Agent Glucose andRO

Membrane .............................................................................................................................. 214

15. Appendix A8: Flat Sheet Unit Experimental Work for Osmotic Agent Sucrose and NF

Membrane .............................................................................................................................. 216

16. Appendix A9: Flat Sheet Unit Experimental Work for Osmotic Agent Glucose and

NF Membrane ........................................................................................................................ 218

17. Appendix A10: Sample Calculations for the Data Parameters (Table 4.1) ............... 220

List of Figures

Figure 1.1. Electrodialysis Process (Fritzmann, et al, 2007) ..................................................... 5

Figure 1.2. Forward and Reverse Osmosis (Krukowski, 2001). Where: ∆P represents the

trans-membrane pressure differential; ∆π represents cross membrane osmotic pressure

differential. ................................................................................................................................. 6

Figure 1.3. Schematic Flow Diagram of the MOD Process. ..................................................... 9

Figure 1.4.MOD Laboratory Test Trial Plant Rig in Surrey University .................................. 13

Figure 2.1. Cumulative Installed Capacities of Desalination Plants: 1945-2005 (Gleick et al.

2006) ........................................................................................................................................ 17

Figure 2.2. RO-Based Desalination: Cost Evolution (Reddy and Ghaffour 2007) ................. 18

Figure 2.3. Desalination Technologies’ Share in Global Installed Capacities (Global Water

Intelligence, 2012) ................................................................................................................... 19

Figure 2.4. Desalination Technologies’ Share in the Middle East (Greenlee et al. 2009) ....... 19

Figure 2.5. Thermal and Membrane Desalination Approaches and Methods (adapted from El-

Dessoukey and Eltounney, 2000) ............................................................................................ 24

Figure 2.6. Multiple Effect Desalination Process (Al-Harbi et al. 2011) ................................ 25

Figure 2.7. Forward Osmosis Process (Thompson and Nicoll 2011) ...................................... 30

Figure 2.8. Water Flux Magnitude and Direction as a Function of Manually Applied

(Hydraulic) and Osmotic Pressures (adapted from Lee et al.1981) ......................................... 31

Figure 2.9. An Example of an OMBR (Achilli et al. 2009) ..................................................... 36

Figure 2.10. A Full Scale Working FO-Based Leachate Treatment System (Water

Development Report 2009) ...................................................................................................... 37

Figure 2.11. Direct Osmosis Process (adapted from Li et al. 2013) ........................................ 41

Figure 2.12. Indirect FO Process for Water Desalination (adapted from Li et al. 2013) ........ 44

Figure 2.13. Cross-Sectional Electronic Microscope Image of HTI’s Membranes (Zhao and

Zou 2011) ................................................................................................................................. 47

Figure 2.14. Cross-Sectional Electronic Microscope Image of a TFCM with Fingerlike Pores

(Yip et al. 2010) ....................................................................................................................... 48

Figure 2.15. ICP and ECP in FO Process (Zhao et al. 2012) ................................................... 53

Figure 2.16. Dilutive External, Concentrative External, Dilutive Internal and Concentrative

Internal Concentration Polarization (Mezher et al. 2011; Schiermeier 2008; Younos and

Tulou 2005) .............................................................................................................................. 57

Figure 2.17. Water Permeability Cefficient as a Function of Draw Solution Concentrations

(Mehta and Loeb 1979) ............................................................................................................ 58

Figure 2.18. Theoretical, FO Mode and PRO Mode Water Flux as a Function of net Osmotic

Pressure (Zhao and Zou 2011). ................................................................................................ 61

Figure 2.19. Continuous MOD Process with Regeneration Step (Adapted from Al Mayahi

and Sharif 2004) ....................................................................................................................... 62

Figure 2.20. Projected Power Consumption of MOD and RO Plants (adapted from Nicoll

2011) ........................................................................................................................................ 64

Figure 2.21. Projected Operating Costs of MOD and RO Plants (adapted from Nicoll 2011) 65

Figure 2.22. Gibraltar MOD Rig by Modern Water Prior to Containerisation (Thompson and

Nicoll 2011) ............................................................................................................................. 66

Figure 2.23. MOD Plant Facility in Al Khaluf, Oman (Thompson and Nicoll 2011) ............. 68

Figure 2.24. Normalised Flow for Al Khaluf RO Plant (above) and MOD Plant (below) in

2010 (Thompson and Nicoll 2011) .......................................................................................... 69

Figure 3.1. Experiment Outline for FO Using Flat Sheet Membranes .................................... 72

Figure 3.2.Design and layout of experimental membrane module for FO process ................. 73

Figure 3.3. Schematic Diagram of theFlat SheetForward Osmosis Bench-Scale Laboratory

Apparatus ................................................................................................................................. 74

Figure 3.4. Disassembled Membrane Testing Apparatus Showing Flow Channels. ............... 74

Figure 3.5. Experiment Outline for Hollow Fine Fibre Membrane ......................................... 79

Figure 3.6. The Experimental Unit Fixed onto aSkid. ............................................................. 80

Figure 3.7. Flow Diagram of The Hollow Fine Fibre Forward Osmosis Process ................... 81

Figure 3.8. The Experiment Hollow Fine Fibre Membrane, FO in Operation ........................ 82

Figure 3.9. Osmotic Pressure as a Function of concentration in sucrose-Sodium Chloride

Solution: a) Sodium Chloride concentration fixed and sucrose concentration Variable; b)

Sucrose concentration fixed and sodium chloride concentration variable. All values obtained

by OLI Analyzer software at 25°C. ......................................................................................... 87

Figure 4.1. Effect of osmotic pressure deferential (bar) on water flux (Jw) with constant

temperature (25 oC) and constant flow rates of feed and draw solutions (2 L/min) ................ 97

Figure 4.2. Effect of OA concentration on Water Flux with RO and NF Using RO and NF

Membrane at Constant Temperature (25 oC) and Feed and Draw Solution Flow Rates (2

L/min) .................................................................................................................................... 132

Figure 4.3. Effect of OAs Concentration on Percent Recovery Using RO and NF Membrane

at Constant Temperature (25 oC) and Feed and Draw Solution Flow Rates (2 L/min) ......... 133

Figure 4.4. Effect of OAs Concentration on Water Permeability Using RO and NF

Membranes at Constant Temperature (25 oC) and Feed and Draw Solution Flow Rates (2

L/min) .................................................................................................................................... 134

Figure 4.5. Comparison of Osmotic Agents’ Concentration Effect on Specific Energy

Consumption for NF and RO Membranes at Constant Temperature (25 oC) and Feed and

Draw Solution Flow Rates (2 L/min)..................................................................................... 135

Figure 4.6.Comparison of Feed Water Flow Effect on Water Flux for NF and RO Membranes

at Constant Temperature (25 oC) and OAs Concentration(275 and 200 g/L) for both sucrose

and glucose respectively ........................................................................................................ 137

Figure 4.7. Comparison of Feed Water Flow Effect on Water Recovery Percentage for NF

and RO Membranes at Constant Temperature (25 oC) andOAs Concentration (275 and 200

g/L) for both sucrose and glucose respectively...................................................................... 138

Figure 4.8. Comparison of Feed Water Flow Effect on Water Permeability for NF and RO

Membranes at Constant Temperature (25 oC) and OAs Concentration(275 and 200 g/L) for

both sucrose and glucose respectively ................................................................................... 139

Figure 4.9. Comparison of Feed Water Flow Effect on Specific Energy Consumption for NF

and RO Membranes at Constant Temperature and OAs Concentration(275 and 200 g/L) for


Figure 4.10. Comparison of Osmotic Agent Flow Effect on Water Flux for NF and RO

membranes at Constant Temperature (25 oC) and OAs Concentration(275 and 200 g/L) for


Figure 4.11. Comparison of Osmotic Agent Flow Effect on Water Recovery Percent for NF

and RO membranes at Constant Temperature (25 oC) and OAs Concentration (275 and 200

g/L) for both sucrose and glucose respectively...................................................................... 143

Figure 4.12. Effect of OAs Flow Rate on Water Permeability Using RO And NF Membranes


and glucose respectively ........................................................................................................ 144

Figure 4.13. Effect of OAs flow rate on SEC using NF and RO membrane at Constant

Temperature (25 oC) and OAs Concentration(275 and 200 g/L) for both sucrose and glucose

respectively ............................................................................................................................ 145

Figure 4.14. Effect of Feed Water Temperature on Water Flux Using NF and RO Membranes

at Constant Feed and Draw Solution Flow Rate (2 L/min) and OAs concentration (200 and

250 g/L) for both sucrose and glucose respectivly ................................................................ 147

Figure 4.15. Effect of Feed Water Temperature on Percent Recovery Using NF and RO

Membranes at Constant Feed and Draw Solution Flow Rate (2 L/min) and OAs concentration

(200 and 250 g/L) for both sucrose and glucose respectivly ................................................. 148

Figure 4.16. Effect of Feed Water Temperature on Water Permeability Using NF and RO


(200 and 250 g/L) for both sucrose and glucose respectivly ................................................. 149

Figure 4.17. Effect of Feed Water Temperature on Specific Energy Consumption Using NF

and RO Membranes at Constant Feed and Draw Solution Flow Rate (2 L/min) and OAs

concentration (200 and 250 g/L) for both sucrose and glucose respectivly .......................... 150

Figure 4.18. Effect of OA Temperature on Water Flux Using RO and NF Membranes at

Constant Feed and Draw Solution Flow Rate (2 L/min) and and OAs concentration (200 and


Figure 4.19. Effect of OA Temperature on Percent Recovery Using NF and RO Membranes

at Constant Feed and Draw Solution Flow Rate (2 L/min) and and OAs concentration (200

and 250 g/L) for both sucrose and glucose respectivly .......................................................... 152

Figure 4.20.Effect of OA Temperature on Water Permeability Using NF and RO Membranes



Figure 4.21. Effect of OA Temperature on Specific Energy Consumption Using NF and RO

Membranes at Constant Feed and Draw Solution Flow Rate (2 L/min) and and OAs

concentration (200 and 250 g/L) for both sucrose and glucose respectivly .......................... 154

Figure 4.22. Osmotic Pressure Differential Effect on Water Flux in a Binary System ......... 157

Figure 4.23. Osmotic Pressure Differential on Percent Recovery in Binary Sucrose and

Glucose System ...................................................................................................................... 158

Figure 4.24. Osmotic Pressure Differential on Water Permeability in Binary Sucrose and

Glucose System ...................................................................................................................... 159

Figure 4.25. Osmotic Pressure Differential Effect on SEC of Sucrose and Glucose OAs in a

Binary System ........................................................................................................................ 160

Figure4.26. Osmotic Pressure Differential Effect on Solute Fluxin Binary Sucrose and Binary

Glucose Systems .................................................................................................................... 161

Figure 4.27. Comparison of Osmotic Pressure Differential Effect on Water Flux in Binary

Osmotic Systems and Seawater as the Feed Water ............................................................... 163

Figure 4.28. Comparison of Osmotic Pressure Differential Effect on Water Recovery Percent

in Binary Osmotic Systems and Seawater as The Feed Water .............................................. 164

Figure 4.29. Effect of Osmotic Pressure on Water Recovery Percent in a Ternary System

(sucrose + NaCl /H2O ) .......................................................................................................... 166

Figure 4.30. Effect of Osmotic Pressure on Water Flux in a Ternary System (sucrose + NaCl /

H2O ) ...................................................................................................................................... 168


(glucose + NaCl / H2O ) ........................................................................................................ 169


(glucose + NaCl /H2O ) .......................................................................................................... 170

Figure 8.1. Calibration Curve for Glucose............................................................................. 202

Figure 8.2. Calibration Curve for Sucrose ............................................................................. 204

Figure 9.1. Viscosity Meter TC-502 ...................................................................................... 206

List of Tables

Table 1.1. Approaches to Seawater Desalination (Van der Vegt et al., 2011) .......................... 2

Table 2.1. Description of Pressure Driven Membrane Processes Source : Mortazvi (2008) .. 29

Table 2.2. Comparison of FO and RO (Adapted and Expanded from Chung et al. 2012; Zhao

et al. 2012) ............................................................................................................................... 63

Table 3.1. TFC-SR2 Specifications (Koch Membrane Systems Ltd ,UK) .............................. 75

Table 3.2. TFC-ULP specifications (Koch membrane systems Ltd ,UK) ............................... 75

Table 3.3. Physico-Chemical Properties of Sucrose and Glucose ........................................... 76

Table 4.1. Effect of Osmotic Agent Concentration on Water Flux, Recovery Percentage,

Water Permeability, Specific Energy Consumption and Solute Flux ...................................... 95

Table 4.2.Effect of Osmotic Pressure Deferential (Bar) on Water Flux (Jw)........................... 95

Table 4.3.Effect of Osmotic Agents’ Concentration on Viscosity .......................................... 96

Table 4.4. Effect of OAs Viscosity on Water Flux .................................................................. 96

Table 4.5.Effect of Feed Water Flow Rate on Water Flux, Recovery Percentage, Water

Permeability, Specific Energy Consumption and Solute Flux............................................... 101

Table 4.6. Effect of Osmotic Agent Flow Rate on Water Flux, Recovery Percentage, Water

permeability, Specific Energy Consumption and Solute Flux ............................................... 104

Table 4.7. Effect of Feed Water Temperature on Water Flux, Recovery Percentage, Water

permeability, Specific Energy Consumption and Solute Flux ............................................... 107

Table 4.8. Effect of OA Temperature on Water Flux, Recovery Percentage, Water


Table 4.9. Results of Changes of Manipulated FO Factors on Key Measured Parameters of

the Process: NF Membrane .................................................................................................... 113

Table 4.10. Effect of Osmotic Agent Concentration on Water Flux, Recovery Percentage,

Water permeability, Specific Energy Consumption and Solute Flux .................................... 115

Table 4.11. Effect of Osmotic Pressure Differential on Water Flux ..................................... 115

Table 4.12. Effect of Feed Water Flow Rate on Water Flux, Recovery Percentage, Water


Table 4.13. Effect of Osmotic Agent Flow Rate On Water Flux, Recovery Percentage, Water


Table 4.14. Effect of Feed Water Temperature on Water Flux, Recovery Percentage, Water


Table 4.15. Effect of Osmotic Agent Temperature on Water Flux, Recovery Percentage,

Water Permeability, Specific Energy Consumption and Solute Flux .................................... 128

Table 4.16. Results of Changes of Manipulated FO Factors on Key Measured Parameters of

the Process: RO Membrane ................................................................................................... 131

Table 4.17. The Best Combinations of Membranes and Omsotic Agents Based on Experiment

Factors and Parameters .......................................................................................................... 155

Table 8.1. Results of the Analyses in mg/ml Glucose ........................................................... 203

Table 8.2. Results of the Analyses in mg/ml Sucrose............................................................ 205

Table 10.1. The Concentration of Sucrose as OA with NaCl vs DW/BW as Feedwater ...... 207

Table 11.1. Sucrose and NaCl vs DW/BW as Feedwater ...................................................... 209

Table 12.1. Glucose and NaCl vs DW/BW as Feedwater ..................................................... 210

Table 13.1. Results for The System of Flat Sheet RO Membrane and Sucrose as OA ......... 212

Table 14.1. Results for The System of Flat Sheet RO Membrane and Glucose as OA ......... 214

Table 15.1. Results for The System of NF Membrane and Sucrose as OA ........................... 216

Table 16.1. Results for The System of NF Membrane and Glucose as OA .......................... 218

Table 17.2. Sample Calculations for the Data Parameters in Table 4.1. with Sucrose OA ... 225

Preface

The purpose of this research was to investigate the effect of different organic osmotic agents

in forward osmosis desalination process. The experimental apparatus was constructed at the

Centre for Osmosis Research and Applications (CORA), University of Surrey, UK to carry

out the experiments. The investigated osmotic agents were sucrose, glucose and sodium

chloride (NaCl). The experiments were conducted for several types of membranes: nano-

filtration (NF), reverse osmosis (RO) and hollow fine fibre membrane used as flat sheets and

modules. The study also investigated binary and ternary combinations involving different

organic osmotic agents. The manipulated factors in the experiments included osmotic agent

concentrations, flow rates and temperatures as well as feed water flow rates and temperatures.

The parameters measuring efficacy of the forward osmosis process were water flux, water

permeability, recovery percentage, solute flux and specific energy consumption where

applicable.

The research is contributed by definitions, theoretical concepts and practically conducted

experiments. All experiments were carried out in laboratory settings and on pilot scales using

calibrated measurements. The details of the experiments are presented in form of images,

figures, graphs and equations.

The study addressed an interesting challenge in the field of water desalination by

investigating forward osmosis process and attempting to develop optimal combinations of

membranes and osmotic agents to enhance the efficacy of forward osmosis desalination

process.

Chapter 1 describes the study background by reviewing the key water desalination

approaches, providing general outlook of the study and describing the rationale, main goals

and objectives of the research.

Chapter 2 presents a comprehensive literature review on desalination, with a focused view

on forward osmosis approach, its main principles and challenges.

Chapter 3 offers a review of the study methods including the experiment design and

procedures, quantitative aspects behind measures and detailed description of the research

elements.

Chapter 4 presents experimental results obtained during the study.

Chapter 5 presents the discussion of the results and offers conclusions based on the study

results.

Chapter 6 offers directions for future work.

Acknowledgements

In the name of Allah, the Most Gracious and the Most Merciful. Alhamdulillah, all

praises to Allah for the strengths and His blessing in completing this Thesis. First of all , I

would like to thank my mother, Moody Alfreedi , without her support and encouragement I

never would have been able to achieve my goals.

Secondly, I sincerely pay thanks and gratitude to my supervisor, Dr. Mohammed

Sanduk, Dr. Sami Al-ibi, Dr.Esat Alpay and Professor Adel Sharif for their kind supervision,

encouragement and constant support in completing this study. I wish to gratefully

acknowledge the special thanks to Professor Rex Thorpe and Mr. David Hawkins and Mrs.

Hilary Mitchell.

Finally , I would like to thank my lovely wife Badryah Aljourboo for supporting me

in this long journey of my life and love me so much. My love and thanks go also to my

beautiful angels, Najla , Rema , Khalid ,Saud for their understanding of why I wasn’t at home

as much as should like to have been.

Author’s declaration

I hereby declare that I (Saleh O. M. AlAswad) am the author of this thesis. I authorize the

University of Surrey to lend this thesis and to imitate in part or total at the request to other

institutions for the purpose of scholarly research

Sign. ........................................................................

Nomenclature

Abbreviations

Description

ABVC absorption vapour process

ADVC Adsorption vapour compression

AL Active layer

Al2(SO4)3 Aluminium Sulphate

C12H22O11 Sucrose

C3H6O3 cellulose acetate and lactic acid

C4H4O4 Maleic acid

C6H12O6 Glucose

CA cellulose acetate

Ca(OH)2 Calcium hydroxide

CaSO4 Calcium sulphate

CD Capacitive deionisation

CECP Concentrative external concentration polarization

CICP Dilutive internal concentration polarization

CMM Chemically modified membranes

CO2 Carbon dioxide

CORA Centre for osmosis research and applications

CVC Chemical vapour compression

DECP Dilutive external concentration polarization

DICP Concentrative internal concentration polarization

DO Direct Osmosis

DS Draw solution

ECP External concentration polarization

ED Electrodialysis

FO Forward osmosis

FS Feed solution

GAC Granular activated carbon

HTI Hydration technology innovation

ICP Internal concentration polarisation

LiBr Lithium bromide

LPRO Low pressure reverse osmosis

MED multi effect distillation

MEE Multi-effect evaporation

MENA Middle East and North Africa

MJ Minecraft joules

MOD manipulated osmosis desalination

MOD Manipulated osmosis desalination

MPD m-phenylenediamine

MSF multistage flash

MVC Mechanical vapour compression

Na2SO4 Sodium sulphate

NaCl Sodium chloride

NCA Non cellulose acetate

NDP Net driving pressure

NH3 Ammonia

NH4HCO3 Ammonium bicarbonate

OA Osmotic agent

OMBR Osmotic membrane bioreactor

PBI Polybenzimidazole

PDS Hydraulic pressure for draw solution

PES Polyethersulfone

PFW Hydraulic pressure for feed water

pH Negative logarithmic value of the hydroxonium ion (H3O+) concentration

PICM Phase inversion cellulosic membranes

PRO Pressure Retarded Osmosis

PS Polysulfonic

PSI Pound per square Inch

PVDF Polyvinylideneflouride

PVP Polyvinylpyrrolidone

RO Reverse osmosis

SEC Specific energy consumption

SEE single effect evaporation

SL Support layer

SO2 Sulfur dioxide

TDS Total dissolved solids

TFC Thin film composite

TFCM Thin film composite membranes

TFC-ULP Thin film composite ultra-low-pressure

TMC Trimesoyl chloride

TVC Thermal vapour compression

UCL University of California Los Angeles

VC vapour compression

ZnCl2 Zinc chloride

Symbols

Symbol Description A Water permeability Constant

Aw Membrane’s permeability coefficient

A Water permeability coefficient

Cb Concentration of the feed solution in the bulk Cm Concentration of Feed Solution at Membrane

D Solute’s diffusion coefficient

Dh Hydraulic diameter

E Specific energy consumption

G Acceleration

Js Salt flux diffusion

Jv Total volumetric

Jw Water flux K Mass transfer Coefficient

k Mass transfer

kdraw Mass transfer coefficient on the draw side

kfeed Mass transfer coefficient on the feed side K Mass transfer Coefficient

Ρ Turbine efficiency

PDin Absolute pressure of the draw solution (osmotic agent) entering the membrane

PFin Absolute pressure of the feed water entering the membrane

Pn Net driving pressure

P Solution pressure

Q Flow

QFin Volumetric flow rate of feed water entering the membrane

QFout Volumetric flow rate of feed water leaving the membrane

% R Recovery percentage

Sh Sherwood number

S Membrane’s structural parameter

Total power generated

Greek Symbol Description

δ Constrictivity

ε Porosity

εeff Effective transport through porosity μw Water Chemical Potential

∆π Osmotic pressure differential

∆πeff Effective driving force

π Osmotic Pressure

πbdraw Osmotic pressure of the draw solution in the bulk

πbfeed Osmotic pressure of the feed solution in the bulk

DS Osmotic pressure in bulk DS side

FW Osmotic pressure in bulk FW side

πmdraw Osmotic pressure of the draw solution at the membrane

πmfeed Osmotic pressure of the feed solution at the membrane

DS-m Osmotic pressure at the membrane interface in DS side

FW-m Osmotic pressure at the membrane interface in FW side

i Osmotic pressure at the transition interface between membrane layers

ρ Water density

Reflection Coefficient

Thickness

τ Tortuosity

Units

Description

g/L Gram per litre

psi Pascal

mg/L Milligram per litre ᴼC Degree Celsius

Note: Solute flux is used instead of sugar flux (sucrose or glucose solution) in the thesis in

various experimental runs.

Saleh Al Aswad Page 1

1. CHAPTER ONE: INTRODUCTION

1.1. Research Background

Insufficient supply of fresh water is gradually becoming one of the most pressing global

issues. The continuing population growth, industrialisation and agriculture activities are

creating increasing demand on the naturally available amounts of drinking water resources,

while pollution and instability of weather patterns decreases the chances of these resources’

recuperation. As a result, nearly one third of the global population today lives in the water-

stressed regions, and by 2025 this number is expected to double (Service, 2006).

In order to counter the imminent shortages of fresh water supply, a number of measures have

been suggested, including water conservation techniques, building of better infrastructure and

water distribution systems. The problem with these approaches, however, is that they are only

able to reduce the use of the available water sources, not offer alternative ways to water

production. From this perspective, the new, unconventional approaches to extract fresh water

can be considered as an essential avenue to expand the existing water resources and resolve

many issues that the lack of fresh water involves (Elimelechand Phillip 2011; Shannon et al.

2008).

One of the most promising alternative methods to produce fresh water is desalination.

Desalination refers to the process of extracting/reducing the amount of solids, minerals, and

organic matter in water to make it suitable for consumption and irrigation (Porteous 1983).

Typically, the desalination process uses either brackish water or seawater as feed water and

produces two streams: one with low concentration of salts and minerals for potable purpose

and the other one with high concentration of salts and minerals, which are either disposed of

or recycled again to receive a pure water flux across the selectively permeable FO and RO

membranes (Lewis 1982). While desalination processes using both feeds are common, it is

the seawater desalination which allows for a virtually unlimited supply of usable water, since

seawater represents 97% of available water on Earth (Kim et al. 2010). The potential of

seawater desalination has been recognised in many water stressed regions of the world.

Within the past several decades, desalination plants working on seawater feed appeared in

Israel, Spain, and a number of Middle Eastern countries (Dreizin et al. 2008; Lee 2010;

Schermeier 2008).


The growth in popularity of seawater desalination plants has been caused by the

technological progress and appearance of commercially available materials to support the

process. The early plants were primarily based on thermal desalination approach, which was

energy demanding and not ecologically feasible (Elimelechand Phillip 2011; Mezher et al.

2011). Since then, thermal processes have been greatly improved, while new approaches to

seawater desalination have emerged as well, which were based on novel approaches to

separate water from different types of solutes.

1.2. Overview of Desalination Processes

Modern theory and practice of desalination recognizes more than a dozen approaches, which

are listed within three main categories in Table 1.1. The detailed discussion of each of these

approaches is beyond the scope of this report; therefore, each process is discussed in general

terms. The exception is the membrane technology approaches, which are the focus of the

current study.

Table ‎1.1. Approaches to Seawater Desalination (Van der Vegt et al., 2011)

Approaches Based on Thermal

Technologies

Approaches Based on

Membrane Technologies

Approaches Based on

Alternative Technologies

Adsorption Vapour Compression

Mechanical Vapour Compression

Multi-Effect Evaporation

Multistage Flash

Solar Distillation

Thermal Vapour Compression

Electrodialysis

Electrodialysis Reversal

Forward Osmosis

Reverse Osmosis

Capacitive Deionisation

Electrodeionisation

Freeze Separation

Gas Hydrates

Ion Exchange

Rapid Spray Evaporation

Vacuum Distillation

1.2.1. Thermal Desalination

Thermal desalination is the oldest and the most widespread approach to water desalination

(Mezher et al. 2011). In simplest terms, the approach assumes boiling the feed water to purify

and collect the vapour. Thermal desalination is a rather energy demanding process, which has

been widely applied in the Middle East, where the cost of petrochemicals is rather low.


1.2.1.1. Multistage Flash

Multistage flash represents a process made of several flashing spaces called stages, each stage

containing an exchanger to heat the feed water and a collector for condensate steam. The

entire sequence has a cool and a hot end, and the stage temperatures between them increase.

Each stage has a different pressure level, which matches the boiling points of water solution

at that stage’s temperature. Feed water travels through each stage, where it is heated by

warmer water vapour. By reaching the final stage-the brine heater-the water solution has

achieved a maximum temperature and flows back through the stages, where certain amount

of it evaporates through the boiling (flashing) process. The evaporating steam is hotter than

the feed water; therefore, it cools down and condenses on the heat exchangers and heats the

feed water (Younos and Tulou 2005).

1.2.1.2. Multi-Effect Evaporation

Multi-effect evaporation is, according to its name, a combination of several effects. The feed

water gets sprayed over the hot tubes, where it evaporates, and the vapour is collected to flow

through the tubes in the next effect, thus using the energy from each stage to heat and

evaporate water at the next stage. With each passing effect, the remaining amount of brine is

collected at the bottom and either passes onto the next stage or gets disposed of. This

approach is, actually, less energy demanding than multistage flash while still being able to

handle large energy capacities (El-Dessouky and Eltounney 2002).

1.2.1.3. Thermal Vapour Compression

This approach is a somewhat improved multi-effect evaporation, where the steam jet ejector

compresses the vapour which enters the first effect tubes. The vapour gets condensed by a

condenser to the final product. An alternative way to compress the vapour is by mechanical

compressors instead of jet ejectors. It is considered that thermal vapour compression has a

higher performance ration in comparison to the multi-effect evaporation (El-Dessouky et al.

2000).

1.2.1.4. Adsorption Vapour Compression

In the essence of this approach is creation of pressure differences between the two reservoirs,

where feed water is transferred. Pressure differences are caused by an exothermic reaction,

where the feed water is preheated by being mixed with a special solution like LiBr. This

approach typically generates pot water (El-Dessouky and Eltounney 2002).


1.2.2. Alternative Approaches

Among the alternative approaches to desalination, ion exchange, is perhaps, the one receiving

most of attention. Ion exchange systems involve interchange of ions between a solid phase

and a surrounding it liquid phase. The solid phase consists of chemical resins, which come in

forms of synthetic or naturally-occurring inorganic materials. The liquid phase consists of

feed water. Desalination through ion exchange is a rather complex process, where salt ions

(such as Na+ and Cl

-) are replaced for other ions when feed water passes through the solid

phase beads (Kim et al. 2010).

In addition to ion exchange, which is already being used in practice, a number of additional

technologies are emerging, which may have some potential for desalination in the future. One

of these approaches is electrodeionisation, which, in effect, is a combination of

electrodialysis and ion exchange. This approach assumes applying electric charges over a

feed water stream which goes through a series of membrane pairs. The membranes are either

anion or cat-ion transferring, but not both. As a result, two streams of liquid are created: one

of brine and one of pure water. Due to the nature of purification process, electrodeionisation

can produce extremely pure waters (Younos and Tulou 2005).

1.2.3. Membrane Approaches

Membrane technologies can be considered a successor to traditionally used thermal

desalination approaches: the vast majority of contemporary desalination plants, as well as

those being planned utilize membranes (Elimelech and Phillip 2011; Fritzmann, et al. 2007).

Membranes represent a thin layer of porous material, which permeates water molecules

through, but holds large molecules, viruses, bacteria and different solutes. They may be made

of different polymeric and non-polymeric materials, although the most commonly used

membranes for desalination are synthetic (Younos and Tulou 2005).

Based on the technological process used for desalination, all membrane based approaches can

be classified as either electrical or pressure driven. Electrical driven approaches include

electrodialysis and electrodialysis reversal, while pressure driven approaches include reverse

and Forward Osmosis, nano-, ultra-, and microfiltration. All these approaches are briefly

described below.


1.2.3.1. Electrodialysis and Electrodialysis Reverse

Electrodialysis is a desalination approach relies on membranes and electrodes adjacent to

them. The separation of water from the particles dissolved in it occurs within membrane pair,

which includes an anion membrane, a cation membrane, and two spacers, which evenly

distribute the feed water flow across the membrane surface (Brunner, 1990). An example of a

complete electrodialysis schema is shown in Figure 1.1. In the picture, the ED stack is

separated into multiple cells by anion and cation exchange membranes labelled AEM and

CEM respectively. The alternating sequence of membrane placement allows to reduce ionic

species in the diluate compartments of the stack and increase them in the concentrate

compartments (Fritzmann et al. 2007). The electric circuit is closed in the first and last cells

of the ED stack. ED process is quite effective with removal of salts. As a rule, the process

requires 70-90 psi pressure and is able to remove 75% to 98% of solids from feed water

(Younos and Tulou 2005).

Figure ‎1.1. Electrodialysis Process (Fritzmann, et al, 2007)

1.2.3.2. Reverse Osmosis

Reverse Osmosis is the passage of liquid from the region of high concentration to the region

with low concentration through a one-way permeable membrane. Consequently, Reverse

Osmosis is the opposite of that: water travels from the regions with high concentration to the

region of low concentration under artificially applied pressure. The difference between

regular and Reverse Osmosis is demonstrated in Figure 1.2. Applied to the process of


desalination, Reverse Osmosis separates water from salt from the seawater feed, as the salts

and minerals are held by membrane pores, which allow only water molecules through. The

concentrated brine left after the water passage is normally disposed of.

Figure 1.2. Forward and Reverse Osmosis (Krukowski, 2001). Where: ∆P represents

the trans-membrane pressure differential; ∆π represents cross membrane osmotic

pressure differential.

The Reverse Osmosis is much more powerful than electrodialysis, as it allows to effectively

remove TDS concentrations up to 45,000 mg/L, which is sufficient to purify both salt and

brackish waters. At the same time, Reverse Osmosis requires energy to operate the pumps

pressing feed water through the membranes, although the amount of energy is generally less

than under thermal desalination approaches. The required pressure levels, however, rise with

the constantly increasing feed water concentration (up to 1,200 psi for seawater), thereby

giving Reverse Osmosis only 30%-80% recovery rates (Younosand Tulou 2005).

Membranes play a crucial role in Reverse Osmosis desalination process. In general, all

membranes can be classified into cellulose acetate (CA) and non-cellulose acetate (NCA).

CA membranes have smooth surfaces, which are resistant to fouling, because the particles

causing fouling do not deposit in the membrane crevices (Nicolaisen 2002). NCA membranes

are commonly referred to as “thin film composite membranes.” They use organic materials,

such as polysulfone. The advantages of NCA are higher flux rates and relative stability over a

long range of pH levels; however, they are quite sensitive to the effects of chlorine (El-

DessoukyandEttouney 2002).

Protection of membranes in Reverse Osmosis and other membrane-based desalination

approaches is necessary to enhance retention of salts and reduce cost of energy. Therefore,

feed water pre-treatment is typically required to remove very large particles, bacteria, oil, and

organic matter that may damage membranes. Pre-treatment also involves addition of


chemicals to avoid membrane scaling and the formation of precipitates as well as adjustment

of pH levels depending on the membrane type. Microfiltration and ultrafiltration are typical

pre-treatment approaches used in Reverse Osmosis. A more detailed overview of various

membrane types used for osmotic desalination process is provided in the Literature Review

section of the report.

1.2.3.3. Nanofiltration

Nanofiltration, in essence, has the same concept of work as Reverse Osmosis. The only

difference is in lower pressure levels required, because of larger pore sizes of the membranes.

Nano-filtration is used to soften water and remove larger solids and dissolved organic carbon.

Although sometimes applied as a standalone technology for low TDS brackish water feeds,

nano-filtration, nevertheless, cannot be effectively used to treat seawaters (Younos 2005). At

the same time, research has shown potential in application of nano-filtration in combination

with Reverse Osmosis approach (Al-Zuhairi 2008; Peng et al. 2004).

1.2.3.4. Forward Osmosis

Forward Osmosis is the direct osmotic approach, where liquid naturally travels from the areas

of low concentration to the areas of high concentration. Just like Reverse Osmosis, Forward

Osmosis approach has been considered for desalination since the 1970s. However, the earlier

studies remained largely theoretical, with few maturing into practical operational systems

(McCutcheon et al. 2005; Zhao et al. 2012). The interest in Forward Osmosis has grown

recently due to development of the commercially available membranes. Still, unlike Reverse

Osmosis, Forward Osmosis remains a poorly researched approach when it comes to

desalination (Cath et al. 2006; Zhao et al. 2012). At the same time, the available studies

confirmed Forward Osmosis potential as a more economically feasible and environmentally

sustainable desalination approach (Kravathand Davis 1975; McCutcheon et al. 2005; Peng et

al. 2004; Sharif and Al-Mayahi 2005).

1.3. Research Rationale

As summarised above (section 1.2), a number of desalination approaches have been presented

in the literature: multistage flash distillation (MSF), multi effect distillation (MED), vapour

compression (VC), reverse osmosis (RO), electrodialysis (ED), Capacitive Deionisation (CD)

and others (Ghaffour et al. 2013; Greenle et al. 2009; Fritzman et al. 2007; Kim et al. 2010;

Mezher et al. 2011; Younosand Toulou 2005). Some of these approaches, like RO MED, and


MSF were successfully implemented in practice. Despite their commercial applications,

however, contemporary desalination methods are still far from ideal in both economical and

environmental terms. For example, thermal desalination approaches like MSF and MED are

economically burdensome due to enormous amounts of energy they require. The existing

popular membrane approaches like RO are more efficient in this regard, although they

provide relatively low recovery rates. Finally, novel approaches like electrodeionisation may

offer advantages of a combination of methods; however, they are still in research and

development stages, with little practical applications to confirm their superiority over the

existing seawater desalination techniques. Therefore, the search for more efficient and

effective means of water desalination continues.

FO is getting increasing attention by researchers and practitioners in the desalination area due

to the many advantages it may provide. One of the major benefits that distinguish it from the

dominant membrane based approaches is the balance between the required amount of energy

and the purity of the produced water. Further, because FO represents a naturally occurring

process of water separation, it requires significantly less energy than RO (McGinnis and

Elimelech 2007). This means better economic efficiency. At the same time, the membranes

used in FO process have the same density than the ones used in RO, which means

comparative levels of water purity. There are also a number of additional contaminants that

are held in the FO process (Cartinella et al. 2006; Cath et al. 2010). Another potential benefit

of FO approach application in seawater desalination is membrane durability. Several recent

studies demonstrated that the degree of fouling in FO is lower, more reversible, and it can be

further minimised by hydrodynamics optimisation (Achilliet al. 2009; Mi and Elimelech

2010; Lee et al. 2010). Finally, due to high levels of osmotic pressure gradient across the

membrane, FO may be able to provide higher levels of water fluxes and, consequently, higher

degrees of water recovery (Zhao et al. 2012). This, in turn, will lead to reduction of the brine

volumes, which so far has been the major environmental issue related to desalination plants

(McCutcheon et al. 2005).

There is growing research on desalination that focuses on application of hybrid (combined)

systems. Optimal combination of different desalination approaches may offer such benefits as

significant cost reductions, reduction of energy use and better handling of the discharge

materials (Al-Mutaz 2003; Hamed 2005; Gude et al. 2010). Ghaffour et al. (2013) showed

that hybrid water desalination systems are especially attractive in places with considerable

fluctuations of power and water demands. Various combinations of desalination techniques


have been proposed, the majority of them include membrane and thermal methods: MSF/RO

and MED/RO (Gude et al. 2010). This research follows the call for more research on the

combination approaches to seawater desalination. The study takes Forward Osmosis as a

basis for the manipulated osmosis desalination (MOD) process.

1.3.1. Manipulated Osmosis Desalination

The Manipulated Osmosis Desalination (MOD) process was developed at the University of

Surrey’s Centre for Osmosis Research and Applications (CORA). The process is comprised

of FO desalination combined in a single cycle with a regeneration step. Specifically, the

principle behind the MOD is manipulation of two fluids with different osmotic pressures to

receive a pure water flux across the selectively permeable FO and RO membranes (Al-

Mayahiand Sharif 2004). The manipulated fluids are seawater and a chosen draw solution.

The seawater passes through a FO unit with a draw solution, which attracts water molecules.

The resulting liquid is then transferred to a membrane-based (RO or nanofiltration) recovery

unit where final separation of the product water takes place. Figure 1.2 presents a basic

schema of the MOD process using RO unit for the recovery step.

Figure ‎1.3. Schematic Flow Diagram of the MOD Process.

In comparison with the traditional RO-based desalination process that operates at 60-80 bar,

the MOD operates at greatly reduced pressures of 2-3 bar at the feed water (seawater) part

(see Figure 1.3). While RO as a recovery step consumes additional energy, the consumption

can be minimised with a careful selection of the osmotic agent and optimisation of operative

conditions. Additionally, these factors can lead to greater recovery efficiency due to control


of the draw solution composition and absence of foulant impurities. Other advantages of the

MOD process are lower costs due to reduced capital expenses due to significant reduction of

contaminants; higher quality water product due to a double membrane barrier; and possibility

of using lower pressure pipe works and fittings (Thompson and Nicoll 2011).

Another important advantage of MOD is lower membrane fouling in comparison with the

traditional desalination methods. The membrane fouling involves the deposition and

adsorption of feed water constituents, such as organic and inorganic compounds, colloidal

particles, and microbes to the membrane surface. As water permeates the semipermeable

membrane, foulants in the feed water solution accumulate on the membrane surface forming

a thin layer which creates hydraulic resistance. This layer enhances concentration polarization

that reduces the net driving force for permeation. The fouling deteriorates membrane

performance and decreases membrane life span resulting in high operating cost for membrane

desalination. As FO is a low fouling process, the fouling layer formed in FO is structurally

different from fouling in pressure-driven processes (Hoek and Elimelech 2003; Lee et al.

2010).

Normally, the MOD process depends on the performance of FO membranes, composition,

concentration and type of draw solution, recirculation rate of the draw solution, performance

of the regeneration step and feed water temperature, composition and flow rate. An important

aspect to consider in any pressure-driven membrane desalination process is concentration

polarisation. It refers to the accumulation of excess solute particles inside and around the

membrane surface during an osmotic cycle (Sablani et al. 2001). The build-up particles

gradually reduce the effective osmotic pressure difference thereby limiting the water flux

through a membrane. Therefore, any desalination approach should take into account this

phenomenon.

Generally, two categories of polarisation are recognised. For dense, symmetrical membranes,

external concentration polarization (ECP) is common. External concentration polarization

takes place at the surfaces of the active membrane layer. On the feed side, as water permeates

through the membrane, solutes concentrate at the surface, while on the draw side, they

become diluted at the surface as water comes from the feed side. These types of ECP are

referred to as concentrative and dilutive respectively (McCutcheon and Elimelech 2006).

When the membranes are asymmetrical internal concentration polarisation takes place: when

the dense rejection layer faces the feed solution, the draw solution inside the support is

diluted by the water permeating through the porous support layer and when the layer faces


the draw solution, water permeating through the membrane concentrate solutes inside the

support (Gray et al. 2006).

1.4. Research Contribution

Selection of the right combination of membrane and osmotic agent remains one of the key

challenges in FO (McCutcheon and Elimelech 2008; Zhao et al. 2012). Studies have analysed

various systems and combinations with different types of membranes and both organic and

non-organic osmotic agents for FO separately and as a part of the MOD process. However,

there is still no consensus with regards to what the best system or the best process conditions

are.

Numerous studies have focused on the FO process as a separate system. For example, Al-

Hemiri et al. (2009) investigated performance of the systems that included TFC-HR and

TFC-UFP membranes with magnesium sulphate hydrate, potassium chloride, calcium

chloride and ammonium bicarbonate as osmotic agents. The study found that the systems

using calcium chloride as the osmotic agent showed higher water flux rate, and the rates

increased with the draw solution concentration and increasing feed solution flow rate. The

study found that use of spiral wound RO membrane was possible for FO process. Zhao and

Zou (2011) investigated the performance of a FO system that included FS-TCA membrane

combined with Na2SO4 osmotic agent under different temperatures. In the experiment,

temperature rise from 25 to 45 oC resulted in increases in permeability and fluxes, although

higher fouling tendency of the membrane at the higher temperatures was also noted. Achilli

et al. (2010) conducted a comprehensive study on the choice of osmotic agents for FO

applications. They rendered seven inorganic substances as suitable osmotic agents; however,

in view of different physico-chemical characteristics of these substances the researchers

insisted on linking the choice of osmotic agents to the specific FO applications and

membranes considered. The various factors studied by Achilli et al. (2010) in their study

were different osmotic agents such as NaCl, KCl and MgSO4, a flat-sheet cellulose triacetate

(CTA) membrane and a water flux ranging from 10.9 L(m2.hr) for KCl to 5.51 L(m

2.hr)for

MgSO4 at 2.8 MPa and at a temperature of 25 oC for feed and draw solutions.

More recent research by Al-Zuhairi et al. (2015) investigated the efficiency of the FO process

within a MOD system A commercially available flat sheet TFC-ULP membrane was

combined with magnesium sulphate (MgSO4) as an osmotic agent. The experiments were

conducted for varying osmotic agent concentrations, flow rates and temperatures of feed and


draw solutions. With the increase in draw solution concentration from 0.35 to 2.0 M water

under the constant temperature (25 oC) and flow rates (3 L/min), flux grew to 7 L(m

2.hr)for

the brackish water with 10,000 ppm NaCl and to 2 L(m2.hr)for the brackish water with

40,000 ppm NaCl. Water flux of nearly 1.9 L(m2.hr) was achieved by increasing OA

concentration, feed water and draw solution rates (3 L/min). Water flux of 2.5 L(m2.hr) was

achieved at the draw solution temperature of 35 oC and feed solution temperature of 25

oC.

The researchers concluded that while the examined system provided acceptable water flux

rates, membrane improvements and better osmotic agent could further increase the process

efficiency.

The major contribution of this research is in examining the new combinations of membranes

and osmotic agents for FO as a part of the MOD process. In addition to water flux, other

important parameters are taken into account when analysing efficacy of the system.

Specifically, the current research applied a MOD approach using two organic osmotic agents

(sucrose and glucose) against deionised water, simulated brackish water and seawater

(synthesized by NaCl salt) as feed waters. The efficiency of specific organic osmotic agents

as stated above in FO process was investigated using more efficient membrane types

[nanofiltration (NF), RO flat sheet and hollow fine fibre (HFF)]. The experiment also used

binary and ternary systems using hollow fine fibre membrane (HFFM) in order to determine

the FO performance efficiency in terms of different parameters, namely water flux, specific

energy consumption, water recovery, water permeability and solute flux. The experimental

conditions included concentration of osmotic agents, temperature and the flow rate of feed

water and osmotic agents.

1.5. Research Aims and Objectives

The purpose of this study is to investigate the performance efficiency of various membrane

types of membranes in FO process using specific organic osmotic agents such as sucrose and

glucose versus deionised and salty feed water.

The main goals this research include:

To study the appropriateness of using a number of selected an organic osmotic agents

such as sucrose and glucose in the Forward Osmosis process using different types of

membranes.

To determine the effects of osmotic agent concentration, flow rate and temperature on


parameters on the performance efficiency of FO process in term of water flux,

recovery percentage, water permeability, specific energy consumption, and solute

flux.

To investigate suitable membrane types that could be used in Forward Osmosis

process.

The experimental procedure involved usage of a rig set (Figure 1.4) to determine and

measure the effect and optimal parameters conditions such as concentration, flow rate,

temperature and fluxes for feed and draw solutions by using different types of the

membranes.

Figure ‎1.4.MOD Laboratory Test Trial Plant Rig in Surrey University

To accomplish the goals of this study, two sets of experimental work have been carried out.

The first set involved using NF and RO membrane flat sheet types small scale forward

osmosis unit. The performance efficiency of the FO process using specific organic

compounds such as sucrose and glucose was examined as osmotic agent whereas deionised

and feed salty water (brackish water) using as feed water. The efficiency was measured in

term of water flux, water recovery, specific energy consumption, and solute flux. Parameters

affecting the forward osmosis process, such as osmotic agent concentration, flow rate and

temperature were also examined. In the second experiment set, hollow fine- fibre membrane

manufactured by Toyobo Company was used in a forward osmosis pilot plant unit.

Performance efficiency was again measured using the above mentioned factors. Different

concentration of sucrose and glucose were used as osmotic agents verses deionised and salty

feed water (brackish water) whereas other parameters were kept constant for using in the


second run. Additionally, in this study the forward osmosis efficiency as well as using

mixtures of sucrose-sodium chloride and glucose- sodium chloride were tested versus

deionised and salty feed water.

1.6. Thesis Organisation

This thesis consists of several distinctive Chapters. Chapter two provides a comprehensive

review of the relevant literature on Manipulated osmosis Desalination process with a specific

focus on forward osmosis process. The relevant issues in FO for desalination such as

concentration polarisation, choice of osmotic agents and selection of membranes are

discussed as well. Chapter three provides a detailed description of the experimental work and

procedures specifically; methods and materials are described for both the flat-sheet and

hallow-fine fibre studies. Chapter four provides the results of the studying for different types

of membranes and osmotic agents, and presents a detailed analysis of the data. Chapter five

presents conclusion. Recommendations for future work are presented in Chapter six.


2. CHAPTER TWO: LITERATURE REVIEW

This chapter presents a review of the relevant literature for this research. The first part of the

chapter explains the importance of desalination process as a means to produce fresh water

and the need for developing novel approaches to water desalination. The second part of the

chapter provides a thorough review of forward osmosis (FO) process, which is the key

process investigated in this thesis. Specifically, the section describes FO process, its main

principles and applications, advantages and challenges related to applications of FO for

commercial water production. The final part of the chapter focuses on manipulated osmosis

desalination (MOD) which combines FO with other desalination methods for a more efficient

process.

2.1. Water Desalination

2.1.1. Importance of Water Desalination

Fresh water is a vital yet very scarce resource existing in natural environment. Among Earth

water resources, only 0.5% is in form of potable, fresh water (Kim et al. 2010). The growth of

global population, development of industrial production and agriculture and the enhanced

living standards are putting increasing pressures on supplies of clean water across the globe.

An additional factor that is expected to increase water shortages in the future is climate

change, which leads to melting of the glaciers, pollution of river deltas by the rising sea levels

and disturbances in seasonal and annual rain cycle processes in many regions. Recent

projections by the United Nations and the Organisation for Economic Cooperation and

Development indicate that by 2025 the total number of people living in regions with absolute

water scarcity will reach 1.8 billion, and at least two thirds will be living in water-stressed

regions (UNDP 2006). Therefore, to alleviate the problems related to fresh water shortages, it

is imperative to look for novel ways to produce it.

Water desalination is one of the perspective ways to address fresh water shortages because

sea water accounts for 97% of all water resources on Earth (Kim et al. 2010). Therefore,

desalination technology may provide a seemingly endless supply of water without altering the

natural fresh water ecosystems. Desalination is also completely independent of climatic

conditions of a region, which makes it applicable even in the most arid regions that have


access to sea. For some regions, where there is shortage of rainfall and no easy access to fresh

water supply, desalination may be the only viable option to sustain the population needs.

The term desalination is applied to processes that remove salts from water to make it suitable

for drinking and other uses (ElDessouky and Ettouney 2002). On average, seawater salinity is

approximately 35,000 mg/L of dissolved solids, while drinking standards require about 500

mg/L (Cooley et al. 2006). Water desalination is an essential part of the natural hydrologic

cycle whereby water is evaporated by sun from the surface of oceans and precipitates in form

of rainfall (Gleick 2000). As a technical concept, desalination refers to a range of

technologies used by human to remove salts from water.

2.1.2. Brief History of Water Desalination

Early references to desalination can be traced to Aristotle who mentioned that sea water

vapour is sweet and does not contain salts when condenses (Gleick 2000). However, at

industrial scale, seawater desalination came to use starting from the second half of the

nineteenth century. In 1869, a British patent was granted for a first industrial device to

convert seawater into drinkable water, and based on this patent the first desalination plant

was constructed in Arden the same year to supply ships travelling through the Red Sea with

fresh water (Norton et al. 2003). The beginning of the twentieth century marked more active

development of desalination plants. In 1903, Australia launched its first small scale plant in

Perth (Heath 2013). In 1912, Egypt installed six desalination plants with a total capacity of 75

m3/day (ElDessouky 2007). Later, in 1928, the island of Curaҫao committed itself to

completely covering own needs for fresh water by installation of desalination plants (Gleick

2000).

Early desalination plant designs were mostly based on evaporation technique and had small

capacities. The capacities increased with the emergence of the oil industry in the 1930s, and

the first large scale plant was introduced in Saudi Arabia in 1938 (Cooley et al. 2006). In the

following three decades, global desalination plant capacities were growing at average annual

rate of 17% (ElDessouky 2007). However, existing at the time plant designs suffered from

high fouling tendencies, limited operating hours and required lengthy cleaning process

(ElDessouky 2007). A breakthrough was achieved in 1957 when a novel multistage flash

distillation plant was launched in Kuwait and provided much higher efficiencies than the

previous plant designs (McArthur 2001). Other milestones were achieved at that time


including on-line ball cleaning systems, co-generation plants and efficient anti-scaling

chemical treatment (ElDessouky 2007). All these achievements leading to much higher

efficiencies in fresh water production prompted an exponential increase in desalination plant

capacities since the 1960s (Figure 2.1).

Figure ‎2.1. Cumulative Installed Capacities of Desalination Plants: 1945-2005 (Gleick et al.

2006)

Alongside evaporation based desalination processes, researchers in the 1960s examined the

possibility of applying membrane-driven approaches. In 1959, a Sidney Loeb and Srinivasa

Sourirajan of UCLA managed to create a functional synthetic membrane from cellulose

acetate polymer that allowed successful passage of water under pressure while retaining salts

and TDS (Denn 2012). Because the process worked in reverse direction of the naturally

occurring membrane processes in nature, it became known as Reverse Osmosis (RO). This

discovery launched a new era in seawater desalination by introducing a new approach that

would later become dominant across the globe. One year after Loeb and Sourirajan’s work,

McCutchan and Loeb started a project that culminated in the construction and launch of the

first small scale pilot RO plant in Coalinga in 1965 (Stevens and Loeb 1967). While the plant

used brackish water as a feed, the possibility of applying membrane processes for

desalination practically demonstrated by Coalinga plant raised worldwide interest in this form

of desalination for seawater.

The first large scale RO plant was constructed in Jeddah, Saudi Arabia with a total production

capacity of 12,000 m3/day (Al-Gholaikah et al. 1978). Practical applications of RO for water

desalination were initially feasible only in the regions with low costs of fuel because earlier


plant designs required enormous amounts of energy. The aforementioned Jeddah plant, for

example, used up to 8 kWhr of electricity to produce a single cubic meter of freshwater

(Prud’homme 2012). However, improvements in membrane technologies, application of

energy recovery devices, reduction of material costs and achievement of economies of scale

have led to substantial decline in RO desalination cost (Figure 2.2). Recent large scale

desalination plants are capable of reaching costs per cubic metre well below a dollar (Bernat

et al. 2010).

Figure ‎2.2. RO-Based Desalination: Cost Evolution (Reddy and Ghaffour 2007)

2.1.3. Current Status of Desalination Technologies

Recent estimates provide that the current global desalination capacities exceed 66 million

m3/day with the projections to reach 100 million m

3/day by 2015 (Ghaffour et al. 2013).

Accordingly, the total market for desalinated water is expected to surpass $31 billion by that

time (Pankratz 2008). Desalination technologies for obtaining freshwater have been actively

applied in the regions with substantial water scarcity, such as the Middle East and North

Africa (MENA). Some countries, like Kuwait and Qatar, fully rely on desalination plants for

their industrial and domestic water supplies (Ghaffour 2009). In some regions, desalination is

also becoming a viable competitive alternative to the traditional sources (dams, canals,

reservoirs) of potable water (Quteishat 2009). Finally, water desalination has been actively

explored by some countries where there is no lack in freshwater supplies, such as Spain,

Australia and the United States. The US currently ranks only below Saudi Arabia in their

share of global desalination production capacity (Greenlee et al. 2009).

Reduced capital costs of freshwater production, lower energy requirements and technology

modularity have made RO the most popular desalination approach today. It currently

accounts for roughly 60% of all desalination capacities worldwide (Figure 2.3). MSF and


MED, however, are likely to remain in use, especially, in the regions with low energy costs,

such as MENA where MSF has been traditionally accepted and relied on. In the Middle East

countries, MSF, unlike the rest of the world, is the dominant water desalination approach

accounting for nearly 87% of all installed capacities (Figure 2.4).

Figure ‎2.3. Desalination Technologies’ Share in Global Installed Capacities (Global Water

Intelligence, 2012)

Figure ‎2.4. Desalination Technologies’ Share in the Middle East (Greenlee et al. 2009)

60%

26%

8% 6%

Reverse Osmosis Multistage Flash Multi-Effect Distillation Other Technologies

86.70%

10.70%

1.80% 0.80%

Multistage Flash

Reverse Osmosis

Electrodialysis

Other Technologies


Currently, there are over 15,000 desalination plants around the world (Greenlee et al. 2009),

the fact that confirms the importance assigned to desalination technologies for freshwater

production. Maturity of desalination technologies is reflected in the large scale capacities of

the modern plants. The largest desalination plant Ras Al Khar in Saudi Arabia, which uses a

combined MSF-RO desalination process, has the capacity of over 1 million m3/day

(Almashabi 2014). Recently, Israel launched the largest RO plant Sorek, which has the

capacity of 624,000 m3/day (Bosecker 2013). However despite the growing popularity of

desalination process and increasing plant capacities, desalination still accounts for less than

1% of the total freshwater production (Cooley et al. 2006). This means that there is still a

large potential for expansion of desalination approaches.

The continuing search for novel ways to desalinate water is prompted by the need to further

reduce costs and improve effectiveness of water recovery (that is, efficiency of salt removal).

While overall cost and effectiveness of plant operations vary from site to site, researchers and

practitioners define a number of factors that are relevant for any process. Energy requirement

is the main parameter considered in desalination process (Ghaffour et al. 2013). It is

estimated that energy accounts for about 50% of desalination plant costs (Herndon 2013).

While large plant sizes allow reduction of energy cost per unit of output due to economies of

scale, construction of such plants is not always feasible due to environmental concerns.

Another possible way to reduce energy is to use an optimised plant design and novel methods

of water extraction.

Optimisation of water desalination process is often achieved by effective combination of two

desalination approaches. A common way to reduce cost of water production has been through

a plant configuration that combines MSF or MED with RO or nano-filtration (Ghaffour et al.

2013; Malivaet al. 2006). The world’s largest desalination plant Ras Al Khar, for example,

uses a combination of MSF and RO processes (Almashabi, 2014). Advantages of such

systems over single approach desalination process are in using common intake and outfall

systems, exchange of thermal heat between the processes and integrated pre- and post-

treatment all of which reduce energy costs and increase water recovery (Ghaffour et al.

2013). However, these parameters could be improved even further by applying novel

approaches to water desalination. One of the most perspective directions in this regard is the

use of Forward Osmosis (FO) a natural membrane process requiring 25-45% of the thermal

energy consumed by the most efficient thermal desalination process (Li et al. 2011). This is

because in the core of FO process is movement feed solution by means of osmotic pressure


from low pressure side consisting of feed water to high pressure side consisting of the draw

solution. Therefore, FO, in perspective, allows to minimise energy costs and increase water

recovery (Martinetti et al. 2009; McGinnis and Elimelech 2007; Zhao and Zou 2011).

While the potential of FO for improving water desalination process has received increased

interest from the research community, practical application of this process for water

desalination has been limited. Research continues to identify the best types of membranes,

osmotic solutions and process configuration to bring FO to par with the commercially

attractive technologies on the market today. This paper presents an experiment for

optimisation of FO process using different membrane types and osmotic agents. The next

section provides a detailed description of FO process, which is followed by presentation of a

novel manipulated osmosis desalination technique that uses FO as the primary step.

2.2. Types of Desalination

Modern theory and practice of desalination recognizes more than a dozen approaches. In

practice, all approaches can be placed within two major categories: thermal and membrane

(Figure2.5). A number of alternative approaches are being tested, including

electrodeionisation, vacuum distillation, freeze separation, rapid spray evaporation, capacitive

deionisation and freezing with hydrates. At present, however, most of these approaches are in

either theoretical or experiment stages and do not have practical applications. Because this

review aims to discuss the alternative desalination approaches that are in use, it is limited to

thermal and membrane categories only.

2.2.1. Thermal Approaches

Thermal desalination is the oldest and the most widespread approach to water desalination

(Mezher et al. 2011). In simplest terms, the approach assumes boiling the feed water to purify

and collect the vapour. Thermal desalination is a rather energy demanding process, which has

been widely applied in the Middle East, where the cost of petrochemicals is rather low.

Common thermal desalination approaches are multistage flash (MSF), single effect

evaporation (SEE), multi-effect evaporation (MEE), freezing, humidification and solar stills.


2.2.1.1. Multistage Flash (MSF)

Multistage flash represents a process made of several flashing spaces called stages, each stage

containing an exchanger to heat the feed water and a collector for condensate steam. The

entire sequence has a cool and a hot end, and the stage temperatures between them increase.

Each stage has a different pressure level, which matches the boiling points of water solution

at that stage’s temperature. Feed water travels through each stage, where it is heated by

warmer water vapour. By reaching the final stage-the brine heater-the water solution has

achieved a maximum temperature and flows back through the stages, where certain amount

of it evaporates through the boiling (flashing) process. The evaporating steam is hotter than

the feed water; therefore, it cools down and condenses on the heat exchangers and heats the

feed water (Younos and Tulou 2005).

The most popular MSF systems on the market today are brine circulation MSFs where the

flashing stages are divided among the heat recovery and heat rejection systems (El-

Dessoukey and Ettouney, 2002). In once through MSF, the feed only used once without

recirculation is it flows through the consecutive evaporating chambers at decreasing

temperatures. In brine mixing MSF processes, the brine gets mixed with the feed in order to

achieve lower feed water flow and increase thermal efficiency of the process. Finally, vapour

compression MSF systems include additional compressors, blowers or jet ejectors to increase

the vapour pressure and thereby make it an additional heating medium for feed solution (El-

Dessoukey and Ettouney, 2002).

2.2.1.2. Single Effect Evaporation (SEE)

By definition, SEE is a process that takes place in a single heating unit. There are several

types of SEE on the market. In thermal vapour compression (TVC), the steam jet ejector

compresses the vapour which enters the first effect tubes. The vapour gets condensed by a

condenser to the final product. An alternative way to compress the vapour is by mechanical

compressors instead of jet ejectors. It is considered that thermal vapour compression has a

higher performance ration in comparison to the multi-effect evaporation (El-Dessouky et al.

2000). Mechanical vapour compression (MVC) units have a similar principle of operation

with the only difference of using mechanical compressors in place of steam jet ejectors

(Younos and Tulou, 2005). Such units are portable and can be easily transported to different

areas. Adsorption vapour compression (ADVC) creates pressure differences between the two

reservoirs, where feed water is transferred. Pressure differences are caused by an exothermic


reaction, where the feed water is preheated by being mixed with a special solution like LiBr.

This approach typically generates pot water (Younos and Tulou, 2005). In absorption vapour

process (ABVC), the feed water is treated with solutions like LiBr-H2O for water absorption

on the shell side of the absorber ((El-Dessouky and Eltounney 2002). Chemical vapour

compression (CVC) approach uses chemicals for the treatment of feed water and vapour (El-

Dessouky and Eltounney 2002).


Desalination processes

ThermalMembrane

RO ED

MSF

MEE

SEE

Freezing

HDH

Solar

Stills

Parallel

Feed

Forward

Feed

Vertical

Stack

TVC

MVC

ADVC

ABVC

CVC

Vapour

Compression

Brine Mixing

Once Through

Brine

Circulation

TVC

MVC

ADVC

ABVC

CVC

Sea Water

RO: Reverse Osmosis

ED: Electro-Dialysis

MEE: Multiple Effect Evaporation

SEE: Single Effect Evaporation

MSF: Multistage Flash

TVC: Thermal Vapour Compression

MVC: Mechanical Vapour

Compression

ADVC: Adsorption Vapour

Compression

ABVC: Absorption Vapour

Compression

CVC: Chemical Vapour

Compression

HDH: Humidification -

dehumidification

NF: Nano -Filtration

MF: Micro-filtration

UF: Ultra filtration

FO: forward Osmosis

Low Salinity Surface

Brackish Well Water

NFMFUFFO

Figure ‎2.5. Thermal and Membrane Desalination Approaches and Methods (adapted from El-Dessoukey and Eltounney, 2000)


2.3.1.3. Multi-Effect Desalination

Multi-effect desalination (MED) is a thermal distillation process which involves several

evaporation stages where feed water is preheated and placed on special surfaces. Figure 2.6

provides a visual presentation of a typical MED process. The initial saline feed is transferred

to the first chamber where steam from a boiler heats the evaporator tubes to drive evaporation

of the preheated feed water. The generated steam is condensed within the evaporator tubes of

the next stage and the cycle of vapour production repeats. The evaporating surfaces in each

stage are heated by the steam generated in the previous stage, and the pressure in each

following chamber is reduced. The steam generated during the final stage of MED process is

cooled by the incoming feed water. MED offers better thermal performance than MSF, and

this process can be modified for both high and low temperature operations, which makes it an

attractive option (El-Dessouky and Eltouney, 2002).

Figure ‎2.6. Multiple Effect Desalination Process (Al-Harbi et al. 2011)

2.2.1.4. Freezing, Humidification and Solar Stills

Freezing, humidification and solar stills are less used but still existent technologies for

thermal water desalination. Freezing process of desalination is based on freezing of pre-

cooled feed water with the subsequent separation of water in form of ice. Vacuum or

chemicals such as the glycol solutions are used to freeze the feed to a slurry form with its

subsequent pumping (Efrat, 2011). Humidification process of desalination is similar to the

basic natural water cycle only with a shorter period of time. Generally, it uses ambient


pressure to evaporate and condensate the feed water thereby speeding up these cycles (El-

Dessoukey and Ettouney, 2002). Solar stills represent a basic configuration of such process.

These are boxes covered with glass or light plastic on the internal side of which water vapour

condenses (Anjeneyulu et al., 2012).

2.2.2. Membrane Approaches

Membrane technologies can be considered a successor to traditionally used thermal

desalination approaches: the vast majority of contemporary desalination plants, as well as

those being planned utilize membranes (Elimelech and Phillip 2011; Fritzmann et al. 2007).

Membranes represent a thin layer of porous material, which permeates water molecules

through, but holds large molecules, viruses, bacteria and different solutes. They may be made

of different polymeric and nonpolymeric materials, although the most commonly used

membranes for desalination are synthetic (Younos and Tulou 2005).

Based on the technological process used for desalination, all membrane based approaches can

be classified as either electrical or pressure driven. Electrical driven approaches include

electrodialysis and electrodialysis reversal, while pressure driven approaches include reverse

and Forward Osmosis, nano-, ultra-, and microfiltration. All these approaches are briefly

described below.

2.2.2.1. Electrical Driven Approaches

Electrodialysis is a desalination approach relies on membranes and electrodes adjacent to

them. The separation of water from the particles dissolved in it occurs within membrane pair,

which includes an anion membrane, a cation membrane, and two spacers, which evenly

distribute the feed water flow across the membrane surface (Brunner 1990).

Dissolved anions such as NO3- and Cl

- are attracted to the anode, while cations such as Na

+

or K+ are attracted to the cathode. Anion exchange or cation exchange membranes between

the anode and cathode control the movement of anions and cations. The ED stack is separated

into multiple cells by anion and cation exchange membranes labelled AEM and CEM

respectively. The alternating sequence of membrane placement allows to reduce ionic species

in the diluate compartments of the stack and increase them in the concentrate compartments

(Fritzmann et al., 2007). The electric circuit is closed in the first and last cells of the ED

stack. ED process is quite effective with removal of salts. As a rule, the process requires 70-


90 psi pressure and is able to remove 75% to 98% of solids from feed water (Younos and

Tulou 2005). The chemical and petrochemical industries often apply ED for treatment of

dilute aqueous solutions (Fritzmann et al., 2007). At the same time, ED has relatively high

energy requirements for the processes involving highly concentrated feeds; therefore, today it

is rarely used for seawater desalination.

2.2.2.2. Pressure Driven Approaches

The pressure driven approaches of desalination include microfiltration (MF), ultrafiltration

(UF), nanofiltration (NF) and reverse osmosis (RO). These approaches are based on filtration

technique using thin film material with thickness ranging from 1 to 250 micron. Membrane

materials, pore size and operating pressure are the major differences between the approaches.

MF membranes are made of ceramic, polysulfonic (PS), polyvinylideneflouride (PVDF) or

cellulose acetate with tubular or hollow fibre membrane module configurations and pore sizes

of 0.02 - 4 micron. MF processes operate at the lowest pressure (2 bar), but they are only

good for removal of larger suspended particles, clay, bacteria and biomass. UF membranes

are made of the same materials as MF, although they have smaller pore size of 0.02 – 0.2

micron. This allows UF membranes additionally reject macromolecules, proteins, viruses and

polysaccharides. However, these membranes also require higher operating pressures, usually

1-10 bar. Membrane configurations include tubular hollow fibre, spiral wound and plate-and-

frame. NF membranes have even smaller pores of less than 2x10-3

microns. This allows

additional rejection of higher molecular weight compounds, mono- and disaccharides, as well

as polyvalent ions. MF, UF and NF are useful for particular types of feed waters and for

different water treatment purposes. MF, for example, is widely applied for fermentation and

biomass recovery, UF is used for fractionation of organic compounds, and NF is widely

applied for colour removal and demineralisation. Although sometimes applied as a standalone

technology for low TDS brackish water feeds, nanofiltration, nevertheless, cannot be

effectively used to treat seawaters (Younos 2005).

Reverse Osmosis approach is based on the phenomenon of osmosis, i.e. the passage of liquid

from the region of high concentration to the region with low concentration through a one-way

permeable membrane. Consequently, Reverse Osmosis is the opposite of that: water travels

from the regions with low concentration to the region of high concentration under artificially

applied pressure. RO allows to effectively remove TDS concentrations up to 45,000 mg/L,

which is sufficient to purify both salt and brackish waters. At the same time, RO requires


energy to operate the pumps pressing feed water through the membranes, although the

amount of energy is generally less than under thermal desalination approaches. The required

pressure levels, however, rise with the constantly increasing feed water concentration (up to

150 bar for seawater), thereby giving only 30%-80% recovery rates (Younos and Tulou

2005). Membranes play a crucial role in Reverse Osmosis desalination process. In general, all

membranes can be classified into cellulose acetate (CA) and non-cellulose acetate (NCA).

CA membranes have smooth surfaces, which are resistant to fouling, because the particles

causing fouling do not deposit in the membrane crevices (Nicolaisen 2002). NCA membranes

are commonly referred to as “thin film composite membranes.” They use organic materials,

such as polysulfone. The advantages of NCA are higher flux rates and relative stability over a

long range of pH levels; however, they are quite sensitive to the effects of chlorine (El-

Dessouky and Ettouney 2002).

Protection of membranes in Reverse Osmosis and other membrane-based desalination

approaches is necessary to enhance retention of salts and reduce cost of energy. Therefore,

feed water pre-treatment is typically required to remove very large particles, bacteria, oil, and

organic matter that may damage membranes. Pre-treatment also involves addition of

chemicals to avoid membrane scaling and the formation of precipitates as well as adjustment

of pH levels depending on the membrane type. Microfiltration and ultrafiltration are typical

pre-treatment approaches used in RO. A summary of pressure driven desalination processes

is provided in Table 2.1.


Table ‎2.1. Description of Pressure Driven Membrane Processes Source : Mortazvi (2008)

2.4. Forward Osmosis

2.4.1. Process Description

Forward Osmosis (FO) is an osmotic process that uses a semi-permeable membrane to

separate water from the solutes dissolved in it. Unlike reverse osmosis (RO), which uses

hydraulic pressure for such separation, FO operates by osmotic pressure whereby water flows

from the region of low concentration to the region with high concentration (Boo et al. 2013;

Membrane

Reverse

osmosis

(RO)

Nanofiltration

(NF) Ultrafiltration (UF)

Microfiltration

(MF)

Asymmetric Asymmetric Asymmetric Asymmetric

symmetric

Thin film

thickness

1 micron

150 micron

1 micron 150

micron

1 micron 150–250 micron 1–150 micron

Rejection High and

low

molecular

weight

compounds,

NaCl,

glucose,

amino acids

High molecular

weight

compounds,

mono-, di- and

oligosaccharides,

polyvalent ions

Macromolecules, proteins,

polysaccharides, viruses

Particles, clay,

bacteria

Membrane

materials

Cellulose

acetate

(CA) thin

film

CA thin film Ceramic, polysulfonic (PS),

poly vinylideneflouride

(PVDF), CA thin film

Ceramic, PS,

PVDF, CA

Pore size <0.002

micron

<0.002 micron 0.02–0.2 micron 0.02–4 microns

Module

configuration

Tubular,

spiral

wound,

plate-and-

frame

Tubular spiral

wound, plate-

and-frame

Tubular hollow fiber spiral

wound, plate-and-frame

Tubular, hollow

fiber

Operating

pressure

15–150 bar 5–35 bar 1-10 bar <2 bar


Coday and Cath 2014; Valladares-Linares et al. 2014). Figure 2.7depicts a basic FO process:

pure water is extracted from source water, which is the area of low osmotic pressure, to draw

solution with an osmotic agent, which is the area of high osmotic pressure. The larger

osmotic pressure exercised by the draw solution is, the stronger is the separation process in

FO (Helix-Nielsen 2010; Zhao et al. 2012). By using primarily osmotic instead of hydraulic

pressure, FO requires much lower energy than RO. In theory, if a sustainable water treatment

solution is available FO can operate at zero hydraulic pressure (Valladares-Linares et al.

2014).

Figure ‎2.7. Forward Osmosis Process (Thompson and Nicoll 2011)

As it is seen from the above figure, in an osmotic process, two elements are present: a semi-

permeable membrane and a draw solution. The differences in water chemical potentials

between the feed solution and the draw solution initiates the flow of water through the

membrane, which, due to a small size of pores, allows only water molecules through, while

holding the molecules of solutes and organic materials.

The primary difference of FO from the widely used RO process is in the direction of flow

between the feed and the draw solutions. FO operates through the osmotic pressure

differential between the solutions: the feed solution gets concentrated, while the draw

solution gets diluted until the equilibrium is achieved. This is in contrast to the RO approach,

where hydraulic pressure needs to be applied to drive the process. To demonstrate how the

process works, consider Figure 2.8for illustration. The picture shows the direction and

magnitude of water flux Jw through an ideal semi-permeable membrane based on differences


between osmotic pressure differential ∆π and hydraulic pressure differential ∆P. When

hydraulic pressure is exerted (that is, ∆P > ∆π), RO takes place and water flows from the area

of lower to higher concentration. When osmotic pressure differential is larger (that is, ∆P <

∆π), FO takes place, and water flows towards the area of higher concentration by means of

osmotic pressure. Pure FO, also referred to as Direct Osmosis (DO), takes place without any

hydraulic pressure (∆P = 0) as water flows through a membrane by means of osmotic

pressure only. When osmotic pressure differential is larger than hydraulic pressure

differential, Pressure Retarded Osmosis (PRO) is said to take place. The slope of the line in

the figure determines the membrane’s permeability coefficient Aw: a 100% salt rejection

when total volumetric flux Jv equals pure water flux Jw which, in turn, is zero at the

equilibrium point (∆π = ∆P).

A simple equation to determine water flux under FO was developed by Lee et al. (1981), who

generalised it for PRO process:

Jw= A(∆π − ∆P) Eqn.2.1

where:

Jw is the water flux;

A is the water permeability constant depending on the membrane used;

∆π is the osmotic pressure differential between the solutes;

Water Flux, Jw

Hydraulic Pressure

Differential, ∆P

DO point: ∆P = 0

Equilibrium (flux reversal)

point: ∆π = ∆P

RO area:

∆P > ∆π

FO area:

∆P < ∆π

Figure 2.8. Water Flux Magnitude and Direction as a Function of Manually Applied

(Hydraulic) and Osmotic Pressures (adapted from Lee et al.1981)


∆P is the hydraulic pressure differential.

Pressure differentials in both cases are derived by differences between feed water (FW) and

draw solution (DS). Therefore, osmotic pressure differential is determined as:

∆π = πFW – πDS Eqn.2.2

where :

πFWis the osmotic pressure of feed water;

πDSis the osmotic pressure of draw solution.

Similarly, hydraulic pressure differential is determined as:

∆P = PFW – PDS Eqn.2.3

where :

PFW is the hydraulic pressure for feed water;

PDS is the hydraulic pressure for draw solution.

Consequently, a general FO equation can be expressed as:

Jw= A[(πFW – πDS) – (PFW – PDS)] Eqn.2.4

The equation above represents a simplistic model of a FO osmosis process. In a typical FO

process, water flux rate is a function of multiple factors including pressure and flow rate

differences between feed and osmotic solutions, the membrane area and type, as well as type

and concentration of the osmotic agent (Petrotos et al. 1998; Salter 2006; Zhao et al. 2012).

For FO, additional issue to consider are external and internal polarisations which have direct

influence of water flux. These issues are discussed later in the chapter.

2.4.2. Advantages of FO

The advantages of FO come from three main characteristics of the process, namely, low

hydraulic pressure, high osmotic pressure, and high rejection rates. Low hydraulic pressure

requirement ensures low energy consumption, which can be theoretically lowered to zero if

appropriate draw solutes and regeneration techniques are developed (McGinnis and

Elimelech 2007, 2008; Zhao et al. 2012). This remains the major attractive point of FO

applications considering finite energy resources and the increasing costs of their extraction.


Moreover, FO allows for energy harvesting by mixing saline and fresh water, since the

natural DO process releases energy (Achilli and Childress 2010; Gerstandt et al. 2008).

Another advantage of FO is low degree of membrane fouling and higher reversibility (Achilli

et al. 2009; Mi and Elimelech 2010). Studies also demonstrated that membrane fouling in FO

can be minimised if the hydrodynamics is optimised (Lee et al. 2010). Low membrane

fouling means that membranes can serve longer and that water flux would decline at a lower

rate. FO can also effectively remove a wider range of contaminants in comparison to other

osmotic processes, which means better quality of water (Cartinella et al. 2006; Cath et al.

2010). High osmotic pressure gradient across the membrane also creates the potential for

higher water recovery rate (Zhao et al. 2012). On the one hand, this means better freshwater

return per unit of feed water; on the other hand, this leads to lower brine volumes discharged

to the environment (McCutcheon et al. 2005). Finally, FO allows to preserve the physical

properties of the feed water (such as nutrition, taste, and aroma) because it does not use

heating or pressurisation of the feed (Jiao, et al. 2004; Petrotos and Lazarides 2001; Yang et

al. 2009). This is an important feature for food and medical applications. Due to a wide range

and variety of the advantages provided by FO, this process found numerous applications,

which are discussed below.

2.4.3. FO Applications

2.4.3.1. Energy Applications

Using osmotic process for power generation was suggested first in the 1950s. Pattle (1954)

described process where fresh and salt water could be mixed in order to produce electricity.

The primary advantages of power generation by osmotic process are renewability and

sustainability, which make it more attractive than the finite and environmentally

compromising fossil fuels (Post et al. 2008). According to Veerman et al. (2009), up to 2.5

MJ of energy can be generated by mixing a cubic metre of seawater and a cubic metre of

river water. In total, it is estimated that the gross potential for osmotic power generation is

equivalent to 2.4-2.6 TW (Zhao et al. 2012).

Pressure retarded osmosis (PRO) has been explored as one of the most perspective directions

in conversing of salinity-gradient energy based on the Gibbs free energy of mixing (Achilli et

al. 2009; Post et al. 2007; Sandler 1999; Thorsen and Holt 2009). As was explained earlier, in

PRO, water is transferred from the feed solution to the draw solution side thereby increasing


the draw solution volume. This, in turn, produces hydraulic pressure that could be used for

pushing an electricity generation turbine. In PRO, the relevant measure is the energy

conversion efficiency of a membrane (W), which is defined as the amount of power

generated per unit membrane of area (Zhao et al. 2012). Energy conversion efficiency is

measured by a product of water flux (Jw) and hydraulic pressure differential ∆P:

W = Jw∆P Eqn.2.5

Using basic water flux equation for FO (Eqn .2.1), this formula becomes:

( ) (

)

By plotting the differentiation of W against ∆P, the highest point is where

.

Therefore, the maximum power density becomes dependent on the water permeability

coefficient (A) and osmotic pressure differential (∆π):

The total power generated by a hydraulic turbine in PRO process can be calculated as (Zhao

et al. 2012):

P = ηρQgP Eqn.2.8

where :

η is the turbine efficiency;

ρ is water density;

Q is the flow;

g is the acceleration;

P is the solution pressure.

The first applications of PRO for energy generation were explored in the 1970s by a number

of studies by Loeb and colleagues (Loeb and Norman 1975; Loeb et al. 1976; Mehta and

Loeb 1979). More interest arose on the wave of search for alternative energy sources in view

of global price hikes of oil products (Achilli and Childress 2010). The first large scale PRO

energy generation plant was launched by Statkraft at Tofte, Norway in 2008 (Statkraft 2014).


In Japan, the Mega-ton Water System project was launched in the beginning of the current

decade. By using seawater reverse osmosis brine instead of seawater, the plant was able to

increase extractable energy from 0.75 kW/m3 to 2.75 kW/m

3 of draw solution (Kurihara and

Hanakawa 2013). In Canada, H2O Innovation company is currently working on launching a

PRO energy generation plant in Quebec (Valladares-Linares et al. 2014). With regards to

power generation, contemporary turbines reach up to 92% efficiency (Yip and Elimelech

2012).

2.4.3.2. Wastewater Treatment

In general, wastewaters have lower osmotic pressure than seawater; however, they cause

much higher fouling (Zhao et al. 2012). Lower fouling propensity under FO is, therefore, one

of the major advantages to use FO for wastewater reclamation and reuse. One of the rapidly

developing areas for FO application in wastewater treatment has been harvesting water from

municipal wastewaters. It should be noted that FO has been mostly used as a pre-treatment

step, not the ultimate treatment solution. However, it plays an important role in reducing

membrane fouling and, therefore, total treatment costs. Xie et al. (2013) described a practical

application of a hybrid forward osmosis-membrane distillation process for sewer mining. The

study demonstrated that FO allowed to reject most of the organic contaminants while

membrane distillation allowed to reject residue contaminants. Zhang et al. (2013) also

described a highly efficient photovoltaic powered hybrid FO-electrodialysis process for

secondary wastewater effluent, which is favourable for using in remote areas. FO was also

tested for dewatering anaerobic digester centrates and activated sludge (Hau et al. 2014;

Holloway et al. 2007).

In the past few years, an increasing number of researchers investigated the concept and

application of Osmotic Membrane Bioreactor (OMBR) (Achilli et al. 2009; Alturki et al.

2012; Cornelissen et al. 2008; Lay et al. 2011; Yap et al. 2012; Zhang et al. 2012). An

example of such system is presented in Figure 2.9. In OMBR, osmotic dilution takes place as

a pre-treatment, after which the diluted draw solution is re-concentrated by an additional step

such as RO to produce fresh water (Zhao et al. 2012). The major advantages of OMBR over

the traditional systems for municipal wastewater treatment are minimisation of fouling

process, a more sustainable flux and more reliable removal of contaminants (Valladares-

Linares et al. 2014). This makes commercial application of OMBR realistic in the near future.

The two major factors in optimising OMBR operations are membrane salt permeability to the


water permeability ratio and hydraulic retention time to sludge retention time ratio (Xiao et

al. 2011).

Figure ‎2.9. An Example of an OMBR (Achilli et al. 2009)

FO also has a number of potential applications in industrial wastewater treatment. Several

studies (Li et al. 2014; Rassoul et al. 2012) investigated FO potential for removal of heavy

metals from industrial wastewaters and showed that FO membranes provided an almost

complete rejection of lead, zinc, copper and cadmium. In the United States, Catalyx Inc. has

successfully piloted a FO plant to recycle wastewaters from textile and carpet mills (Catalyx

2009). The majority of FO applications in wastewater treatment, however, have been

developed for oil and gas industry. An FO-RO combination process in a closed loop was

successfully applied to reclaim drilling wastewater from gas exploration operations (HTI

2011). Abousnina (2012) described an effective FO process for dissolved organic

components removal from oily wastewaters. Nelson and Ghosh (2011) showed that a

properly designed FO process can substantially improve treatment of produced water from oil

natural gas operations. Hickenbottom et al. (2013) experimentally demonstrated that a FO-

based process can increase rejection of both organic and inorganic contaminants, reduce

membrane fouling and successfully recover up to 80% of water from oil and gas drilling

waste. McGinnis et al. (2013) pilot tested application of an NH3/CO2 FO membrane brine

concentrator for treatment of produced waters from natural gas extraction and showed that up

to 42% energy savings could be achieved in comparison with the conventional evaporator

techniques while preserving the same quality of reclaimed water. Duong and Chung (2014)

conducted a bench scale testing of a customised FO-based membrane for fresh water

recovery from a heavily concentrated synthetic oily solution to achieve a 99.88% rejection

rate.


Finally, FO was tested for treatment of landfill leachates: highly variable liquids that usually

consist of several types of pollutants, including heavy metals, organic compounds, nitrogen

and total dissolved solids. The traditional wastewater treatment approaches usually do not

deal with the total dissolved solids; in fact, in some cases, after treatment, their concentration

in the solution even increases (Water Development Report 2009). An evaluation of FO’s

potential use was investigated in the UN Water Development Report, and it was found that

FO could be an effective solution for landfill leachate treatment because of its ability to reject

total dissolved solids. In the end of 1990s, an American company Osmotek conducted a three

month pilot testing of a FO-based membrane for landfill leachate in Corvallis, Oregon

(Gleick 1996). The tested FO membrane with the NaCl draw solution were able to achieve

94-96% water recoveries, high contaminant rejection and nearly full flux restoration after

cleaning. This prompted the launch of a full-scale system, which is presented in Figure 2.10.

The full scale system managed to achieve a nearly 92% water recovery with 99% rejection

rate for most of the contaminants (Water Development Report 2009).

Figure ‎2.10. A Full Scale Working FO-Based Leachate Treatment System (Water

Development Report 2009)


2.4.3.3. Pharmaceutical Applications

Two major applications of FO in the pharmaceutical industry have been drug delivery and

product enrichment (Lin and Ho 2003; Santus and Baker 1995; Thombre et al. 1999). Drug

delivery systems using FO use its natural osmotic process for medicine administration,

primarily orally. There are many types of such systems including capsules and tablets that are

coated with semipermeable membranes, polymer drug matrix systems and osmotic pumps –

self-formulating in-line systems such as basic osmotic pump, Rose-Nelson pump and

Higuchi-Leeper pump (Lin and Ho 2003; Nayak and Rastogi 2010; Santus and Baker 1995;

Thombre et al. 1999; Wang et al. 2011). Two studies conducted by Talaat (2009, 2010)

showed great potential of FO for use in ambulatory dialysis systems.

Product enrichment with FO covers such pharmaceutical products as lysozymes and proteins

(Zhao et al. 2012). These products have large molecular structures and they are sensitive to

heat. FO does not require external pressure that creates heat and they can reject large

molecule sizes. Moreover, the final products in the pharmaceutical applications of FO are

concentrates, which eliminates the need for an additional step of separating water from the

diluted draw solution (Zhao et al. 2012). These factors make FO an attractive process in

pharmaceutical applications. Yang et al. (2009) managed to create high quality enriched

lysozyme solutions without the change or denaturation with FO process. Wang et al. (2011)

demonstrated that similar process was possible for protein concentration. Nayak and Rastogi

(2010) demonstrated that FO had higher stability and lower browning index in the process of

anthocyanin concentrating.

2.4.3.4. Food Processing

In general, water removal in the food industry is used to improve the food’s shelf life,

increase its stability and reduce its transportation costs (Petrotos and Lazarides 2001).

Because FO operates at low hydraulic pressure and lower temperatures, it has an advantage

of retaining the sensory and nutritional values of the processed foods (Coday et al. 2014).

Popper et al. (1966) were the first group of researchers to investigate FO applications for food

processing, specifically, in the production of concentrated beverages. Two studies by Petrotos

and colleagues explored FO application for tomato juice concentration (Petrotos et al. 1998;

Petrotos et al. 1999). The studies noted the potential for FO application after pretreatment of

juice by ultrafiltration. Recently, more studies were conducted for different types of fruit

juice: (Babu, et al. 2006; Garcia-Costello and McCutcheon 2011; Garcia-Costello, et al.


2009; Jiao et al. 2004; Nayak et al. 2011). Researchers also investigated FO applications for

processing of fruits and vegetables including carrots, potatoes, peppers, papayas, pears,

pineapples, strawberries and mushrooms (Changrue et al. 2008; Eren and Kaymak-Ertekin

2007; Garcia et al. 2010; El-Aouar et al. 2006; Khoyi and Hesari 2007; Lombard et al. 2008;

Ozdemir et al. 2008; Park et al. 2002; Petrotos et al. 2010; Torringa et al. 2001; Uddin et al.

2004). As with juice concentration applications, these studies showed solid potential for FO

use in food processing, specifically, for dehydration of the products.

2.4.3.5. Other Applications of FO

Besides energy generation, wastewater treatment, pharmaceutical applications and food

processing, FO has been explored in other fields. Phuntsho et al. (2011) described FO

application for direct fertigation by the diluted draw solution. FO-based products have been

proposed and even commercially implemented for emergencies that bring water scarcity. For

example, FO hydration bags have been used by the military and by emergency relief

personnel (Cath et al. 2006; HTI 2011). Cath et al. (2005) proposed a FO-RO compact system

for wastewater treatment in a closed loop, which has been tested by NASA with a potential

for application in long term space travel.

FO has also been actively proposed and tested for green technologies. Several studies

identified FO’s important role in generating biomass energy and biofuels in processes such as

separation of algae biomass (Hoover et al. 2011; Zou et al. 2011). Zhang et al. (2011) also

introduced a novel process by for wastewater treatment and bioelectricity production.

Similarly, FO is being actively investigated for environmental applications. Hoover et al.

(2011) proposed that dilution of the desalination brine by FO prior to discharge into the

oceans would greatly reduce damage to the marine ecosystems. FO can also be applied as a

membrane cleaning process as an alternative to the traditional chemical treatment (Qin et al.

2010; Ramon et al. 2010).

2.4.4. FO Applications for Water Desalination

The initial interest in FO for desalination arose along with the development of RO in the

1970s (Kessler and Moody 1976; Kravath and Davis 1975; Moody and Kessler 1976). Two

major studies can be noted at that time. Kravath and Davis (1975) were the first to suggest the

use of semi-permeable membranes together with concentrated solutions. Kessler and Moody


(1976) investigated practical feasibility of batch FO desalination process by comparing

experimental results with the theoretical models. Similar to Kravath and Davis, they

suggested that FO could be an appropriate process for water supply in emergency lifeboats.

For quite a while, however, FO for desalination research has not been realised in practice, and

many related patents remained on the paper only (i.e. Frank 1972; McGinnis 2002; Stache

1989).

With the progress of technologies and appearance of commercially available membranes for

the process, FO for water desalination has gained closer attention from both researchers and

practitioners. As was discussed earlier, the existing water desalination approaches have

serious environmental problems and relatively high cost of water production. FO, in theory,

can help overcome these difficulties due to its low sensitivity to salt concentrations, higher

water recovery rates, and use of the natural osmotic process that requires low to none manual

pressure to proceed. Since the first half of the current decade, the researchers started

exploring FO process more intensively. FO has been explored both as a standalone

desalination process and in combination with other methods (Bamaga et al. 2011; Cath et al.

2006; Valladares-Linares et al. 2014; Zhao et al. 2012).

Modern approaches to FO water desalination can be placed into two broad categories: direct

and indirect desalination approaches (Valladares-Linares et al. 2014). The difference between

these groups is in what is taken as a feed solution (FS) and a draw solution (DS). The

processes and the differences between them are described below.

2.4.4.1. Direct FO Desalination

In direct FO desalination freshwater is directly extracted from seawater using an osmotic

membrane process. A typical layout of this approach is presented in Figure 2.11. As is seen,

seawater is the FS which is separated from a DS by a semipermeable FO membrane. By

means of osmotic pressure differences, water is drawn towards the DS, while salts and other

solvents are rejected by the membrane. The diluted DS is then treated to extract pure water

and returned back to the desalination cycle, while the FS brine is discharged.


In direct FO the FS (seawater) is common in all processes, there is much variety in DS used.

One method uses thermolytic DS that can be later decomposed by heating into volatile gases

such as CO2 or SO2 (Zhao et al. 2012). After freshwater is recovered, these gases are recycled

by thermal decomposition. An example of such approach for water desalination is a patent by

McGinnis (2002) where KNO3 and SO2 combination is used as a DS. Saturated KNO3 has a

tendency to precipitate out of the cooled diluted DS. The precipitated diluted KNO3 is then

transferred to an additional FO unit where SO2 acts as a DS. Following the osmotic process,

the dissolved SO2 is removed by heating (McGinnis et al. 2002).

More commonly described and tested is the method that uses a mixture of NH3 and CO2 as a

DS. The approach was proposed and investigated in a series of papers by McCutcheon and

colleagues (McCutcheon et al. 2005; McCutcheon et al. 2006). It was shown that the

appropriate mixture of ammonium and carbon dioxide can ensure high water flux (up to 36

Lm-2

h-1

) and feed water recoveries. McGinnis and Elimelech (2007) also demonstrated that

this process required substantially lower amount of equivalent work in comparison to the

traditional desalination approaches such as RO. They concluded that energy savings could be

between 72% to 85% (McGinnis and Elimelech 2007). In a recent study, Chanukya et al.

Figure 2.11. Direct Osmosis Process (adapted from Li et al. 2013)


(2013) analysed the effect of concentration polarisation for this approach and concluded that

FO could a feasible solution for water desalination.

Besides the thermolytic DS used in FO desalination process, studies explored water-soluble

salts and solutions to enhance water flux and recovery with different desalination

technologies. Li et al. (2011) proposed to use special hydrogels with carbon filler particles in

DS and managed to achieve 1.2L/(m2.hr)flux. However, Cai et al. (2013) showed a much

lower maximum flux (about 0.2 L/(m2.hr)using hydrogels in a similar process. Better results

were demonstrated by Ling and Chung (2011) who used hydrophilic nanoparticles and

ultrafiltration technique to regenerate DS: they managed to achieve 6 L/(m2.hr)flux and

92.7% recovery using synthetic water as a FS (Ling and Chung, 2011). Zhao et al. (2012)

showed even better flux of up to 10 L/(m2.hr)and up to 98% of rejection rate by using

divalent salts like Na2SO4 and nano-filtration for the DS recovery process.

Although FO applications for desalination are still in their infancy, the available bench-scale

literature reveals that by using appropriate draw solutions and suitable FO membranes,

effective desalination of seawater is practically possible (McCutcheon et al. 2005; Miller and

Evans 2006; Valladares-Linares et al. 2014; Zhao et al. 2012). A typical example in this case

is the FO pilot plant constructed and tested by a team from Yale (Cath et al. 2006). The draw

solution was formed by mixing ammonium carbonate and ammonium hydroxide, and the

formed salt species included ammonium carbonate, bicarbonate, and carbonate. The analysis

of the pilot plant work showed that by using the ammonia/carbon dioxide draw solution it

was possible to achieve osmotic pressure of 238 bar for a 0.05M NaCl feed water

concentration and 127 bar for a 2M NaCl concentration as well. Considering that the latter

concentration is equivalent to brine from the seawater desalination at 70% recovery (Cath et

al. 2006), the obtained pressure is very promising.

Modern Water is the only company so far that has implemented FO on a commercial basis.

The first direct FO plant was constructed in Gibraltar in 2008. In 2012, the company

completed and launched a larger direct FO plant in Al Khaluf, Oman. Operations of both

plants demonstrated higher than predicted water quality and lower than predicted membrane

fouling despite the bad quality of the sweater FS (Zhang et al. 2012). In a special report

devoted to practical FO applications, Modern Water (2013) mentioned that besides lower

fouling propensity and higher quality of produced water, FO process has led to energy

savings up to 30% due to pressure savings and lower chemicals use. Both plants in Gibraltar

and Oman operate as hybrid FO-RO plants where FO is used as a pre-treatment step for


seawater desalination. This hybrid process is referred to as Manipulated Osmosis

Desalination, and it is described in more details later in the chapter.

Despite the fact that the studies and pilot tests demonstrated certain positive moments for FO

applications in seawater desalination, they also reveal a number of barriers that currently exist

to wide scale practical applications of FO for desalination. Three limitations most commonly

mentioned in the literature relate to concentration polarisation as well as developing better

membranes and draw solutions. A typical example is a recently filed patent by Trevi Systems

Inc. for a FO-based desalination process that is expected to bring up to 87.5% savings in

electricity use in comparison to RO systems (Valladares-Linares et al. 2014). The process

uses a retrograde soluble solute as an osmotic agent and applies a two-stage DS recovery

process: first, a coalesce is used to produce a single phase water rich stream, which is then

purified by nano-filtration to produce fresh water (Carmignani et al. 2012). However, the

required draw solute is still not commercially available, which means that operational costs

are likely to increase substantially. The key issues related to direct FO commercialisation for

seawater desalination are reviewed later in the chapter.

2.4.4.2. Indirect FO Water Desalination

Indirect FO desalination uses wastewater as a FS and seawater as a DS. The process is

attractive because of the virtually unlimited supply of DS, which is one of the main issues in

direct FO as discussed above. Fig. 2.12 presents a schematic diagram for indirect FO

desalination process. Wastewater effluent as a FS is separated from seawater by a FO

semipermeable membrane. Because of different levels of FS and DS salinities, the osmotic

pressure differential causes water flow towards DS while wastewater contaminants are

rejected by the membrane. The diluted seawater is then transferred for an additional

desalination step that produces freshwater, while the rejected wastewater brine is directed to

post-treatment. The regeneration step of the diluted seawater can be performed by low

pressure reverse osmosis (LPRO) thereby leading to low energy requirements (Cath et al.

2010; Yangali-Quintanilla et al. 2011). Valladares-Linares et al. (2014) estimated that in

comparison to pure seawater, the osmotically diluted seawater contains nearly three times

less total dissolved solids, which leads to significant reduction of energy consumption by the

RO process.


Indirect FO desalination experiments have demonstrated good rejection of wastewater

contaminants, in particular phosphate and chemical oxygen demand (Cath et al. 2009;

Valladares-Linares et al. 2013). Valladares-Linares et al. (2013) also showed that a

submerged FO membrane module can reject up to 98% of heavy metals in wastewater,

although nitrogen compound rejection was moderate, and biopolymers and protein like

substances caused substantial fouling. These issues could be addressed by inclusion of a pre-

treatment process in indirect FO desalination. However, this step has yet to be investigated

(Valladares-Linares et al. 2014).

This paper focuses on the direct desalination approach, which is so far a more developed area

for seawater desalination. However, in general, indirect FO desalination can be a promising

process because of integration of wastewater treatment and desalination. It can be especially

attractive for the cities located along the coastlines, where there is steady unlimited supply of

seawater which can provide a more cost-effective way of wastewater treatment.

Waster brine

post-treatment

SEAWATER (DS)

Desalination of

diluted

seawater

Wastewater (FS)

FRESHWATER

Reu

sed D

S

Semi-permeable FO

membrane

Figure 2.12. Indirect FO Process for Water Desalination (adapted from Li et al. 2013)


2.4.6. Membranes for Forward Osmosis Desalination

Similarly to an RO process, membranes play one of the key roles in the success of FO

applications. There are two major demands for the FO desalination membranes: the ability to

allow water flow with minimum cross contamination from salts; and the highest possible flux

rates across the membrane (Miller and Evans 2006). The first requirement is based on the

need to minimise the environmental impact. The second requirement is based on the need to

reduce the costs of operation. In addition, membranes are expected to be stable at different

temperatures and various osmotic agents applied. Earlier versions of tested FO membranes

were produced from such materials as rubber, porcelain, nitrocellulose and animal bladders

(Baker 2004). In the 1960s, when Loeb and Sourirajan introduced their membrane, studies of

FO mostly used RO membranes for experiments (Cath et al. 2006). In the late 1990s,

however, specialised membranes were introduced for FO applications. Based on the methods

of fabrication, all FO membranes can be grouped into three large categories: phase-inversion

cellulosic membranes, thin film composite membranes, and chemically modified membranes.

2.4.6.1. Phase Inversion Cellulosic Membranes (PICM)

Phase Inversion Cellulosic Membranes (PICM) are prepared via conventional phase inversion

by using cellulose acetate as the dip-coating polymer (Zhao et al. 2012). Initially developed

for pharmaceutical applications, these types of membranes have recently been tested for

treating aqueous solutions in the FO process as well (Wang et al. 2005). Wang et al. (2007)

were the first researchers who developed and applied PICM specifically for FO applications.

Their membranes represented a type of polybenzimidazole (PBI) NF hollow fibre membrane

created by dry-jet wet phase inversion. The advantages of the membrane are self-charged

properties, great chemical stability, and solid mechanical strength (Wang et al. 2007).

However, the membranes performed poorly with NaCl, Na2SO4 and MgSO4 solutes. MgCl2

proved to work well as a draw solution; however, its relatively high price tag may not be a

suitable alternative (Achilli et al. 2009). A modified version of PICM, which significantly

improved salt rejection rates by using p-xylylene dichloride, was presented a couple years

later (Wang et al. 2009). The membrane was further optimised by applying polyethersulfone

(PES) and polyvinylpyrrolidone (PVP) to the casting solution to improve FO performance

(Yang et al. 2009).

Cellulose acetate, in general, has a number of beneficial characteristics, including high

hydrophilicity that enables stronger flux, low fouling propensity, and relatively good


resistance to various oxidants including chlorine (Geise et al. 2010; Zhang et al. 2010). These

benefits were observed as early as in 1959 (Reid and Bretton 1959; Reid and Kuppers 1959).

Because of this, cellulose acetate has been widely used in production of both FO and RO

membranes. In fact, the Loeb-Sourirajan membranes were produced by phase inversion and

used cellulose acetate polymer (Loeb and Sourirajan 1963). Such membranes have been

widely applied for energy generation through PRO (Gerstandt et al. 2008). Cellulose ester-

based hollow fibre and flat sheet membrane modules have been recently introduced by

Chung’s group (Su et al. 2010; Wang et al. 2010; Zhang et al. 2010). Preliminary tests

showed that the membranes with two layers are able to reduce the ICP effect in the

membrane support layer (Wang et al. 2010; Zhang et al. 2010). Sairam et al. (2011) also

developed and tested phase inversion FO flat sheet membranes with cellulose acetate and

lactic acid (C3H6O3), maleic acid (C4H4O4) and zinc chloride (ZnCl2) as pore-forming agents.

The zinc chloride membrane showed the best FO performance. However, cellulose acetate

membranes have a number of significant drawbacks as well. They have showed low

resistance to hydrolysis and biological attach, and perform well only within relatively short

range of temperature and pH (Baker 2004; Geise et al. 2010; Mulder 1996).

2.4.6.2. Thin Film Composite Membranes (TFCM)

TFCM have been introduced by HTI Corporation on the commercial basis several years ago

(Zhao et al. 2012). Corresponding to their definition, TFCM are the membranes that consist

of several layers. HTI’s patent for a composite FO membrane describes it as having a

minimum of three layers: a skin layer of polymeric material (8-18 µm), a porous scaffold

layer (25-75 µm), and a layer of hydrophilic support fabric (Herron, 2008). Figure2.16 shows

cross-sectional electronic microscope images of two such membranes. Both membranes are

asymmetric and made of cellulose triacetate, although they have different thickness. Most of

FO studies have been using the thinner (~50μm) type because it provides the better flux,

although it should be noted that it also rejects fewer salts. Unlike most of the commercially

available RO membranes, this membrane type is mechanically supported by an embedded

polyester mesh (HTI, 2011).


Figure ‎2.13. Cross-Sectional Electronic Microscope Image of HTI’s Membranes (Zhao and

Zou 2011)

Specifically for FO, both hollow fibre and flat sheet membranes have been developed (Song

et al. 2011; Wang et al. 2009; Wei et al. 2011; Yang et al. 2009). In many cases, the

procedure for FO TFCM production is the same as for a RO: phase inversion is used for the

preparation of a porous substrate with subsequent polymerisation to form a thin polyamide


active layer (Petersen 1993). Flat sheet TFCMs for FO were initially produced in Elimelech’s

lab by m-phenylenediamine (MPD) and trimesoyl chloride (TMC) polymerisation on porous

polysulfone (Yip et al. 2010).The manufacturers determined that an important point for

membrane performance optimisation was formation of the substrate by phase inversion where

highly porous macrovoids are covered by a thin spongelike layer (Tiraferri et al. 2011).

Wang et al. (2010) used a similar method to develop a hollow fibre FO membrane. They also

found that the optimal membrane structure should have a small sponelike layer in a thin,

highly porous substrate. Further, Wang et al. observed that for a hollow fibre FO membrane

the substrate structure was an essential factor for the performance. Experimental tests showed

that the membranes with straight fingerlike pores (Figure2.17) performed better due to

minimisation of ICP (Wang et al. 2010). Such membranes have been applied in PRO power

generation process and showed power densities up to 10.6 W/m2 (Chou et al. 2012). Widjojo

et al. (2011) used sulfonated polymer material and similar to Wang et al’s. approach for

hollow fibre membrane production. They found that such material allowed to form a sponge-

like structure for a better permeate flux.

Figure ‎2.14. Cross-Sectional Electronic Microscope Image of a TFCM with Fingerlike Pores

(Yip et al. 2010)

For FO TFCM it is typical to observe a trade-off between salt rejection and water

permeability. Wei et al. (2011), for example, reported that membrane permeability could be

enhanced by either increasing the TMC concentration or decreasing the MPD concentration;

however, this also led to a lover degree of salt rejection. A series of recently conducted

studies explored variations of TFCM in an attempt to achieve the balance between water

permeability and salt rejection levels (Bui et al. 2011; Song et al. 2011; Wei et al. 2011). The


results of these studies generally predict that in TFCM for FO the support layer largely

determines the ICP, while membrane active layer determines the rates of reverse solute flux

and the salt rejection rate. Due to relatively positive results of the studies covering TFCM it is

highly likely that these types of membranes will be further researched for FO applications. It

is also likely that the experiments will be focusing on improving salt rejection rates and

decreasing reverse solute diffusion.

2.4.6.3. Chemically Modified Membranes (CMM)

CMM represent one of the most recent innovations in the FO field. Basically, these

membranes are created by chemical enhancement of different types of membranes. For

example, Arena et al. (2011) applied a hydrophilic polymer polydopamine to improve the

performance of some commercially available RO TFC demonstrating in the process reduction

of ICP and improvement of water flux in FO tests. Setiawan et al. (2011) modified a hollow

fibre FO membrane to add a selective layer by polyelectrolyte post-treatment of

apolyamideimide microporous substrate using polyethyleneimine and suggested that because

of the positively charged property the new FO membrane was better suited for heavy-metal

laden waters. Qiu et al. (2011) created another type of a positively charged flat sheet FO

membrane that had a nano-filtration like selective layer on a woven fabric embedded

substrate. FO membranes on a layer-by-layer basis have also been introduced recently (Saren

et al. 2011). Specifically, such membranes were developed by combining phase-inversion

process to create polyacrylonitrile substrate with post-treatment process where sodium

hydroxide was used to enhance hydrophilicity and the negative charge density of the

membrane’s surface.

CMM still represent a novel, albeit very promising way to enhance the performance of the

existing FO membranes. In addition to the aforementioned membrane production methods,

Zhao et al. (2012) suggested that future FO studies could explore additional technologies

such as polyelectrolyte dip-coating, UV-photographing and layer-by-layer assembly. So far,

however, most of FO membranes are produced by using conventional techniques that have

been practically applied for pressure-driven RO and nano-filtration processes.

2.4.6.4. Selection and Evaluation of FO Membranes

The majority of the contemporary FO membrane producing methods, as is seen from above,

still rely on conventional techniques used for pressure-driven RO membranes. The process of


developing new, high performance membranes with the production techniques strictly

oriented toward FO is still in its infancy. From this standpoint, there are suggestions from FO

researchers to borrow from the solid history of RO membrane preparations and put efforts

into new and existing techniques to prepare high performance FO membranes (Miller and

Evans 2006; Zhao et al. 2012).

The choice of FO membranes will highly depend on their performance. As was noted above,

concentration polarisation, high levels of salt rejection, and anti-fouling properties are the

basic requirements to FO membranes. As was demonstrated in equations ( ), –

( ), above, osmotic water flux is the function of the water permeability (Aw), the

membrane structure coefficient (S), the solute resistivity coefficient (k), and the solute

permeability (B). Because the relationship between k and S is independent from the general

flux equation, these parameters can be determined on a separate basis. k was demonstrated as

a reflective to ICP coefficient, which increases its importance in the evaluation process.

In general, when selecting and evaluating a newly developed membrane for FO application, it

is essential to consider its ICP and fouling resisting properties, the level of water

permeability, rates of salt rejection, and its structural parameter. Some researchers also

recommend using specific reverse solute flux calculations (Hancock and Cath 2009; Zhao et

al. 2012).

2.4.7. Draw Solutions for Forward Osmosis

Basically, a draw solution (osmotic agent) is used to initiate the osmotic process. However,

during this process, water moves across the membrane surface due to the concentration

gradient from highly concentrated solution to low concentrated solution in order to achieve

the state of equilibrium between the draw solution and feed solution. Selection of the optimal

draw solution is another main challenge in FO process. According to Miller and Evans

(2006), regardless of their application, draw solutions should be inert, stable, non-toxic, and

pH neutral. They further continued that such solutions should have minimal negative impact

on the membranes, human health, or the environment, while being able to ensure high

osmotic pressures at low cost of production. The criteria such as pH and stability of the draw

solution may vary with the experimental conditions it is required that the solution has

properties of higher diffusion and low viscosity (Zhao and Zou 2011) .


Several studies have recently discussed the approaches to selecting the optimal draw

solutions (Achilli et al. 2009; Kim et al. 2012). In general, three main criteria can be

identified across the studies for selecting the optimal FO draw solution: high osmotic

pressure, ease and low cost of recovery, and minimised ICP effect (Achilli et al. 2009; Cath

et al. 2006; Kim et al. 2012; Yen et al. 2010). At the same time, the interrelation between

these factors does not always make the choice of the solution easy. For example, Zhao and

Zou (2011) demonstrated that the draw solution’s viscosity, diffusion coefficient, and

ion/molecule size have significant impact on ICP. Specifically, it is required that the solution

has properties of high diffusion and lower viscosity (Zhao and Zou 2011).

The first references to draw solution choices for FO could be traced back to the 1960s.

Batchelder (1965) patented a FO desalination process that used highly volatile gases (NH3

and SO2) diffused in water as a draw solution. Volatile gases were used for the ability to

separate or recover the thermolytic draw solution. The diluted solution was to be heated in

order to remove pure water. In the same year, Glew (1965) patented a desalination process

that used a mixture of water and either SO2 or aliphatic compounds as a draw solution. Fank’s

(1972) patent for desalination described the use of a precipitable salt (Al2(SO4)3) dissolved in

water to create a draw solution with high osmotic pressure. The salt’s precipitation was

achieved by adding Ca(OH)2 to the diluted draw solution. Unfortunately, these earlier patents

did not report on any empirical measurements such as flux, salt rejection or membrane

fouling propensity with these solutions.

Later osmotic studies started investigating sugars and simple salts as potential candidates as

draw solution agents in FO. Sugars were noted for the ability of their solutes to become easily

reconcentrated at lower pressure levels. Major applications of such solutions were in small

scale osmotic processes. Kravath and Davis (1975) proposed to use glucose as the draw

solution to apply specifically with seawater desalination. Their research, however, was

focused on emergency situations and only considered desalination to the point of dilution of

draw solution because it could be safely consumed in the short term. Based on Kravath and

Davis’ research, Kessler and Moody (1976) proposed use of a nutrient mixture with a low

molecular weight as a draw solution for osmosis based water extraction on life boats. Further,

Stache (1989) used a concentrated fructose and glycine solution for his semi-permeable bag

design that used osmosis process for producing potable water. A relatively novel approach to

draw solutes was proposed by many investigators who combined a water soluble NH3 with

CO2 containing NH4HCO3 (McCutcheon et al. 2005, 2006; Achilli et al. 2009; Martinetti et


al. 2009).It was shown that the mix is capable of providing higher water fluxes and only

moderate heat (approximately 60 ºC) is required to recover the solute. The issue with this

method, however, is the strong ammonia odour preserved in the purified water (Miller and

Evans 2006). Some other natural chemicals and even in-organics have been tested as the

draw solutes; although, according to Zhao et al. (2012), there is hardly a naturally existing

chemical that can become a perfect draw solute.

Because of the difficulties with finding a perfect naturally available draw solute, a number of

synthetic and organic materials have been proposed (Adham et al. 2007; Ling and Chung

2011; Yen et al. 2010). Especially interesting in this regard is the experiment described in

Zhang et al. (2011), where a stimuli-responsive polymer hydrogel was applied as a FO draw

solution (Zhang et al. 2011). According to the study description, based on temperature and/or

pressure levels, the gel has the capacity to either swell or shrink, thereby pulling or releasing

water. The researchers were able to achieve higher level of performance after adding light

absorbing carbon particles to the gel, which significantly improved swell/shrink ratios (Zhang

et al. 2011). A similarly innovative idea was described in a study by Liu et al. (2011), where

the initial draw solute, A2(SO4)3 was combined with CaO to create a gel like mixture of

positively charged particles Al(OH)+ and CaSO4

+. Super magnetic nanoparticles, negatively

charged, were then added to begin the sedimentation process and enhance separation

efficiency. While these innovative methods have potential as eco-sustainable approaches to

water desalination, they are still poorly developed to be considered seriously.

It follows from the discussion that there is a wide range of draw solutions available for FO

desalination process. None of the existing solutions, including the most technologically

advanced ones, seems to be perfect, however. Therefore, the quest for the optimal draw

solution continues, while it is always useful for the researchers to remember the basic criteria

that define the draw solution.

2.5. Concentration Polarisation

2.5.1. Introduction

Concentration polarisation is a common and unavoidable phenomenon in both osmotically-

driven and pressure-driven membrane processes. It is referred to as the accumulation of

excess solute particles in a thin layer bordering the membrane surface (Sablani et al. 2001;

Zhao et al. 2012). In osmotically-driven membrane processes, concentration polarisation is

caused by the concentration difference between the feed solution and the draw solution


through an asymmetric FO membrane. Inherent to all membrane processes, concentration

polarisation is one of the major factors limiting the reduction of permeate flux (McCutcheon

and Elimelech 2006; Zhao and Zou 2011). In the FO process, two types of concentration

polarisation effects occur: External Concentration Polarisation (ECP) and Internal

Concentration Polarisation (ICP). These are demonstrated in Figure 2.15.

Feed Solution Draw Solution

Jw

Cdraw

Support

Layer

Active

Layer

Cfeed

ICP ECP

∆πeff

FO Flow

Figure ‎2.15. ICP and ECP in FO Process (Zhao et al. 2012)

In the Figure, C feed stands for concentration of feed solution; C draw for concentration of

draw solution; ∆πeff for effective driving force; and Jw for water flux. The effective driving

force diminishes due to polarisation concentration profiles. Both polarisation types are

described in details below.

2.5.2. External Concentration Polarisation

ECP takes place at the surface of the active membrane layer. In general, no matter what the

nature of the free stream is, a thin layer of fluid close to the fluid-channel interface will be in

laminar flow (Welty et al. 2001). Water transport within this layer and the transport of other


solutes will be based solely on convection and molecular diffusion (Hsiang 2011). As a

result, a phenomenon of polarisation takes place, which is known as ECP. There are two

possible types of ECP depending on membrane orientation: concentrated ECP occurs when

the support layer is facing the draw solution, while dilutive ECP takes places when the active

layer faces the feed solution (Zhao et al. 2012). In general, ECP leads to the increased

osmotic pressure at the interface of the membrane active layer on the feed side, thus reducing

the net driving force. The same can happen if the osmotic pressure at the membrane active

layer surface on the draw solution side decreases (Zhao et al. 2012). According to Cath et al.

(2006), the negative impact of ECP on the flux can be decreased by either optimising the

water flux or by increasing the overall flow velocity. They further stated that optimisation of

water flux may not be suitable in all cases due to low initial flow level.

The effects of ECP on water flux can be modelled by using the general equations for

pressure-driven membrane processes (Mulder 1996). The basic concentration polarisation

equation is expressed as:

(

) Eqn.2.9

where :

Cm is the concentration of the feed solution at the membrane;

Cb is the concentration of the feed solution in the bulk;

Jw is the water flux.

k in the equation is the mass transfer coefficient, which depends on the properties of the

applied solutions and hydrodynamics of the systems. Typically, such coefficients are

estimated by using empirical correlation of dimensionless numbers, such as Sherwood

number, Reynolds number and Shmidt number. Eqn. 2.10 shows the general mass transfer

coefficient calculation for concentration polarisation.

where:

D represents the solute diffusion coefficient;

Dh represents the hydraulic diameter.


The general correlation formula for Sherwood number can be presented as:

Eqn 2.11

where:

Re is Reynolds number;

Sc is Schmidt number.

Depending on geometry and conditions of the flow, different K, a and b are applied which

leads to a great variety of possible correlations. No single Sherwood number can describe all

systems for mass transfer coefficient calculations, and up to 27 Sherwood relations are

recognised in literature (van den Berg et al. 1998). The most commonly used Sherwood

relationship for RO and ultrafiltration processes, according to Mulder (1996), are:

Sh = 0.04Re0.75

Sc0.33

for turbulent flow systems

and

Sh = 1.85(Re*Sc*Dh/L)0.33

for laminar flow systems where

L is the characteristic length.

McCutcheon and Elimelech (2006) applied these correlations when measuring the mass

transfer coefficient for modeling ECP in FO with satifactory results which makes the

aforementioned correlations suitable for ECP modeling in FO processes.

In FO, the concentration of the feed solution is small, which makes it possible to replace it by

the osmotic pressures. Therefore, equation [Eqn.2.9] becomes:

(

)

Where:

πmfeed is the osmotic pressure of the feed solution at the membrane;

πbfeed is the osmotic pressure of the feed solution in the bulk;

kfeed is the mass transfer coefficient on the feed side.

Consequently, the equation for the dilutive ECP in FO can be expressed as:

(

)


Where:

πmdraw is the osmotic pressure of the draw solution at the membrane;

πbdraw is the osmotic pressure of the draw solution in the bulk;

kdraw is the mass transfer coefficient on the draw side.

The general water flux equation (Eqn.2.1) for FO replaces the value of (σ∆π − ∆P) with the

osmotic pressure differences from the solutes at the membrane surfaces to become:

Jw = A(πmdraw – πmfeed) Eqn.2.14

By inserting the values for πmdraw and πmfeed from equations (Eqn.2.12) and (Eqn.2.13)

respectively, the final flux process for ECP is expressed by the following equation:

(

) (

)

Equation (Eqn.2.14) applies both concentrative and dilutive ECP to model the water flux.

This is a simplified equation, however, because it makes such assumptions as the negligible

value of the solute permeability coefficient, similar mass transfer coefficients on the feed and

draw solutions, and the presence of a dense symmetric film. In practice, asymmetric FO

membranes are used more often, which makes the effects of ICP much more important.

2.5.3. Internal Concentration Polarisation

ICP is similar to ECP process with the only difference that it takes place within the

membrane’s porous layer; therefore, cross-flow cannot diminish its negative effects on the

flux (Cath et al. 2006). ICP is considered one of the major issues in osmotically-driven

membrane processes. Studies show ICP is the major factor causing decline in water flux

(Gray et al. 2006; McCutcheon and Elimelech 2006). In similar manner to ECP,

concentrative and dilutive ICP can occur depending on the orientation of the membrane.

Figure 2.14 demonstrates both processes: dilutive ICP occurs when the active membrane

layer faces the feed solution, while concentrative ICP occurs when the active membrane layer

faces the draw solution.


Draw Solution (DS) side Membrane Feed Water (FW) side

DS: Osmotic pressure in bulk DS side.

DS-m: Osmotic pressure at the membrane interface in DS side.

i: Osmotic pressure at the transition interface between membrane layers.

FW-m: Osmotic pressure at the membrane interface in FW side.

FW: Osmotic pressure in bulk FW side.

DECP, CECP, DICP, CICP: Dilutive External, Concentrative External, Dilutive internal and

Concentrative Internal Concentration Polarization, respectively.

DS-m FW-m

DS

FW

DS-m

i

FW-m

DS

FW

Skin layer Porous substrate

DS-m

i

FW-m

DS

FW

Porous substrate

Skin layer

(a):

Symmetric dense

membrane

(b):

Asymmetric

membrane with the

dense skin layer

against the draw

solution

(c):

Asymmetric

membrane with the

dense porous

substrate against

the draw solution

Water flow

DS Salt flow (diffusion)

Water flow


Water flow


DECP

CECP

DECP

CICP

DICP

CECP

Figure ‎2.16. Dilutive External, Concentrative External, Dilutive Internal and Concentrative

Internal Concentration Polarization (Mezher et al. 2011; Schiermeier 2008; Younos and Tulou

2005)

The adverse effect of ICP on water flux and permeability rates is a well-known issue in

retarded pressure driven osmotic processes. Mehta and Loeb (1979) studied the effect of the

porous support layer on ICP and the effect of high draw solution concentrations on the


overall permeability coefficient of the membrane. They showed that upon swapping the

working fluids on the two sides of the membrane, flux sharply declines due to ICP (Fig.

2.17).

Figure ‎2.17. Water Permeability Cefficient as a Function of Draw Solution Concentrations

(Mehta and Loeb 1979)

Evaluation of the ICP effect on the FO flux is commonly done with the equations of the

classical solution/diffusion theory (Mulder 1996). The dilutive ICP effect is represented by

the following equation:

where :

Jw is the water flux;

A is the water permeability constant depending on the membrane used;

πdraw is the osmotic pressure of the draw solution;

πfeed is the osmotic pressure of the feed solution;

B is the solute permeability coefficient of the membrane;

k is a resistivity coefficient reflecting the degree of ICP in the support layer.

The equation for concentrative ICP is expressed in similar way:


For both processes, smaller k coefficient translates into lower ICP and higher water flux. The

coefficient is calculated as:

where :

D stands for the solute’s diffusion coefficient through water;

S for the membrane’s structural parameter;

t, τ, and ε are the thickness, tortuosity, and porosity of the support layer of membrane,

respectively.

As is seen, S is the product of the three membrane parameters, and its value is important to

know for any newly developed membrane.

It is important to remember that ICP and ECP usually occur at the same time. Consequently,

for the FO process, the effects of concentrative ECP and dilutive ICP on the flux can be

determined by the following equation (MacCutcheon and Elimelech 2006, 2007):

(

) ( )

Based on equation( ), ICP is dependent on membrane properties such as thickness,

porosity and tortuosity as well as on diffusion solute properties such as diffusion coefficient

of the solute. Zhao and Zou (2011) proposed that ICP could also be dependent on

constrictivity, which is a factor of the ratio of solute molecule diameter to the membrane pore

diameter:

Consequently, constrictivity is derived from an empirical equation by Beck and Schultz

(1970):

δ = (1 - λ)4

Eqn.2.21


According to Zhao and Zou (2011), constrictivity defines how diffusion is slowed down due

to viscosity increase in the porous structure as a result of greater proximity to the average

pore wall.

k coefficient from equation ( ) is then calculated as:

Other models for ICP estimation have been proposed. For example, Tan and Ng (2008)

proposed that in order to account for higher concentration solutions to avoid over-prediction

of the water flux at high NaCl concentrations, a modified film theory should be used. They

consequently reported that their model was in better agreement with the data obtained

experimentally and that it applied to a wider range of solutes. However, the derived model is

also much more complex than the one proposed by McCutcheon and Elimelech (2006, 2007)

and is beyond the scope of this report.

Due to the importance of ICP for osmosis process investigations, there have been a number

of studies to cover this topic recently (Gruber et al. 2011; Li et al. 2011; Qin et al. 2010; Tang

et al. 2010; Zhao and Zou 2011). Qin et al. (2010) found that dilutive ICP was an important

obstacle to the osmosis process application due to its ability to reduce the water flux by

99.9%. Tang et al. (2010) studied the impact of ICP on FO flux behaviour and observed that

ICP’s negative effect on flux increased substantially with membrane fouling. Zhao and Zou

(2011) also found that the ICP effects on flux in FO increases with the increase of draw

solution concentrations (Figure 2.18).


Figure ‎2.18. Theoretical, FO Mode and PRO Mode Water Flux as a Function of net Osmotic

Pressure (Zhao and Zou 2011).

As is seen, CP is an important factor to consider in FO applications for desalination. It has

been shown that CP is dependent on membrane structure; therefore, choice of the right

membranes plays an essential role in spearheading FO desalination research and practical

applications.

2.6. Manipulated Osmosis Desalination (MOD)

2.6.1. Process Description

The Manipulated Osmosis Desalination (MOD) process was developed at the University of

Surrey’s Centre for Osmosis Research and Applications (CORA). The process is comprised

of FO desalination combined in a single cycle with a regeneration step. Specifically, the

principle behind the MOD is manipulation of two fluids with different osmotic pressures to

receive a pure water flux across the selectively permeable membrane (Al-Mayahi and Sharif

2004). The manipulated fluids are seawater and a chosen draw solution. The seawater passes

through a FO unit with a draw solution, which attracts water molecules. The resulting liquid

is then transferred to a membrane-based (RO or nano-filtration) recovery unit where final

separation of the product water takes place. Figure 2.19 presents a basic schema of the MOD


process using RO unit for the regeneration step. A semi-permeable FO membrane is placed

between the feed solution (seawater) and the draw solution (water containing osmotic agent).

The higher osmotic pressure from the draw solution allows to extract water from the feed. As

in other membrane processes, the result is rejected feed brine, which is discharged, and the

diluted draw solution. The diluted draw solution is then transferred to the regeneration unit

where RO process is used to extract water. The rejected osmotic solution is then transferred

for reuse in the first step.

Figure ‎2.19. Continuous MOD Process with Regeneration Step (Adapted from Al Mayahi and

Sharif 2004)

2.6.2. Advantages and Considerations

The MOD allows to capitalise on the advantages of FO and RO processes while at the same

time reducing their weaknesses (Table 2.2). If FO is used as a pre-treatment step in

desalination, it can on the one hand dilute the feed stream without energy consumption and,

on the other hand, help avoid irreversible membrane fouling. In the second step, low pressure

RO will be required, although it will enable steady flux and high quality water product.


Table ‎2.2. Comparison of FO and RO (Adapted and Expanded from Chung et al. 2012; Zhao

et al. 2012)

Process Advantages Disadvantages Challenges

Forward

Osmosis

-low energy

requirements for water

transport

-reversible fouling

-low operational cost

-possibility of energy

production

-few working projects;

-does not work well as a

standalone process;

-ICP;

-requires relatively high

capital costs

-increase water flux;

-reduce ICP;

-develop effective draw

solutes;

-produce effective

specialised membranes;

Reverse

Osmosis

-mature technology

-low capital costs

-high water quality

-hydraulic pressure

consumes a lot of energy;

-fouling is irreversible;

-operational costs are

high;

-reduce energy losses;

-reduce membrane

fouling;

In comparison with the traditional RO-based seawater desalination process that operates at

60-80 bar, the MOD operates at greatly reduced pressures of 2-3 bar. While RO as a recovery

step consumes additional energy, the consumption can be minimised with a careful selection

of the osmotic agent and optimisation of operative conditions. Additionally, these factors can

lead to greater recovery efficiency due to control of the draw solution composition and

absence of foulant impurities (Choi et al. 2009). Other advantages of the MOD process are

lower costs due to membrane longevity and lower fouling propensity; reduced capital

expenses due to significant reduction of contaminants; higher quality water product due to a

double membrane barrier; and possibility of using lower pressure pipe works and fittings

(Achilli et al. 2009; Boo et al. 2013; Thompson and Nicoll 2011). Recent studies have also

demonstrated the potential for hybrid FO-RO systems to recover impaired water from a

recycle feed (Cath et al. 2009; Yangli-Quintanilla et al. 2011).

In order to better understand economic benefits of MOD, it is useful to consider capital and

operational expenses from a comparative perspective. Such analysis was provided by Nicoll

(2011) who described the advantages of MOD over RO for desalination on the bases of

capital investment, power consumption and operational costs. According to Nicoll (2011), a

typical MOD system would require roughly the same amount of initial capital as RO if

operated on the same feed source. While overall capital expenses were difficult to project for

equivalent plants, Nicoll mentioned that an MOD plant can at least save on basic

components. Unlike the traditional RO plant that used duplex stainless steel components, an

MOD plant could use plastic components that cost less.


A more definitive analysis provided by Nicoll (2011) was for operating expenses of MOD

and RO plants. The following assumptions were made for both plants: feed water temperature

is 25 oC; the same feed water pre-treatment requirements; overall plant efficiency at 70%; and

energy recovery efficiency at 70% as well. A conservative assumption was made with regards

to MOD with a 30% conversion rate. Cost of osmotic agent was taken as $75 per tonne. For

RO plant the conversion rate was between 41% to 80% based on the total dissolved solids

(TDS) in the feed. A low estimation was made for energy cost at 0.075 kW/hr. Nicoll’s

energy consumption and total operating cost projections were based on the amount of TDS.

Since MOD does not require hydraulic pressure to operate in the pre-treatment stage, its

power consumption remained lower regardless the level of TDS in feed water, while RO

process required an increasing amount of power to operate. However, for MOD, osmotic

agent cost and the cost of osmotic agent recovery were relevant. Figure 2.20 and 2.21 show

Nicoll’s (2011) projections for power consumption and operating costs respectively based on

TDS in feed water. It can be seen that in comparison to RO MOD requires lower power

consumption and operating expenses, and that the difference increases as the feed gets more

challenging with higher amount of TDS.

Figure ‎2.20. Projected Power Consumption of MOD and RO Plants (adapted from Nicoll

2011)

0

1

2

3

4

5

6

7

10000 20000 30000 40000 50000

kW/h

r

Total Dissolved Solids in the Feed

RO

MOD


Figure ‎2.21. Projected Operating Costs of MOD and RO Plants (adapted from Nicoll 2011)

Despite its promising parameters, MOD has yet to mature as a wide scale desalination

technology. As in other membrane processes, MOD needs to be optimized to achieve the best

efficiency and product quality. Optimisation of the MOD process depends on a number of

factors such as number of performance of FO membranes, composition, concentration and

recirculation rate of the draw solution, performance of the regeneration step and feed water

factors such as temperature, composition and flow rate. The type of membrane and the choice

of an osmotic agent in the draw solution can further increase efficiency of the process.

Specifically, a properly selected osmotic agent may enable use of lower resistance

membranes that require lower feed hydraulic pressures while membranes with higher

permeability of water may lead to a more efficient separation process.

2.6.3. Existing MOD Projects

2.6.3.1. University of Surrey Rig

As was mentioned earlier, MOD process was developed and patented by a group of

researchers at the University of Surrey’s Centre for Osmotic Research and Applications

(CORA), which is a part of Modern Water’s development programme. The test rig installed

at the University’s lab became the first operational MOD unit, and it has been used for

research and development since, including identification of membrane configuration and

parameters, osmotic agents and other relevant processes. In 2008, based on the trials

0

0.05

0.1

0.15

0.2

0.25

0.3

0.35

0.4

0.45

0.5

10000 20000 30000 40000 50000

Op

era

tin

g C

ost

Total Dissolved Solids

RO

MOD


performed at the University of Surrey, Modern Water commissioned a practical MOD system

in Gibraltar. The key objectives of the plant were to demonstrate a possibility of stable

continuous MOD cycle operations; practically test the theorised models for MOD cycle

performance; and analyse real-world issues related to MOD that cannot be apparent in

laboratory settings (Thompson and Nicoll 2011).

2.6.3.2. Gibraltar MOD Pilot Plant

The MOD system in Gibraltar has been operational since May 2009. The system applies the

same process structure as the test rig at the University of Surrey, although on a larger scale

with 18 m3/day output (Figure 2.22). The rig uses the same seawater as the RO plant located

nearby, and the water has the silt density index between 3 and 4. The treatment process

involves the FO unit for desalination and RO for osmotic agent regeneration. Within the first

two years of operations, several on site modifications were implemented, mostly to improve

the FO step. Membrane modifications were introduced which allowed to improve the

facility’s productivity by 30% thereby reducing capital expenses and environmental footprint

of the plant. The product water supplied by the Gibraltar MOD facility has less than 200 TDS

mg/L and less than 0.6 mg/L of boron level.

Figure ‎2.22. Gibraltar MOD Rig by Modern Water Prior to Containerisation (Thompson and

Nicoll 2011)


Because of the convenient location next to a RO plant, it was possible to run a number of

tests for comparative performance of the MOD plant’s RO step as well. It was found that

MOD plant allowed for a more efficient use of RO membranes mostly because of the ability

to control the regeneration unit feed solution coming from the FO step. It was estimated, for

example, that the MOD plant’s cycle allows to achieve 50% feed water conversion with only

three membrane elements whereas RO plant required 6-7 elements to achieve a 40% recovery

rate (Thompson and Nicoll 2011). An important advantage of the MOD system installed in

Gibraltar, according to Thompson and Nicoll (2011), has been the ability to modify the

product quality by changing the properties of the osmotic agent. Further, focused

modifications to the operating flow conditions allowed to reduce boron levels by over 40%.

2.6.3.3. Production Facilities in Oman

Following the success of the Gibraltar plant, Modern Water deployed a larger scale (100

m3/day) MOD plant in Al Khaluf region, Oman. The site already had an operating RO plant

of the same capacity, and both plants would share pre- and post-treatment equipment (Figure

2.23). The choice of the site was made based on the primary goal of the project, which was to

directly compare the performance of a MOD-based plant to the performance of a

conventional RO plant of the same capacity and operating in the same conditions. Other goals

included increasing the region’s water supply while observing operations of the plant under

harsher conditions. The location of the plant was notable for the poor quality feed water and

harsh environmental conditions such as soaring temperatures (up to 50 oC), constant winds

and sand storms. Additional challenges included unplanned power outages and additional

disruptions associated with external services.


Figure ‎2.23. MOD Plant Facility in Al Khaluf, Oman (Thompson and Nicoll 2011)

Despite the environmental and operational challenges in Al Khaluf, the MOD plant

demonstrated the superior performance in comparison to the RO plant. The major practically

demonstrated advantage was reliable membrane performance. Figure 2.24 shows a

comparative chart for RO and MOD membrane performances during 2010. As is seen from

the charts, the RO plant membranes showed decreasing normalised permeate flow with up to

30% decline in the 5 month period. In contrast, the normalised output for the FO system

shows relatively stable flow even without membrane replacement and chemical cleaning.

Other advantages of the MOD plant design over the RO facility included:

- higher feed water recovery: 35% for MOD versus 25% for RO;

- higher water flow: 4.2 m3/hr for MOD versus 3.0 m

3/hr for RO;

- lower feed water pump energy consumption: 1.6 kWhr for MOD versus 25.3 kWhr

for RO;

- lower overall specific energy consumption per final product unit: 4.9 kWhr/m3 for

MOD and 8.5 kWhr/m3 for RO. (Thompson and Nicoll, 2011)


Figure ‎2.24. Normalised Flow for Al Khaluf RO Plant (above) and MOD Plant (below) in

2010 (Thompson and Nicoll 2011)

The proven operational record for Al Khaluf MOD project prompted Omani government to

seek a larger facility construction. A ₤420,000 contract was signed with Modern Water to

install and operate the first MOD commercial facility in Al Najdah, and the project was

completed in 2012 (Lauwers 2012). Al Najdah plant has the capacity to produce 200 m3 of

water per day (Water World 2013).

2.7. Chapter Summary

This chapter presented a comprehensive review of the existing literature on desalination with

the focus on forward osmosis process and its current commercial applications. With the

increasing demand for freshwater because of population growth and industrial development,

searches for unconventional ways to produce water intensified in the past decades. Osmotic

processes have received a particular interest due to their potential for efficient conversion of

seawater into drinkable product. So far, the majority of commercial osmotic plants have used

reverse osmosis process. However, this process requires substantial energy use and

propensity for membrane fouling. Forward osmosis allows to alleviate these issues because it


is based on a natural process of osmotic pressure to extract water from the feed. As such, it

does not require energy to operate and is not prone to irreversible membrane fouling.

Forward osmosis found a wide range of applications in such areas as energy production,

wastewater treatment, pharmaceutical and food industries. Its applications for seawater

desalination, however, have remained limited until recently. The University of Surrey Centre

for Osmotic Research and Applications pioneered a novel desalination process called

manipulated osmosis desalination (MOD). MOD relies strongly on forward osmosis to dilute

the seawater feed with the subsequent treatment of the diluted solution by reverse osmosis.

Such system ensures substantial energy savings, lower environmental impact and lower

operating costs in comparison to the same capacity reverse osmosis plats. This was

successfully demonstrated at fully operational plants in Gibraltar, Al Khaluf and Al Najdah.

MOD was especially effective in the areas with poor quality of seawater feed and harsh

environmental conditions.

Despite the recent commercialisation of MOD, there are still issues that require attention in

order to increase efficiency of the process. These issues involve internal concentration

polarisation, membrane development and search for effective osmotic agents. This paper adds

to the existing knowledge of MOD and its applications by testing different membranes’ and

osmotic solutions’ performances.


3. CHAPTER THREE: MATERIALS AND METHODS

3.1. Introduction

This chapter describes the materials and methods used in to investigating the efficiency of

organic osmotic agents in the forward osmosis desalination process. Two types of

membranes: flat sheet and hollow fine fibre are used. Osmotic pressure measurements were

taken using OLI Analyser Software. The structure of the Chapter follows the experiments

conducted for the different types of membranes. The first section of the Chapter (section 3.2)

describes the flat sheet FO bench scale laboratory apparatus and the dissembled membrane

testing apparatus, the flat sheet membranes used for the experiments and the basic parameters

under which the experiments were conducted. The second section of chapter (section 3.3)

describes the FO rig used for hollow fine fibre membrane measurements, the combinations of

osmotic agents and draw solutions and the parameters under which the experiments were

conducted. The final section of the chapter (section 3.4) describes the criteria used for

calculation of FO performance.

3.2.Flat Sheet Membrane Study

3.2.1. Experiment Outline

In this series of experiments, forward osmosis was studied using two types of membranes

(nano-fitration and reverse osmosis) and two types of osmotic agents (glucose and sucrose)

against deionized water (DW). Concentration, flow rate and temperature were the variable

parameters tested in the experiment .The experimental framework of this work is presented in

Figure 3.1.


FO Process

NF

Membrane

RO

Membrane

SucroseGlucoseSucroseGlucose

Concentration

Flow rate

Temperature

Concentration

Flow rate

Temperature

FO Flat Sheet Membrane Process

Figure ‎3.1. Experiment Outline for FO Using Flat Sheet Membranes

3.2.2.Experiment Set Up

The membrane module used for the experiment was designed and prepared at the Centre for

Osmosis Research and Application (CORA), University of Surrey, England. Figure 3.2

presents the setup of flat sheet FO module. Figure 3.3 shows the schematic diagram of the

bench-scale laboratory apparatus; and demonstrates the disassembled membrane testing

apparatus with flow channels. Note that the valves [14 and 15] were fully open in order for

the feed and draw solution to have full contact with both the active layer and support layer of

the membrane surface for the FO process. However, the flow rate was controlled by valve [4]

and [6] manually with the help of rotameter installed on the rig. This ensured the presence of

the close loop system without any pressure drops.

From Figure 3.4, it is seen that the experimental FO cell consisted of two compartments for

the FO process: flat sheet cell manufactured from Perspex material and the unit also had two

polyethylene tanks with 10 litres capacity each. Tank A was used as osmotic agent storage,

and Tank B was used as feed water storage. Both tanks were placed on 3 digits balances

supplied by (KERH and SOHN GmbH) to determine the amount of water flux. The balances

were capable for measuring mass changes to an accuracy of one gram. In order to control the


temperature, each tank was equipped with an immersion heating coil connected to supply by

a Cole-Palmer temperature conductor. Two small centrifugal pumps (Aqua Ltd with capacity

of 0-10 L/min) were used to pump the solutions into the FO cell. Each stream of the FO cell

was equipped with a pressure gauge and rotameter to measure the hydraulic pressure and the

flow rate of solutions, respectively. All the pipes, fittings and valves were made of

polyvinylchloride (PVC). The unit was designed to recycle the outlet solution of osmotic

agent and feed water to their feed tanks in order to reduce the materials consumption. A

magnetic stirrer (IKA RCT) was used to mix and prepare osmotic solutions. OLI Analyser

Software was used to determine the osmotic pressure for solutions. A conductivity meter

supplied by (Mettler-Toledo Ltd) was used to measure the conductivity and salinity of feed

water.

Figure ‎3.2.Design and layout of experimental membrane module for FO process

Tank B Tank A

FO Membrane

Analytical

scale

Pump

Pump

Temperature adjustable Heater

Flow meters

Analytical

scale

Pressure gauges


5

1314

6

8

1

7

10

15

9

16 17 1819

1 Tank B (Draw solution) 6 Rotameter (Feed water in) 11 Digital Balance

2 Digital Balance 7 Rotameter (Feed water out) 12 Tank A (Feed water)

3 Rotameter (Draw out) 8 Pump 13 FO membrane cell

4 Rotameter (Draw in) 9 Heat regulator 14-15 Valves

5 Pump 10 Heating coil 16-19 Pressure Gauge

Figure ‎3.3. Schematic Diagram of theFlat SheetForward Osmosis Bench-Scale Laboratory

Apparatus

Figure ‎3.4. Disassembled Membrane Testing Apparatus Showing Flow Channels.

3.2.3. Process Operation

The flat sheet membrane with a surface area of 0.0621 m2 was placed between the two

compartments of FO cell. The membrane test cell assembly was operated in a counter-current

flow mode arrangement with the osmotic agent and process solution flows each set at 10

L/min. The experiments were carried out using deionised water as the feed solution (FS) for

two types of flat sheet membranes: nano-filtration and reverse osmosis. The feed solution was

placed against the active layer (AL),whereas the osmotic agent solution was placed against

the support layer (SL). Both osmotic agent solution and feed water solution were run in


closed loop mode to diminish the internal concentration polarization (ICP) and chemical

consumption for osmotic agent (Esperanza et al.,2009).

The effect of operating parameters such as the osmotic agent concentration, flow rates of

osmotic agent and feed water and temperature on water flux (Jw), water recovery percentage

(%R), water permeability (Aw), specific energy consumption (SEC) and solute flux(Js) were

studied. The specifications of NF and RO membranes used in the experiments are

summarised in Table 3.1 and 3.2 respectively.

Table ‎3.1. TFC-SR2 Specifications (Koch Membrane Systems Ltd ,UK)

Type of membrane Low pressure, selective rejection, NF element

Membrane chemistry Proprietary TFC membrane

Typical operating pressure 3.45 – 7 bar

Maximum operating pressure 34.5 bar

Maximum operating temperature 45°C

Rejection 97.53

Flux (GFD) 57.2

Allowable pH – continuous operation 4-9

Table ‎3.2. TFC-ULP specifications (Koch membrane systems Ltd ,UK)

Type of membrane Ultra-Low Pressure RO element

Membrane Chemistry Proprietary TFC polyamide

Typical operating pressure 7-12 bar

Maximum operating pressure 24 bar

Maximum operating temperature 45°C

Rejection 98.68

Flux (GFD) 29.91

Allowable pH – continuous operation 4-11

Test conditions: 2000 mg/L NaCl solution, 8.6 bar, 15% recovery, 25°C and pH 7.5


3.2.4. Osmotic Agents

Two organic osmotic agents (sucrose and glucose) were used to create draw solutes with

different concentrations. The solutions were prepared by dissolving the osmotic agents in

deionised water.

Sucrose is a disaccharide composed o f -D-glucose and -D-fructose as monosaccharides

with chemical formula (C12H22011) and commonly known as table sugar. It has molecular

weight of 342.30 g/mol and density of 1.59 g/cm3. For this experiment, sucrose powder was

supplied in 5kg bag by the Tate and Lyle Company in the UK.

Glucose (C6H12O6) [also known as D-glucose, dextrose, or grape sugar] is a simple aldosic

mono saccharide found in plants. It is one of the three dietary monosaccharides, along with

fructose and galactose. Glucose’s molecular weight is 180.16 g/mol and its density is 1.54

g/cm3. For this experiment, 5kg containers of glucose were supplied by Sigma-Aldrich

Company located in the UK.Table 3.3 summarises the physic-chemical properties of sucrose

and glucose.

Table ‎3.3. Physico-Chemical Properties of Sucrose and Glucose

Property Sucrose Glucose

Chemical Formula C12H22O11 C6H12O6

Molecular Weight (g/mol) 342.30 180.16

Density (g/cm3) 1.59 1.54

Solubility in water at 25oC (g/l) 2074 2000

The osmotic pressures of the solutes were evaluated by using OLI Analyser's software.

3.2.5. Osmotic Agent Concentration Effect Study

To measure the effect of osmotic agent concentration in FO process, sucrose and glucose

solutions with different concentrations as shown in Table 3.4 were used. Sufficient amount of

sucrose solution at different concentrations ranging from 170 to 400 g/L was prepared and

placed in osmotic agent tank, while adequate amount of deionized water with conductivity of

10-15 μS/cm was placed in feed water tank. After filling both compartments with solution by

pumps, the samples of solution were collected and the weights of osmotic and feed solutions

in the respective tanks were recorded. The osmotic agent solution with feed water were

http://en.wikipedia.org/wiki/Aldose

http://en.wikipedia.org/wiki/Monosaccharide

http://en.wikipedia.org/wiki/Fructose

http://en.wikipedia.org/wiki/Galactose


pumped into the FO cell at the same time, while, the temperature(25 °C) and the flow rate of

osmotic agent solutions(1.5 L/min) and the flow rate of feed water solution(2.5 L/min) were

kept constant throughout the run. After 2 hrs, the run was stopped and the weight of solutions

was recorded and the samples of solution were also collected. The above mentioned

procedure was repeated using glucose solution of different concentrations ranging from 90 to

250 g/L. The concentration of sucrose and glucose were determined using HPLC equipment.

Table 3.4. Sucrose and Glucose Concentrations for the Concentration Effect Study

Osmotic Agent Concentrations g/L Osmotic pressure bar

Sucrose

Run 1 170 10.7

Run 2 200 12.6

Run 3 250 17.9

Run 4 400 27.5

Glucose

Run 1 90 12.3

Run 2 120 21.3

Run 3 180 25

Run 4 250 28.1

3.2.6. Feed Water Flow Rate Effect Study

To investigate the effect of OA solution flow rate at constant feed deionised water flow rate

(1.5 L/min) several experiments were carried out. In these experiments 275 g/L sucrose OA

concentration was prepared and pumped to the FO membrane cell at different flow rates from

2-4 L/min accordingly. However, in feed water flow rate effect study different feed deionised

water flow rates from 2-4 L/min accordingly were used while the OA concentration (275 g/L)

and the temperature (25oC) were kept constant during these experiments. Similarly in glucose

osmotic agent flow rate effects study, glucose solution with concentration ranging from 90-

250 g/L was prepared and pumped to the FO membrane at different flow rate values from 2-4

L/min accordingly, while the feed deionised water flow rate was kept constant (1.5

L/min) throughout these experiments .However , in feed water flow rate study, different

feed water flow rates from 2-4 L/min accordingly against (275g/L)glucose OA concentration

at constant flow rate (1.5 L/min) were used. The abovementioned experiments were

conducted at 25°C.


3.2.7. Temperature Effect Study

In this study, several experiments were carried out to investigate the effect of sucrose OA

temperature at constant feed deionised water temperature 25°C and the flow rate (2 L/min).

In these experiments, sucrose OA solution with a concentration of 200g/L was prepared and

pumped to the FO membrane at different temperatures(20-35°C) for 2hr.However,in feed

water temperature effect study,(2 L/min) feed deionised water at different temperatures (20-

35°C )were used, while, OA concentration (200g/L), temperature (25°C) and the flow rate

(2.0 L/min) were kept constant throughout these experiments. Similarly, in glucose OA

temperature effect study, glucose solution with a concentration of 90 g/L was prepared and

pumped into the FO cell membrane at different temperatures (20-35°C), while the feed

deionised water flow rate (2.0 L/min) and the temperature (20°C) were kept constant

throughout these experiments. However, in the feed water temperature study, different feed

water temperatures (20-35°C) at the constant flow rate (2.0 L/min) were used against 90 g/L

glucose OA concentration at temperate 20°C

3.3. Hollow Fine Fibre Membrane Study

3.3.1. Experiment Outline

In this series of experiments, forward osmosis (FO) process efficiency was tested using a

hollow fine fibre membrane, three types of osmotic agents (glucose, sucrose and sodium

chloride) and their combinations were selected. Figure 3.5 presents the experimental

framework.


HFFM

Mixture

Glucose + NaCl/

DW

Sucrose / DI

Mixture

Sucrose + NaCl /

DI

Glucose / DI NaCl / DI

Sucrose / SWMixture

Sucrose + NaCl/

BW

Mixture

Glucose + NaCl/

BW

Glucose / SW

Figure 3.5. Experiment Outline for Hollow Fine Fibre Membrane

3.3.2. Experiment Set Up

For the experiment involving the hollow fine fibre membrane, a FO unit shown in Figure3.6

was used. Figure3.7 depicts the flow diagram of the unit. The OA in the experimental unit

was feed from a 30 litre capacity container whereas feed solution (FW) was feed from a 100

litre capacity container. Two centrifugal pumps were used to transfer feed water and osmotic

agents to the FO membrane. Pump-1 was connected to Din to pump the draw solution into the

fibre side of the membrane, whereas Pump-2 was connected to Finto pump the feed solution

into the shell side of the membrane. The diluted OA was transferred to Container Dout with

the capacity of 60 litres whereas the stripped of water feed solution was transferred to

Container Fout with the capacity of 60 litres. The temperature of the containers Din and Fin was

kept constant as 25 oC and controlled by an electrical heating coil with thermostat installed in

each container. All four containers were made of stainless steel and equipped with a level

glass to measure the level of solution inside tanks

To measure the flow rate of the solutions, four flow rotameters were connected to each inlet

and outlet line of the membrane. All these lines were equipped with pressure gauges and

temperature indicators. The whole FO unit was fixed onto a skid as shown in Figure 3.6

Overall, the FO process as described above was operated in batch mode for a maximum time

of 30 minutes.


Figure ‎3.6. The Experimental Unit Fixed onto aSkid.


Figure ‎3.7. Flow Diagram of The Hollow Fine Fibre Forward Osmosis Process

3.3.3. Process Operation

A hollow fine fibre membrane (HFFM) with an effective area of 4.5 m2 was used. The

membrane, DURASEP®, was supplied by Toyobo Co., Ltd (Japan), and its basic

characteristics are summarised in Table 3.5. Figure3.8 shows a close up photograph of the

membrane within the experiment unit.

F I

L I

T I

P I

1

2

3

4

Osmotic agent tank

Draw out solution tank

Feed in solution tank

Feed out solution tank

Pump

control valve

Flow indicator

Level indicator

Temperature

indicator

Pressure indicator

LEGEND

5 Forward osmosis unit


Table ‎3.5. Specifications of the Hollow Fibre DURASEP® Membrane Produced by

ToyoboCo., Ltd (Japan).

Type of membrane Hollow fibre, FO

Membrane chemistry Cellulose triacetate

Housing material Polysulphide

Operating pressure Shell side 6 bar

Fibre side < 1 bar

Allowable pH 3-8

Membrane area 4.5 m2

Rejection 96% - 98 %

Figure ‎3.8. The Experiment Hollow Fine Fibre Membrane, FO in Operation

Initial temperature and initial water level of both feed and osmotic agents were recorded.

Water samples were collected from FW and DS containers before starting the experiment.

Also, the flow of fluids in both the channels was allowed to stabilise for 10-15 minutes to

ensure flow rate uniformity of different solutions. Each experimental run was carried out for

30 minutes, then the pumps were turned off and the process was allowed to settle to normal

conditions. Then the solution level was recorded in all the four containers (Fin, Fout, Din and

Dout). At the end of each experimental run, water samples were collected from Fout container

and Dout container.


3.3.4. Osmotic Agents

The main objective of the experiment was to exam the efficiency of the FO unit using

different osmotic agents and their concentrations at constant temperature. The performance

was evaluated based on water flux, water permeability and water recovery percentage.

Description of the osmotic agent solutions and experimental measurements are provided

below.

Two organic osmotic agents (sucrose and glucose) and one inorganic osmotic agent (sodium

chloride) were used to create draw solutes with different concentrations. Physico-chemical

properties of sucrose and glucose were described in Section 3.2.4, as they were used for the

flat sheet membrane experiment as well. Sodium chloride, also known as sea salt was an

additional osmotic agent for the hollow fine fibre membrane experiment. Sodium chloride is

an inorganic salt with chemical formula NaCl. It has the molecular weight of 58.44 g/mol and

density of 2.17 g/cm3.Different solutions were prepared by dissolving the osmotic agents in

deionised water and using combinations of the agents with deionised and brackish water. All

the reagents used in the study were of Laboratory Reagent Grade and the Deionised water

available at the research laboratory was used to prepare different solutions. The following

binary and ternary osmotic agent systems were selected for experiment.

Binary Osmotic Agents Systems:

a. Sucrose / H2O (Deionised water)

b. Glucose / H2O (Deionised water)

c. Sucrose / Sea water (SW)

d. Glucose / Sea water (SW)

e. NaCl / H2O (Deionised water)

Ternary Osmotic Agents Systems:

a. Sucrose + NaCl / ( DW )

b. Sucrose + NaCl / ( BW)

c. Glucose + NaCl / ( DW )

d. Glucose + NaCl / (BW )

The binary organic agent solutions (sucrose + H2O, glucose + H2O; NaCl + H2O) were tested

separately at the same osmotic pressures (10- 60 bar) but different concentrations in order to

achieve the same initial osmotic pressure. Details of the experimental conditions of different

binary osmotic solutions are shown in Table 3.6. The ternary mixtures involving sucrose +

NaCl + H2O and glucose + NaCl / H2O were tested separately at the osmotic pressures of 10,


20, 30, 40 and 50 bar. Details of the experimental conditions of different ternary osmotic

solutions are shown in Table 3.7.

Table 3.6. Experimental Conditions for FO Process: Binary Osmotic Solutions.

OA bar C Ds (g/l) Fw*

NaCl + H2O

10 12.73 DW

15 19.08 DW

20 25.38 DW

30 37.75 DW

40 49.78 DW

50 61.46 DW

Sucrose+ H2O

10 134.2 DW

15 198.29 DW

20 260.66 DW

30 380.66 DW

40 495.1 DW

50 604.65 DW

Glucose + H2O

10 71.15 DW

15 105.4 DW

20 139 DW

30 203.9 DW

40 266.5 DW

50 326.65 DW

*DW stands for deionised water

For all experiments VDs = 20 L , VFW = 20 L


Table ‎3.7. Experimental Conditions for FO Process: Ternary Osmotic Solutions

OA bar C Ds (g/l) Fw *

Sucrose +NaCl

10 134.2 BW

15 198.29 BW

20 260.66 BW

30 380.66 BW

40 495.1 BW

50 604.65 BW

Glucose +NaCl

10 71.15 BW

15 105.4 BW

20 139 BW

30 203.9 BW

40 266.5 BW

50 326.65 BW

*BW stands for brackish water

3.3.5. Experimental Measurements

The experiments for different draw solutions and feed water types were designed to maintain

a set of specific requirements. For the experiments involving brackish water, the osmotic

pressure was not allowed to exceed 50 bar, and the salinity of the feed water was not allowed

to exceed 5000 ppm using sodium chloride. The osmotic pressure of the draw solution ranged

between 10 and 60 bar. As was mentioned earlier, the osmotic agents had different

concentrations of the solute to maintain 10 bar pressure compared to the overall osmotic

pressure of the system. By doing so, the increase of osmotic pressure is given by only

sucrose, glucose or sodium chloride, respectively.

Quantitative analysis of all the samples of osmotic solutions was performed to estimate the

reverse flux of the membrane that causes diffusion of the solutes to the feed water. The

concentration of sodium chloride was analysed by Ion-Exchange Chromatography. The

analysis was performed by University of Surrey's Chemistry Department. The samples were

analysed by a Dionex IC5000 ion chromatograph after diluting 100 times with deionised

water. The quantitative analysis of sucrose and glucose was carried out by High-Performance

Liquid Chromatography (HPLC), but an interference takes placed in the sucrose analysis by


HPLC was observed in the presence of sodium chloride in the solution. First, the osmotic

pressure of each of the osmotic agent was evaluated followed by the ternary mixture (osmotic

agent and sodium chloride).

Provided below are the combinations of sucrose and sodium chloride with variable

concentrations used in the experiment for the two types of feed waters (deionised and

brackish water) and the measured osmotic pressures.

Table ‎3.8. Combinations of Sucrose and Sodium Chloride Solutions Used in The Experiment

a. Concentration of sucrose varies, deionised water is the feed

Osmotic pressure Mixture (bar) OA Mixture [g/l] Feed water

∑ bar Sucrose NaCl Sucrose NaCl DW

20 10 10 134.200 12.730 DW

25 15 10 198.290 12.730 DW

30 20 10 260.660 12.730 DW

40 30 10 380.660 12.730 DW

50 40 10 495.100 12.730 DW

60 50 10 604.650 12.730 DW

b. Concentration of sodium chloride varies, deionised water is the feed

Osmotic pressure Mixture (bar) OA Mixture [g/l] Feed Water

∑ bar Sucrose NaCl Sucrose NaCl DW

20 10 10 134.20 12.73 DW

25 10 15 134.20 19.083 DW

30 10 20 134.20 25.380 DW

40 10 30 134.20 37.750 DW

50 10 40 134.20 49.780 DW

60 10 50 134.20 61.460 DW


c. Concentration of sucrose varies, brackish water is the feed


∑ bar Sucrose NaCl Sucrose NaCl BW

20 10 10 134.200 12.730 BW

25 15 10 198.290 12.730 BW

30 20 10 260.660 12.730 BW

40 30 10 380.660 12.730 BW

50 40 10 495.100 12.730 BW

60 50 10 604.650 12.730 BW

d. Concentration of sodium chloride varies, brackish water is the feed


∑ bar Sucrose NaCl Sucrose NaCl BW

20 10 10 134.200 12.73 BW

25 10 15 134.200 19.083 BW

30 10 20 134.200 25.380 BW

40 10 30 134.200 37.750 BW

50 10 40 134.200 49.780 BW

60 10 50 134.200 61.460 BW

Figure 3.9. Osmotic Pressure as a Function of Cconcentration in sucrose-Sodium

Chloride Solution: a) Sodium Chloride concentration fixed and sucrose concentration

Variable; b) Sucrose concentration fixed and sodium chloride concentration

variable.All values obtained by OLI Analyzer software at 25°C.


Provided below are the combinations of glucose and sodium chloride with variable

concentrations used in the experiment for the two types of feed waters (deionised and

brackish water) and the measured osmotic pressures.

Table ‎3.9. Combinations of Glucose and Sodium Chloride Solutions Used in The Experiment

a. Concentration of glucose varies, deionised water is the feed


∑ bar Glucose NaCl Glucose NaCl DW

20 10 10 71.15 12.73 DW

25 15 10 105.4 12.73 DW

30 20 10 139 12.73 DW

40 30 10 203.9 12.73 DW

50 40 10 266.5 12.73 DW

60 50 10 326.65 12.73 DW

b. Concentration of sodium chloride varies, deionised water is the feed


∑ bar Glucose NaCl Glucose NaCl DW

20 10 10 71.15 12.73 DW

25 10 15 71.15 19.083 DW

30 10 20 71.15 25.380 DW

40 10 30 71.15 37.750 DW

50 10 40 71.15 49.780 DW

60 10 50 71.15 61.460 DW

c. Concentration of glucose varies, brackish water is the feed


∑ bar Glucose NaCl Glucose NaCl BW

20 10 10 71.15 12.73 BW

25 15 10 105.4 12.73 BW

30 20 10 139 12.73 BW

40 30 10 203.9 12.73 BW

50 40 10 266.5 12.73 BW


60 50 10 326.65 12.73 BW

d. Concentration of sodium chloride varies, brackish water is the feed


∑ bar Glucose NaCl Glucose NaCl BW

20 10 10 71.15 12.73 BW

25 10 15 71.15 19.083 BW

30 10 20 71.15 25.380 BW

40 10 30 71.15 37.750 BW

50 10 40 71.15 49.780 BW

60 10 50 71.15 61.460 BW

3.3.6. Synthesis of Brackish Water and Seawater

The seawater and the brackish waters were synthesized by dissolving NaCl salt in water. To

simulate seawater composition, 30 g of pure NaCl salt was dissolved in one litre of deionised

water, whereas to simulate brackish water composition, 5 g of pure NaCl salt was dissolved

in one litre of deionised water for experimental purposes (Zhang et al. 2014). Brackish water

has more salinity than fresh water, but not as much as seawater.

Figure 3.10. Osmotic pressure as a Function of Concentration in Glucose-Sodium Chloride

Solution: a) Sodium chloride concentration fixed and glucose concentration variable; b) Glucose

concentration fixed and sodium chloride concentration variable. All values obtained by OLI

Analyzer software at 25°C.


3.4. Calculations of FO Performance Parameters

The performance of FO system for different osmotic agent solutions was measured based on

the parameters of water flux (Jw), water permeability (Aw), recovery percentage (% R),

specific energy consumption (SEC) and solute flux (Jw). Additionally net driven pressure (Pn,

bar) was measured to determine the differential osmotic potential between the feed water and

the osmotic agent for its effect on the performance of the FO process.

3.4.1. Water Flux

Water flux (Jw) of feed water through the membrane was estimated by the following

equation:

Where:

QFinis the volumetric flow rate of feed water entering the membrane

QFout is the volumetric flow rate of feed water leaving the membrane

Am is the active surface area of membrane.

Note: the volume of the sugar plus water remained constant throughout the experiment, i.e., 6

litres (weight of sugar + deionised water) due to the close circulation system.

3.4.2. Water Permeability

Water permeability (Aw) through the membrane in the FO process was estimated using the

following equation.

Eqn.3.2

Where

Jw is the water flux

Pn is the net driven pressure (bar) across the membrane surface during the FO process.

Note: this equation assumes that there is no concentration polarisation as the deionised water

was used as feed water with zero salt concentration (no concentration) where its absolute

pressure is zero.


3.4.3. Recovery Percentage

Recovery percentage (% R) of the feed water was estimated using the following equation:

(

)

Where

QFin is the quantity of feed water entering the membrane

QFout is the quantity of water leaving the membrane during the FO process.

3.4.4. Salt Flux

The salt flux [g/(m2.hr]of the diffused osmotic agent was calculated by the following

equation.

Where

CFout is the concentration (g/L)of feed water leaving the membrane

QFout is the volume (L) of feed water leaving the membrane

CFin is the concentration (g/L) of feed water entering the membrane

QFin is the volume (L) of feed water entering the membrane.

3.4.5. Specific Energy Consumption

The specific energy consumption (E)[kWhr/m3]of the FO process was calculated by using the

following equation:

( )

Where:

PDin is the absolute pressure of the draw solution (osmotic agent) entering the membrane.

QDin is the volumetric flow rate (L) of draw solution (osmotic agent) entering the membrane

PFin is the absolute pressure (bar) of the feed water entering the membrane


QFinis the volumetric flow rate (L) of feed water entering the membrane and

QFoutis the volumetric flow rate(L) of feed water leaving the membrane during the FO

process.

3.4.6. Membrane Flow Rate Factor

The membrane flow rate factor (Kf) was calculated as:

Where

Aw is the water flux (L)

Am is the membrane surface area (m2).

3.4.7. Net Driving Pressure

The net driving pressure (Pn) across the membrane was estimated using the following

equation.

Where

P is the net pressure difference

∆π is the osmotic pressure difference across the membrane surface during FO process.

In this equation P was calculated by the following equation:

22

DSoutDSinFWoutFWin PPPPP

Eqn.3.8

22

DSoutDSinFWoutFWin

Eqn.3.9

As the deionised water was used as feed water and its P is equal to zero, the NDP was

calculated from the modified equation:

Eqn.3.10


4. CHAPTER FOUR: RESULTS AND DISCUSSION

4.1 Introduction

This chapter presents the results of the experiments carried out within the current study. The

ultimate purpose of the study was to identify the best combination of an osmotic agent and a

membrane type to provide the most efficient forward osmosis (FO) process performance.

Accordingly, the experiments were carried out to investigate the performance of FO process

using different osmotic agents (glucose and sucrose) and two sets of membranes (nano-

filtration and reverse osmosis flat sheet) under different process conditions (concentration,

flow rates and the temperature). The FO process performance was measured by a number of

specific parameters such as water flux, water permeability, recovery percentage, specific

energy consumption and the salt flux.

The chapter is organised in several distinctive sections. The first section presents the

experiment results for the NF membrane study. Specifically, the results are presented for the

experiments varying osmotic agents’ concentrations, feed water and osmotic agent flow rates,

and feed water and osmotic agent temperatures. The effects were measured for the parameters

of water flux (Jw), percent recovery (%R), water permeability (Aw) and specific energy

consumption (SEC). All the FO parameters were calculated using the equations given in

Chapter 3, specifically, Eqn.3.1 for water flux (Jw), Eqn3.2 for water permeability (Aw),

Eqn.3.3 for recovery percentage (% R), Eqn.3.4 for salt flux (Js), Eqn.3.5 for specific energy

consumption (SEC), Eqn.3.6 for membrane flow rate factor (kf), Eqn.3.7 for net driving

pressure ( ), Eqn.3.8 for pressure difference (∆P) and Eqn.3.9 and 3.10 for net driving

pressure ∆π.

The second section of the chapter reports the results of the same experiments for the reverse

osmosis flat sheet membranes. The third section of the chapter compares the results obtained

for the two sets of membranes. The fourth section of the chapter presents the experiments

results of osmotic agents’ binary and ternary systems. For the binary system using HFFM, the

effects of osmotic pressure differentials are reported verses deionised and seawater feeds. For

the ternary system, the effects of osmotic pressure differentials on percent recovery (%R) and

water flux (Jw) are reported. The chapter concludes with a summary of the experiment results

for all positions described in the chapter.


4.1 NF Membrane Study

The following sections present the results of the experiments using NF membrane and

sucrose and glucose as osmotic agents. In this work, performance of FO process conditions

were manipulated by changing osmotic agents’ concentrations, feed water and osmotic

agents’ flow rates and temperatures. The effects were measured for water flux, water

recovery percentage, water permeability, specific energy consumption and solute flux where

appropriate.

4.1.1‎The‎Effects‎of‎Osmotic‎Agents’‎Concentrations

The effect of osmotic agent concentrations was measured for water flux, water recovery

percent (%R), water permeability (Aw), specific energy conception (SEC) and solute flux (Js).

Different concentrations of osmotic agents at the same osmotic pressure values were used in

the FO experiment to determine their effect on the performance of FO process.

4.1.1.1. Effect on the Water Flux

The effect of osmotic agents’ concentrations on water flux is presented in Table 4.1. It can be

seen from the table that as glucose’s concentration increased from 90 to 250 g/L the water

flux decreased from 3.5 to1.5 L/(m2.hr). Correspondingly, as sucrose’s concentration

increased from170 to 400g/L, the water flux decreased from 1.2 to 0.3 L/(m2.hr). The results

of the experiment, therefore, showed overall decrease in water flux as a result of increasing

osmotic agents’ concentrations. The decrease, however, was much sharper for glucose than

for sucrose.

The apparent decrease in water flux as a result of increasing osmotic agents’ concentrations

can be explained by an increase in the osmotic agents’ potential differential and viscosity.

Data in Table 4.2 and Figure 4.1 show the effect of osmotic potential differential on water

flux with a visible decrease in water flux for both osmotic agents as a result of increasing

OPD. Specifically, water flux decreased from 0.8 to 0.3 L/(m2.hr) for sucrose based solution

and from 3.5 to 1.5 L/(m2.hr) for glucose based solution with OPD increase from 12.5 to 32

bar. Likewise, Table 4.3 shows how osmotic agents’ viscosity changed (increased) as a result

of increasing concentration, while results in Table 4.4 show a sharp decrease in water flux for

both osmotic agents as a result of increasing viscosity.


Table ‎4.1. Effect of Osmotic Agent Concentration on Water Flux, Recovery Percentage, Water Permeability, Specific Energy Consumption and

Solute Flux

Note: temperature was held constant at 25 oC, feed solution and draw solution flow rates were held constant at 2L/min

Sample calculations of the parameters for this table are presented in Appendix A10.

Table ‎4.2.Effect of Osmotic Pressure Deferential (bar) on Water Flux (Jw)

Sucrose Glucose

Jw Osmotic pressure (bar) Jw Osmotic pressure (bar)

0.8 12.5 3.5 12.5

0.4 25 3 25

0.3 32 1.5 32

Glucose

OA (g/L)

Sucrose

OA(g/L)

Osmotic

Pressure

(bar)

Osmotic

Pressure

(bar)

Water Flux

(Jw)[L/m2.hr]

Recovery Percent

(% R)

Water

Permeability (Aw)[

L/(m2.hr.bar]

Specific Energy

Consumption

(kWhr/m3)

Solute Flux

(Js)[g/(m2.hr]

Glucose Sucrose Glucose Sucrose Glucose Sucrose Glucose Sucrose Glucose Sucrose Glucose Sucrose

90 170 12.3 10.7 3.5 1.2 0.5 2.5 0.3 0.1 0.3 1.2 2.6 2.6

120 200 21.3 12.6 3.5 0.8 0.3 2 0.2 0.07 0.4 1.3 2.6 2.6

180 275 25.0 17.9 3 0.5 0.2 1.3 0.12 0.04 0.5 2.1 2.6 2.6

250 400 28.1 27.5 1.5 0.3 0.15 0.8 0.09 0.03 0.9 3.3 2.6 2.6


Table ‎4.3.Effect of Osmotic Agents’ Concentration on Viscosity

Table ‎4.4. Effect of OAs Viscosity on Water Flux

Glucose Sucrose

Jw Viscosity (cP) Jw Viscosity (cP)

3.5 1.55 1 1.89

3.5 1.66 0.8 2

3 1.88 0.5 2.6

1.5 2.22 0.3 3

Sucrose Glucose

OA Con. (g/L) Viscosity (cP) OA Con. (g/L) Viscosity (cP)

139 1.4 50 1.3

203.9 1.8 71.15 1.38

260.7 2.4 105.4 1.44

380.7 3 139 1.4

400 3.1 263.9 1.8

495.1 4.5 266.5 1.99

500 4.6 384 2.3

600 6.8 400 2.5

604.7 6.9 500 3.9

700 8.5 600 5.7


Figure ‎4.1. Effect of osmotic pressure deferential (bar) on water flux (Jw) with constant

temperature (25 oC) and constant flow rates of feed and draw solutions (2 L/min)

4.1.1.2 Effect on the Water Recovery Percent

The experimental data in Table 4.1 show the effect of osmotic agents’ concentrations on

water recovery percentage. It can be seen from the table that as glucose concentration

increased from 90 to 250 g/L, the water recovery percentage decreased from 0.5 to 0.15%.

Correspondingly, when sucrose concentration increased from 170 to 400 g/L, the water

recovery percentage decreased from 2.5% to 0.8%. The results of the experiment indicated

that water recovery percentage decreased with increasing the sucrose and glucose osmotic

agents’ concentrations versus deionised feed water. At the same time, the decrease was much

sharper for sucrose than for glucose.

4.1.1.3. Effect on the Water Permeability


water permeability (Aw). It can be seen from the table that as glucose concentration increased

from 90 to 250 g/L, the water permeability (Aw) decreased from 0.3 to 0.08 L/(m2.hr.bar).

Correspondingly, as sucrose concentration increased from 170 to 400g/L, the water

permeability decreased from 0.1 to 0.03 L/(m2.hr.bar). The results of the experiment showed

that in general an increase in osmotic agents’ concentrations leads to lower water

permeability. Because with increasing concentration, the viscosity of the solution increases

and that leads to the mobility of water molecules of dense solutions becomes less across the

0

0.5

1

1.5

2

2.5

3

3.5

4

4.5

10 12.5 15 17.5 20 22.5 25 27.5 30 32.5 35

J w (

L (

m2.h

r)

Osmotic pressure deferential (∆𝜋 in bar)

Sucrose OA

Glucose OA


semi permeable membrane .The relative low decrease for sucrose may be due to less effect of

increasing concentration on the solution viscosity as sucrose is initially more viscous than

glucose .

4.1.1.4 Effect on the Specific Energy Consumption

The experimental data in Table 4.1 show the effect of osmotic agents’ concentrations on the

specific energy consumption (SEC). It can be seen from the table that as glucose

concentration increased from 90 to 250 g/L, the SEC increased from 0.3to 0.9 kWhr/m3.

Correspondingly, as sucrose concentration increased from 170 to 400g/L, the specific energy

consumption increased from 1.2to 3.3 kWhr/m3.When the concentration of osmotic agent is

increased simultaneously, the solution becomes more viscous this led to reduce the water

flux.

4.1.1.5 Effect of OA Concentration on the Solute Flux


solute flux diffusion (Js). It can be seen from the table that as glucose concentration increased

from 90 to 250 g/L, the solute flux remained constant at 2.6 g/(m2.hr). The same results were

obtained while changing the concentration of sucrose from 170 to 400g/L. Therefore, the

results of the experiment showed that osmotic agent concentration has constant effect on

solute flux.

4.1.1.6. Discussion

The study on the effect of osmotic agent concentration showed that parameters such as Jw, %

R, Aw decreased, while the SEC increased and the Js remained constant. As explained earlier,

increasing concentration of OAs increases the osmotic potential (driving force: ∆π) on the

active layer side of the membrane surface thus resulting in a decreasing trend of all these

parameters except SEC and Js when the concentration of OAs was increased. These effects

may be attributed to osmotic pressure differential across the membrane surface as a result of

the concentration polarisation and the NF membrane properties such as pore space and

structure.

As was discussed in Section 2.4.5 of this thesis, when the solute particles accumulate on the

thin layer near the membrane surface, the concentration polarisation occurs which affects the


different FO parameters (McCutcheon and Elimelech 2006; Sablani et al. 2001; Zhao and

Zou 2011). Due to membrane processes such as diffusion, osmotic pressure potential,

movement of water molecules across the membrane surface, the concentration polarisation is

one of the major factors limiting the reduction of permeate flux (McCutcheon and Elimelech

2006; Zhao and Zou 2001). In the experiment process applied for this study, concentration

polarisation was the outcome of the concentration difference between the feed and the draw

solutions. NF membrane pore space and structure also contributed to osmotic pressure

differential. NF membranes are often considered “loose membranes” because of relatively

large pore size that increase water permeability but also have lower recovery percentage

(Pabby et al., 2015).

4.1.2 The Effects of Feed Water Flow Rate

The effect of feed water flow on water flux, water recovery percentage, water permeability,

specific energy consumption and solute flux were measured. The results are summarized in

Table 4.5.


The experimental data in Table 4.5 show the effect of feed water flow rate on the water flux.

It can be seen from the table that as deionised feed water flow rate was increased from 2 to 4

L/min for sucrose and glucose osmotic agents, the water flux increased from 0.4 to 1.4

(L/m2.hr) and from 0.1 to 0.5 L/(m

2.hr) , respectively. Therefore, the results of the experiment

showed that increase in feed water flow rate leads to an increase in the water flux for both

types of osmotic agents. For sucrose, however, the initial flow rate and the following water

flux rate increase were higher.


The experimental data in Table 4.5 show the effect of feed water flow rate on the water

recovery percent. It can be seen from the table that as deionised feed water flow rate was

increased from 2 to 4 L/min for sucrose and glucose osmotic agents, the water recovery

percentage increased from 0.1% to 1.0 % when using glucose as an osmotic agent and

increased from 0.01% to 0.05 % when using sucrose as an osmotic agent. The results of the

experiment showed in general higher water flow rates lead to an increased water recovery

percentage when using both osmotic agents.


4.1.2.3 Effect on the Water Permeability


permeability. It can be seen from the table that as deionised feed water flow rate was

increased from 2 to 4 L/min for sucrose and glucose osmotic agents, the water permeability

increased from 0.005 to 0.02 L/(m2.hr.bar) when glucose was used as an osmotic agent, and

increased from 0.01 to 0.05 L/(m2.hr.bar) when sucrose was used as an osmotic agent. The

results of the experiment showed in general higher water flow rates lead to an increased water

permeability when using both osmotic agents. This difference in the water permeability may

be attributed to the increasing osmotic pressure of sucrose than glucose which caused high

water permeability for sucrose with increasing the flow rate than glucose osmotic agent.


The experimental data in Table 4.5 show the effect of feed water flow rate on the specific

energy consumption. It can be seen from the table that as deionised feed water flow rate was

increased from 2 to 4 L/min for sucrose and glucose osmotic agents, the specific energy

consumption decreased from 12 to 1.8 kW.hr/m3

when sucrose was used as an osmotic agent

and decreased from 2.2 to 0.6 kW.hr/m3 when glucose was used as an osmotic agent. The

results of the experiment showed that in general an increase in the feed water flow rate leads

to the lower degree of the specific energy consumption. There was a very sharp decrease

when sucrose was used as an osmotic agent, and a moderate decrease when glucose was used

as an osmotic agent.


Table ‎4.5.Effect of Feed Water Flow Rate on Water Flux, Recovery Percentage, Water Permeability, Specific Energy Consumption and Solute

Flux

Note: the results were obtained at constant temperature (25 oC) and osmotic agent concentration (275and 200 g/L) for sucrose and glucose

respectively .

Flow

Rate

(L/min)

Water Flux

(Jw)[L/(m2.hr]

Recovery Percent

(% R)

Water Permeability

(Aw)[L/(m2.hr.bar]

Specific Energy

Consumption

(kWhr/m3)

Solute Flux

(Js)[g/(m2.hr)]

Glucose Sucrose Glucose Sucrose Glucose Sucrose Glucose Sucrose Glucose Sucrose

2 0.1 0.4 0.1 0.01 0.005 0.01 2.2 12 2.6 2.6

2.5 0.15 0.6 0.15 0.01 0.002 0.01 1.5 8 2.6 2.6

3 0.2 0.8 0.2 0.02 0.003 0.02 1 5 2.6 2.6

3.5 0.4 1.1 0.4 0.03 0.01 0.03 0.64 3 2.6 2.6

4 0.5 1.4 1.0 0.05 0.02 0.05 0.6 1.8 2.6 2.6


4.1.2.3. Effect on the Solute Flux

The experimental data in Table 4.5 show the effect of feed water flow rate on the solute flux

(Js). It can be seen from the table that as deionised feed water flow rate was increased from 2

to 4L/min for sucrose and glucose osmotic agents, the solute flux remained the same at 2.6

g/(m2.hr). Therefore, the results of the study showed that feed water flow rate had constant

effect on solute flux for either of the osmotic agents with the NF membrane.

4.1.2.4. Discussion

It can be seen from the feed water flow rate study that the values of all the FO parameters

namely Jw, % R, Aw increased, while the SEC decreased when the feed water flow rate

increased from 2 to 4 L/min, while the Js remained constant. Similar conclusions were

reported by McCutcheon (2006) who stated that concentrative ECP on the feed side of the

membrane has a minor effect on the driving force (∆π) unless the feed concentration or the

permeate flux is relatively high. He also mentioned that increasing the feed water flow rate

decreases the concentrative ECP effect on the feed side of the membrane which will allow

more water flow through the membrane pores coupled with increasing the osmotic pressure

of the feed water near the membrane surface.

4.1.3. The‎Effects‎of‎Osmotic‎Agents’‎Flow‎rate

The effect of osmotic agents’ flow rate on water flux, water permeability and solute flux was

measured when using sucrose and glucose osmotic agents versus deionised water. Table 4.6

summarises the outcomes of the experiments.


The experimental data in Table 4.6 show the effect of osmotic agent’s flow rate on the water

flux. It can be seen from the table that as the OA’s flow rate increased from 2 to 4 L/min, the

water flux rate increased from 0.2 to 0.6 L/(m2.hr) when sucrose was used as an osmotic

agent and increased from 0.4 to 0.9 L/(m2.hr)when glucose was used as an osmotic agent.

This may be due to high ratio of flow rate on draw solution osmotic potential difference

across the membrane surface with increasing the flow rate. This may be due to the ratio of

feed water to draw solution tended to rinse the experimental setup.


4.1.3.2. Effect of Recovery Percentage

The data in Table 4.6 illustrate the effect of osmotic agent flow rate on recovery percentage.

It can be seen from the table that as the OA flow rate increased from 2 to 4 L/min, the

recovery percentage increased from 0.2 to 0.9 % when sucrose was used as osmotic agent and

it increased from 0.1 to 0.5 % when glucose was used as an osmotic agent. It appears from

the experimental results that generally, the recovery percentage showed increasing trend with

increasing the osmotic agent’s flow rate. However, the initial and the final recovery

percentage were higher when sucrose was used as an osmotic agent.



flux. It can be seen from the table that as the OA flow rate increased from 2 to 4 L/min, the

water permeability rate increased from 0.07 to 0.3 L/(m2.hr.bar) when sucrose was used as an

osmotic agent and increased from 0.1 to 0.35 L/(m2.hr.bar)when glucose was used as an

osmotic agent. The results of the experiment showed that in general, an increase of the

osmotic agents’ flow leads to a higher water permeability rate. The initial and the final water

permeability rates were higher when glucose was used as an osmotic agent.


Table ‎4.6. Effect of Osmotic Agent Flow Rate on Water Flux, Recovery Percentage, Water permeability, Specific Energy Consumption and

Solute Flux

Note: the results were obtained at constant temperature (25 oC) and osmotic agent concentration (275 and 200 g/L) for both sucrose and glucose

respectively

Flow

Rate

(L/min)

Water Flux

(Jw)[L/(m2.hr)]

Recovery Percent

(% R)

Water

Permeability

(Aw)[L/(m2.hr.bar)]

Specific Energy

Consumption

(kWhr/m3)

Solute Flux

(Js)[g/(m2.hr)]


2 0.4 0.2 0.1 0.2 0.1 0.07 3.17 7.8 2.6 2.6

2.5 0.45 0.25 0.15 0.3 0.15 0.11 2 4.45 2.6 2.6

3 0.55 0.35 0.2 0.45 0.2 0.15 1.6 3 2.6 2.6

3.5 0.65 0.45 0.32 0.66 0.25 0.2 1.2 2 2.6 2.6

4 0.9 0.6 0.5 0.9 0.35 0.3 1.04 1.5 2.6 2.6


4.1.3.4. Effect on Osmotic Agent Flow Rate on Specific Energy Consumption (SEC)

The results in Table 4.6 show the effect of osmotic agent flow rate on specific energy

consumption. It is evident from Table 4.6 that the SEC decreased from 7.8 to 1.5 kW.hr/m3

when the sucrose osmotic agent flow rated increased from 2 to 4 L/min, the SEC decreased

from 3.17 to 1.04 kW.hr/m3 when the glucose osmotic agent flow rate increased from 2 to 4

L/min. Overall, it was observed that the initial and the final SEC decrease was more when

using sucrose osmotic agent compared to glucose osmotic agent with corresponding increase

in osmotic agent flow rate from 2 to 4 L/min.

4.1.3.5 Effect on the Solute Flux

The experimental data in Table 4.6 show the effect of osmotic agent’s flow rate on the solute

flux. It can be seen from the table that as the osmotic agent’s flow rate increased from 2 to 4

L/min, the solute flux remained constant at 2.6 g/(m2.hr) Therefore, the results of the study

indicated that osmotic agent’s flow rate had constant effect on solute flux for either of the

osmotic agents with the NF membrane. This could be attributed to the type of membrane used

in this study and the nature of solute molecules of both the osmotic agents.

4.1.3.6. Discussion

It was observed from the OAs flow rate study that the values of FO parameters such as Jw,

Aw, % R increased, while the SEC decreased and Jw remained constant. The constant increase

in Jw, Aw and % R is due to the increased osmotic pressure differential (OPD) across the NF

membrane surface with increasing OAs flow rate. Correspondingly the SEC requirement

decreased for efficient FO performance. According to Cath et al. (2006), the negative effect

of ECP on flux can be decreased by either optimizing the flow rate or by increasing the

overall flow velocity. They also stated that ICP is similar to ECP process, with the only

difference that it takes place within the membrane porous layer, and is considered as one of

the major issues in osmotically-driven membrane processes.

4.1.4 The Effects of Feed Water Temperature

The effect of feed water temperature on water flux, water recovery percent, water

permeability, specific energy consumption and solute flux were measured using sucrose and

glucose osmotic agent versus deionised water in NF membrane. The results of the

experiments are demonstrated in Table 4.7.


4.1.4.1. Effect on Water Flux

The experimental data in Table 4.7 show the effect of feed water temperature on water flux. It

can be seen from the table that as the temperature increased from 20 to 35 oC, the water flux

increased from 0.2 to 0.8L/(m2.hr)when sucrose was used as an osmotic agent and increased

from 0.5 to 1.1L/(m2.hr)when glucose was used as an osmotic agent. Therefore, the results of

the experiment indicated that in general an increase in feed water temperature leads to an

increase in water flux for both types of osmotic agents. However, both the initial and the final

water flux rates were higher for the processes involving glucose as an osmotic agent.

4.1.4.2. Effect on the Water recovery percentage

The experimental data in Table 4.7 show the effect of feed water temperature on water

recovery percentage. It can be seen from the table that as the temperature increased from 20

to 35 oC, the water recovery percentage increased from 0.5 % to 2.7 % when sucrose was

used as an osmotic agent and increased from 0.8 % to 4.0 % when glucose was used as an

osmotic agent. Therefore, the results of the experiment showed that in general an increase in

feed water temperature leads to an increase in water recovery percentage for both types of

osmotic agents. Because as the temperature of the solution increases, the speed and

movement of water molecules as well as the diffusion rate increases across the membrane

surface thus resulting in high water recovery percentage compared to low temperature

solutions in FO process. Also, the effect of temperature is more on glucose than sucrose due

to the difference in viscosity between the two osmotic agent solutions.


Table ‎4.7. Effect of Feed Water Temperature on Water Flux, Recovery Percentage, Water permeability, Specific Energy Consumption and Solute

Flux

Note: the results were obtained at constant feed and draw solution flow rates (2 L/min) and osmotic agent concentrations (200 and 250 g/L) for

both sucrose and glucose respectively

Temperature

(o C)

Water Flux

(Jw)[L/(m2.hr)]

Recovery Percent

(% R)

Water Permeability

(Aw)[L/(m2.hr.bar)]

Specific Energy

Consumption

(kW.hr/m3)

Solute Flux

(Js)[g/(m2.hr)]


20 0.5 0.2 0.8 0.5 0.02 0.08 8.0 6.5 2.6 2.6

25 0.6 0.25 1.5 0.7 0.15 0.18 5 3.5 2.6 2.6

30 0.8 0.4 2.5 1 0.25 0.28 2.5 2 2.6 2.6

35 1.1 0.8 4.0 2.7 0.4 0.6 1.5 1.1 2.6 2.6




permeability. It can be seen from the table that as the temperature increased from 20 to 35 oC,

the water permeability increased from 0.08 to 0.6 L/(m2.hr.bar) when sucrose was used as an

osmotic agent and increased from 0.02 to 0.4 L/(m2.hr.bar) when glucose was used as an


feed water temperature leads to an increase in water permeability for both types of osmotic

agents. Similar to the effect of temperature on water recovery percentage, the water

permeability showed increases with increasing the temperature mainly due to its effect on

speed and movement of water molecules, diffusion rate and the potential difference across the

membrane surface resulting in higher water permeability of glucose osmotic agent.


The experimental data in Table 4.7 show the effect of feed water temperature on the specific

energy consumption. It can be seen from the table that as the temperature increased from 20

to 35 oC, the energy consumption decreased from 6.5 to 1.1 kWhr/m

3 when sucrose was used

as an osmotic agent and decreased 8.0to1.5 kWhr/m3 when glucose was used as an osmotic

agent. Therefore, the results of the experiment showed that in general an increase in feed

water temperature leads to a decrease in specific energy consumption for both types of

osmotic agents. Since the increase of temperature mainly effects the rate of diffusion ,speed

and movement of water molecules and osmotic potential difference across the membrane

surface. Therefore, less energy is required for efficient performance of FO process using

glucose osmotic agent than sucrose.


The experimental data in Table 4.7 show the effect of feed water temperature on the solute

flux. It can be seen from the table that as the temperature increased from 20 to 35 oC, the

solute flux remained constant at 2.6 g/(m2.hr) for both types of osmotic agents. Therefore, the

results of the experiment showed that there was constant effect of feed water temperature on

the solute flux.


4.1.4.6. Discussion

It was found from the feed water temperature study that water flux, water permeability and

recovery percentage increased, the specific energy consumption decreased and the solute flux

was not affected with increasing the FO process temperature from 20 to 35oC. Higher

temperatures in osmotic processes have been found to positively affect the major’s market

parameters. Phuntsho et al. (2012) reported an average increase in the water flux by up to 1.2

% for each degree rise in temperature from 25to 35oC, but this increase was around 2.3 %

with increasing the temperature from 25to 45oC. In another study, You et al. (2012) found a

positive correlation between water flux and the bulk solution temperature ranging between

20-40oC. The positive effect of feed water temperature on osmotic parameters is often

explained by increases in the diffusion coefficients resulting from decreasing viscosity of the

solutions (Nayak and Rastogi 2011; Petrotos et al. 1998). This leads to higher permeability of

water through the membrane. However, higher diffusion rates also apply for salts in this case,

which leads to lower salt rejection rate. Furthermore, it is extremely important that the

membrane can tolerate higher temperatures (Pabby et al. 2015).

4.1.5 The Effects of Osmotic‎Agents’‎Temperatures

The effect of osmotic agent temperature was measured for water flux, water recovery percent,

water permeability, specific energy consumption and solute flux. The results of the

experiments are presented in Table 4.8.


The experimental data in Table 4.8 show the effect of osmotic agents’ temperatures on the

water flux. It can be seen from the table that as the temperature increased from 20 to 35 oC,

the water flux increased from 0.1 to 0.15 L/(m2.hr)when sucrose was used as an osmotic

agent and increased from 0.11 to 0.17 L/(m2.hr)when glucose was used as an osmotic agent.

Therefore, the results of the experiment showed that in general an increase in osmotic agents’

temperatures leads to an increase in water flux for both types of osmotic agents. However,

both the initial and the final permeability rates were higher for the processes involving

glucose as an osmotic agent.


Table ‎4.8. Effect of OA Temperature on Water Flux, Recovery Percentage, Water Permeability, Specific Energy Consumption and Solute Flux

Note: the results were obtained at constant feed and draw solution flow rates (2 L/min) and osmotic agent concentration (200 and 250 g/L) for


Temperature

(o C)

Osmotic

Pressure

(bar)

Osmotic

Pressure

(bar)

Water Flux

(Jw)[L/(m2.hr)]

Recovery Percent

(% R)

Water

Permeability

(Aw)[L/(m2.hrbar)]

Specific Energy

Consumption

(kWhr/m3)

Solute Flux

(Js)[g/(m2.hr)]


20 12.3 10.7 0.11 0.1 0.2 0.4 0.02 0.01 1.5 1.1 2.6 2.6

25 21.3 12.6 0.13 0.12 0.6 0.8 0.035 0.02 1.43 0.9 2.6 2.6

30 25.0 17.9 0.15 0.135 1 1.3 0.05 0.025 1.35 0.75 2.6 2.6

35 28.1 27.5 0.17 0.15 1.7 2 0.09 0.03 1.2 0.6 2.6 2.6



The experimental data in Table 4.8 show the effect of temperature of the draw solution on the


the water recovery percent increased from 0.4% to 2 % when sucrose was used as an osmotic

agent and increased from 0.2% to1.7% when glucose was used as an osmotic agent.

Therefore, the results of the experiment showed that in general increase in the temperature of

the draw solution leads to an increase in the water recovery percent. This may be attributed to

the effect of increasing temperature on the physical properties of the various osmotic agents

under study .The main effect of increasing temperature is on the viscosity, speed and

movement of molecules, rate of diffusion and osmotic potential difference across the

membrane surface in FO process.


Data in Table 4.8 show the effect of osmotic agent temperature on water permeability. It can

be seen that water permeability increased from 0.02 to 0.09L/(m2.hr.bar) when glucose

osmotic agent was used and increased from 0.01 to 0.03 L/(m2.hr.bar) when sucrose osmotic

agent was used. In both experiments the temperature increased up to 35 oC. Overall, it

appears from the results that increase in water permeability was higher using glucose osmotic

agent compared to sucrose osmotic agent due to the difference in viscosity.


The experimental data in Table 4.8 show the effect of osmotic agents’ temperatures on the

specific energy consumption. It can be seen from the graph that as the temperature increased

from 20 to35oC, the specific energy consumption rate declined from 1.1 to 0.6 kWhr/m

3when

sucrose was used as an osmotic agent and decreased from 1.5 to 1.2 kWhr/m3when glucose

was used as an osmotic agent. Therefore, the results of the experiment showed that in general

increasing the temperature of the draw solution leads to a decrease in special energy

consumption. This may due to the fact that the temperature affects the solution properties

namely speed, movement and diffusion rate of water molecules along with osmotic potential

difference across the membrane surface thus requiring less energy for performance of FO

process compared to low temperature solutions. Since the viscosity of sucrose is higher than

glucose, that is way the effect is clearer for sucrose than glucose agent.



The experimental data in Table 4.8 show the effect of temperature of the draw solution on the

solute flux. It can be seen from the graph that as the temperature increased from 20to 35 oC, a

constant solute flux of 2.6 g/(m2.hr) remained when each of the osmotic agents (glucose and

sucrose) was used. Therefore, the experiment results showed constant effect of temperature of

the draw solution on the solute flux.

4.1.5.6. Discussion

It was observed from the results of the study of temperature of the draw solution that FO

parameters such as water flux, recovery percentage and water permeability increased, the

SEC decreased, while the solute flux remained constant . Increasing the temperature of the

SEC deceased, while the temperature of the draw solution increased from 20 to 35 oC. That

the temperature of the draw solutions could substantially affect the outcomes of the emerging

FO was demonstrated elsewhere (Arsuaga et al. 2008; Sharma and Chellam 2006; Xie et al.

2013; Zhao and Zou 2012). The general explanation is that an increase temperature of draw

solutions leads to a decrease in viscosity rates. Draw solutions with lower viscosity

demonstrate higher-level of the solubility and diffusivity coefficients (Xie et al. 2013).

Accordingly, the outcomes of these study shown in Figure 4.24 indicated that the viscosity of

both the glucose and sucrose decline with rising the temperature in FO process. Therefore,

the flow of the experiment and its outcomes confirmed the previously reported findings.

4.1.6. Overview of the NF Membrane System Results

Several conclusions are drawn from this experiment for the NF membrane (Table 4.9). First,

it can be concluded that for the given osmotic agents (sucrose and glucose) and the given

membrane (NF) changes applied for feed water had better effect on the process outcomes.

With the changes to feed water flow and temperature, the experiments in both cases

demonstrated increases in water flux, water recovery rates and water permeability while also

leading to lower specific energy consumption levels. Changes applied to draw solutions

resulted in worsening of at least one measured parameter. Specifically, changes in draw

solution flow rate and temperature resulted in higher energy consumption, while changes to

draw solution concentration levels resulted in decline of all relevant parameters. Overall, it

can be concluded that manipulation of the forward osmosis process for the NF filtration


membrane can improve effectiveness of the FO process across all relevant parameters. The

only factor that is not desirable for increase is the draw solution concentration.

Another important conclusion is that neither glucose nor sucrose showed greater comparative

efficacy as an osmotic agent for the NF membrane. Alhemiri et al. (2009) recommended that

the best osmotic agent for FO should be based on the water flux parameter. Glucose showed

better flux in comparison to sucrose for NF membrane in all cases except for increase in feed

water flow. At the same time, glucose-based draw solutions did not show superiority over

sucrose-based solutions across all measured parameters when the relevant factors were

manipulated during the experiment. Therefore, from the results of this research, it can be

concluded that neither sucrose nor glucose are effective as osmotic agents for NF membranes.

Table ‎4.9. Results of Changes of Manipulated FO Factors on Key Measured Parameters of the

Process: NF Membrane

Parameter

Change

Water Flux Water Recovery

Rate

Water

Permeability

Specific Energy

Consumption

Solute Flux

Increase in DS

concentration

Decreased

Glucose superior

Decreased

Sucrose superior

Decreased

Glucose superior

Increased

Glucose superior

No change

Increase in FW

flow rate

Increased

Sucrose superior

Increased

Glucose superior

Increased

Sucrose superior

Decreased

Glucose superior

No change

Increase in DS

flow rate

Increased

Glucose superior

Increased

Sucrose superior

Increased

Glucose superior

Decreased

Glucose superior

No change

Increase in FW

temperature

Increased

Glucose superior

Increased

Sucrose superior

Increased

Sucrose superior

Decreased

Sucrose superior

No change

Increase in DS

temperature

Increased

Glucose superior

Increased

Sucrose superior

Increased

Glucose superior

Decreased

Sucrose superior

No change

4.2. RO Flat Sheet Membrane Study

The following sections present the results of the experiments when using RO membrane and

sucrose and glucose as osmotic agents verses deionised water. The FO process conditions

were manipulated by changing osmotic agents’ concentrations, feed water and osmotic

agents’ flow rates and temperatures. The effects were measured for water flux, percent

recovery, water permeability, specific energy consumption and solute flux where appropriate.

4.2.1.‎The‎Effects‎of‎Osmotic‎Agents’‎Concentrations

The effect of osmotic agent concentrations was measured for water flux (Jw), water recovery

percent (%R), water permeability (Aw), specific energy conception (SEC) and solute flux (Js).

Different concentrations for the osmotic agents were used to obtain the same osmotic


pressure in the FO experiment to determine their effect on the performance of FO process.

The results of the experiment are provided in Table 4.10.


The effect of osmotic agents’ concentrations on water flux is Table 4.10. It can be seen from

the table that as glucose’s concentration increased from 90 to 250 g/L the water flux

decreased from 2.0 to 0.3L/(m2.hr). Correspondingly, as sucrose’s concentration increased

from 170 to 400g/L, the water flux decreased from 3.7 to 0.9 L/(m2.hr). The results of the

experiment, therefore, showed overall decrease in water flux as a result of increasing osmotic

agents’ concentrations. The processes with both osmotic agents showed a decreasing trend in

water flux as osmotic concentrations increased. Because with increasing the concentration,

the osmotic potential difference between the active layer and the support layer increases

across the membrane surface thus resulting in low water flux. Similarly in Table 4.11, the

water flux showed a decreasing trend as increase in osmotic potential difference (bar) across

the membrane surface. Since RO membrane is a compact membrane, the effect of increasing

concentration is clearer than using NF membrane being a loose membrane.


Table ‎4.10. Effect of Osmotic Agent Concentration on Water Flux, Recovery Percentage, Water permeability, Specific Energy Consumption and

Solute Flux

Note: the results were obtained at constant temperature (25 oC) and feed and draw solution flow rates (2 L/min)

Table ‎4.11. Effect of Osmotic Pressure Differential on Water Flux

Osmotic Pressure (bar) Sucrose Glucose

Water Flux (Jw)

12.9 3.7 2

15.5 3 1.5

25.3 1.8 0.5

31.6 0.9 0.3

Glucose

OA (g/L)

Sucrose

OA(g/L)

Osmotic

Pressure

(bar)

Osmotic

Pressure

(bar)

Water Flux

(Jw)[L/(m2.hr)]

Recovery Percent

(% R)

Water Permeability

(Aw)[L/(m2.hr.bar)]

Specific Energy

Consumption

(kWhr/m3)

Solute Flux

(Js)[g/m2.hr)]


90 170 12.3 10.7 2 3.7 0.5 2 0.1 0.7 1 2.9 2.6 2.6

120 200 21.3 12.6 1.5 3 0.4 1.7 0.096 0.4 0.98 3 2.6 2.6

180 275 25.0 17.9 0.5 2 0.265 1 0.06 0.2 1.3 4 2.6 2.6

250 400 28.1 27.5 0.3 0.9 0.13 0.3 0.04 0.1 2 5.6 2.6 2.6


4.2.1.2. Effect on the Percent of Water Recovery


water recovery percentage. It can be seen from the table that as glucose concentration

increased from 90 to 250 g/L, the water recovery percentage (% R) decreased from 0.5% to

0.13 %. Correspondingly, when sucrose concentration increased from 170 to 400 g/L, the

water recovery percentage decreased from 2.0% to 0.3 % when glucose concentration

increased from 90 to 250 g/L. The results of the experiment indicated that water recovery

percentage decreased more when sucrose osmotic agent was used compared to glucose

osmotic agent.



water permeability (Aw). It can be seen from the table that as glucose concentration increased

from 90 to 250 g/L, the water permeability (Aw) decreased from 0.1 to 0.04L/(m2.hr.bar).

Correspondingly, as sucrose concentration increased from 170 to 400g/L, the water

permeability decreased from 0.7 to 0.1 L/(m2.hr.bar).The results of the experiment showed

that in general an increase in osmotic agents’ concentrations leads to lower water

permeability. For the processes involving sucrose as an osmotic agent, however, the decrease

was much sharper.

4.2.1.4. Effect on the Specific Energy Consumption


specific energy consumption (SEC). It can be seen from the table that as glucose

concentration increased from 90 to 250 g/L, the specific energy consumption SEC increased

from 1.0 to 2.0 kWhr/m3. Correspondingly, as sucrose concentration increased from 170 to

400g/L, the specific energy consumption increased from 2.9 to 5.6 kWhr/m3. In general, the

results of the experiment showed that sucrose and glucose as osmotic agents provided

different effects on the specific energy consumption. This may be due to the difference in the

viscosity of the osmotic agents. With the higher viscosity, sucrose requires higher energy

levels to maintain an efficient FO process performance.




solute flux (Js). It can be seen from the table that as glucose concentration increased from 90

to 250 g/L, the solute flux remained constant at 2.6 g/(m2.hr).The same results were obtained

while changing the concentration of sucrose from 170 to 400g/L. Therefore, the results of the

experiment showed that both osmotic agents concentration has constant effect on solute flux.

4.2.1.6. Discussion

The data in Table 4.10 indicate a decline in water flux, recovery percentage and water

permeability, an increase in specific energy consumption, but no response on solute flux

when the OAs concentration was increased from 70 to 250 g/L for glucose and 170 to 400

g/L for sucrose. As was in case with the NF membrane discussed above, increasing the OAs

concentration increases the osmotic pressure differential affecting the ICP and ECP in the

vicinity of the membrane surface which resulted in a decrease in some parameters except an

increase in SEC requirement for efficient FO performance. However, it is interesting to note

that solute flux was not affected, but remained constant with increasing the concentration for

both OAs (glucose and sucrose). Also, sensitivity of sugars (sucrose and glucose) is not

affected by the OAs concentration, but is mainly related to flux, and the set-up which

produces the highest flux, generally generates the lowest draw solution sensitivity. The study

findings agree with those of McCutcheon et al. (2005) who reported that FO is promising to

achieve high rate water fluxes and water recovery with increasing OAs concentration thus

significantly reducing the amount of rejected brine and the required post-treatment facilities.

They further indicated a significant improvement in relation to RO in terms of being

environment friendly. Sucrose based OA in this experiment showed better performance in

terms of water flux, recovery and permeability; however, it also consumed higher amounts of

energy.

4.2.2. The Effects of Feed Water Flow rate

The effect of feed water flow rate was measured for water flux (Jw), water recovery percent

(%R), water permeability (Aw), specific energy conception (SEC) and solute flux (Js). The

results of the experiments are demonstrated in Table 4.12.


Table ‎4.12. Effect of Feed Water Flow Rate on Water Flux, Recovery Percentage, Water Permeability, Specific Energy Consumption and Solute

Flux

Note: The results were obtained at constant temperature (25 oC) and osmotic agent concentration (275 and 200 g/L) for both sucrose and glucose

respectively

Flow

Rate

(L/min)

Water Flux

(Jw)[L/(m2.hr)]

Recovery Percent

(% R)

Water Permeability

(Aw)[L/(m2.hr.bar)]

Specific Energy

Consumption

(kWhr/m3)

Solute Flux

(Js)[g/(m2.hr)]


2 0.09 0.89 0.5 4.3 0.02 0.1 3.1 0.7 2.6 2.6

2.5 0.25 1.3 0.7 5.5 0.06 0.25 2.39 0.85 2.6 2.6

3 0.5 1.5 1 6.5 0.08 0.34 1.69 1 2.6 2.6

3.5 0.7 1.7 1.17 7 0.1 0.43 1.1 1.5 2.6 2.6

4 0.9 1.8 1.3 7.5 0.2 0.5 1.0 2 2.6 2.6



The experimental data in Table 4.12 show the effect of feed water flow rate on the water flux.

It can be seen from the table that as deionised feed water flow rate was increased from 2 to 4

L/min for sucrose and glucose osmotic agents, the water flux increased from 0.89 to 1.8

L/(m2.hr) and from 0.09 to 0.9 L/(m

2.hr) respectively. Therefore, the results of the experiment

showed that increase in feed water flow rate leads to an increase in the water flux for both

types of osmotic agents. For sucrose, however, the initial flow rate and the following water

flux rate increase were higher. The range of water flux increase was also higher for the

processes where sucrose was used as an osmotic agent.

4.2.2.2. Effect on the Water Recovery Percentage


recovery percent. It can be seen from the table that as deionised feed water flow rate was

increased from 2 to 4 L/min for sucrose and glucose osmotic agents, the water recovery

percentage (%R) increased from 4.3 % to 7.5% when using sucrose as an osmotic agent and

increased from 0.5% to 1.3% when using glucose as an osmotic agent. The results of the

experiment showed that in general higher water flow rates lead to an increased water

recovery when using both osmotic agents. However, the initial and the final recovery rates for

sucrose were higher than for glucose at the same osmotic pressures. This could be due to the

difference in viscosity between these two solutions at the same concentration.



permeability. It can be seen from the table that as deionised feed water flow rate was

increased from 2 to 4 L/min for sucrose and glucose osmotic agents, the water permeability

increased from 0.02 to 0.2 L/(m2.hr.bar) when glucose was used as an osmotic agent, and

increased from 0.1 to 0.5 L/(m2.hr.bar) when sucrose was used as an osmotic agent. The

results of the experiment showed in general higher water flow rates lead to an increased water

permeability when using both osmotic agents. However, the initial and the final permeability

rates for sucrose were higher than for glucose at the same osmotic pressures. The processes

where sucrose was used as an osmotic agent also showed a steeper increase in the water

permeability rate.



The experimental data in Table 4.12 show the effect feed water flow rate on the specific

energy consumption. It can be seen from the table that as deionised feed water flow rate was

increased from 2 to 4 L/min for sucrose and glucose osmotic agents, the specific energy

consumption decreased from 3.1 to 1.0 kWhr/m3when glucose was used as an osmotic agent

and slightly increased from 0.7 to 2 kWhr/m3when sucrose was used as an osmotic agent.

This slight increase in SEC for sucrose osmotic agent may be due to its higher viscosity as

the movement of dense and viscous solution is very slow through the membrane.


The experimental data in Table 4.12 show the effect of feed water flow rate on the solute flux

(Js). It can be seen from the table that as deionised feed water flow rate was increased from 2

to 4 L/min for sucrose and glucose osmotic agents, the solute flux remained the same at 2.6

g/(m2.hr).Therefore, the results of the study showed that feed water flow had constant effect

on solute flux for either of the osmotic agents with the RO flat sheet membrane.

4.2.2.6. Discussion

The data in Table 4.12 show that FO parameters such as water flux, recovery percentage and

water permeability increased, while specific energy consumption decreased, but solute flux

was not affected when the feed water flow rate was increased from 2 L/min to 4 L/min for

both OAs (glucose and sucrose). Because when the feed water flow rate increases, it

increases the osmotic pressure due to high water volume at the membrane surface thus

affecting the ECP and ICP at the membrane surface resulting a substantial high osmotic

potential difference (OPD) between the active layer and the support layer of the membrane.

This phenomenon caused increases in some parameters and reduces the SEC requirements.

At the same time, increasing flux reduces the salt sensitivity in FO process. Recently Zhao et

al. (2012) reported high water flux is the function of water permeability (Aw), membrane

structure, solute resistivity and the solute permeability. Also it is important to consider ICP of

the test membrane for higher efficiency.


4.2.3.‎The‎Effects‎of‎Osmotic‎Agents’‎Flow‎rate

The effect of osmotic agents’ flow was measured for water flux (Jw), water permeability

(Aw), water recovery percent (%R) and specific energy consumption (SEC). The results of the

experiment are demonstrated in Table 4.13.


The experimental data in Table 4.13 show the effect of osmotic agents’ flow rate on the water

flux (Jw). It can be seen from the table that as the osmotic agent’s flow rate increased from 2

to 4 L/min, the water flux decreased from 1.0 to 0.4 L/(m2hr) when sucrose was used as an

osmotic agent and decreased from 1.30 to 0.55 L/(m2.hr) when glucose was used as an

osmotic agent. Therefore, the results of the experiment showed that in general increase in the

draw solution flow rate leads to a decrease in the water flux. At the same time, the initial

water flux and the final water flux were higher when glucose was used as an osmotic agent.

The eventual decline in water flux was moderate for both osmotic agents. The decrease in Jw

for both osmotic agents might have been due to an increase in the concentration of the

osmotic agents which has led to development of higher osmotic pressure on the feed side of

the membrane than on the support layer of the membrane surface.



permeability (Aw). It can be seen from the table that as the draw solution’s flow rate increased

from 2 to 4 L/min, the water permeability rate decreased from 0.35 to 0.066L/(m2.hr.bar)

when sucrose was used as an osmotic agent and decreased from 0.26 to 0.04L/(hm2.hr.bar)

when glucose was used as an osmotic agent. The results of the experiment showed that

osmotic agents exhibited different behaviour as a result of an increase in osmotic agents’ flow

rate. The initial water permeability rate was higher when sucrose was used as an osmotic

agent. In the end, however, the difference between the processes in terms of water

permeability was minimal.


Table ‎4.13. Effect of Osmotic Agent Flow Rate On Water Flux, Recovery Percentage, Water Permeability, Specific Energy Consumption and

Solute Flux

Note: the results were obtained at constant temperature (25 oC) and osmotic agent concentration (275 and 200 g/L) for both sucrose and glucose

respectively

Flow

Rate

(L/min)

Water Flux

(Jw)[L/(m2.hr)]

Recovery Percent

(% R)

Water Permeability

(Aw)[L/(m2.hr.bar)]

Specific Energy

Consumption

(kWhr/m3)

Solute Flux

(Js)[g/(m2.hr)]


2 1.3 1 1 0.66 0.26 0.35 0.6 0.9 2.6 2.6

2.5 0.99 0.83 0.9 0.58 0.15 0.2 0.9 1.1 2.6 2.6

3 0.85 0.7 0.66 0.5 0.06 0.1 1 1.2 2.6 2.6

3.5 0.7 0.55 0.46 0.3 0.044 0.077 1.1 1.25 2.6 2.6

4 0.55 0.4 0.25 0.1 0.04 0.066 1.15 1.3 2.6 2.6


4.2.3.3. Effect on the Water Recovery Percentage (%R)

The experimental data in Table 4.13 show the effect of osmotic agents’ flow rate on the water

recovery percent (%R). It can be seen from the table that as the draw solution’s flow rate

increased from 2 to 4 L/min, the water recovery percent decreased from 0.66 % to 0.1 %

when sucrose was used as an osmotic agent and decreased from 1.0 % to 0.25 % when

glucose was used as an osmotic agent. Therefore, the results of the experiment showed that in

general increasing osmotic agent’s flow leads to a decrease in water recovery percent. Similar

to Jw, this decrease may be due to the increased osmotic potential difference between the feed

layer and the support layer of RO membrane surface. The initial and the final water recovery

percentages were higher for the processes where glucose was used as an osmotic agent.


The experimental data in Table 4.13 show the effect of osmotic agent’s flow rate on the

specific energy consumption (SEC). It can be seen from the table that as the draw solution’s

flow rate increased from 2 to 4 L/min, the specific energy consumption increased from 0.9 to

1.3 kWhr/m3 when sucrose was used as an osmotic agent and increased from 0.6 to 1.15

kWhr/m3 when glucose was used as an osmotic agent. Therefore, the results of the

experiment showed that in general increasing osmotic agent’s flow leads to an increase in

specific energy consumption. It is notable, however, that while the process involving sucrose

as an osmotic agent showed a steady increase in the specific energy consumption, the initial

rise in SEC was replaced by a steady decline when using glucose as an osmotic agent.

4.2.3.5. Discussion

The results of the experiment showed that water flux (Jw), recovery percentage (% R) and

water permeability (Aw) decreased and the specific energy consumption (SEC) showed

increases when the OAs flow rate was increased from 2 L/min to 4 L/min in FO process.

Increasing the OAs flow rate increased the concentration of solution near the membrane

surface which increases the osmotic pressure near the membrane surface thus resulting in

high net osmotic potential difference between the active layer and the support layer of the

membrane. This high NPD caused low water flux and increased the SEC requirement of FO

process. The low fluxes can be related to dilutive ECP on the OAs side of the membrane.

While glucose based solution demonstrated slightly better water flux and recovery rates,


sucrose based solution demonstrated higher water recovery rate and lower energy

consumption.

4.2.4. The Effects of Feed Water Temperature

The effect of feed water temperature was measured for water flux (Jw), water recovery

percent (%R), water permeability (Aw), specific energy consumption (SEC) and solute flux

(Js). The results of the experiments are demonstrated in Table 4.14.

4.2.4.1. Effect on Water Flux

The experimental data in Table 4.14 show the effect of feed water temperature on water flux.

It can be seen from the table that as the temperature increased from 20 to 35 oC, the water

flux increased from 0.7 to 2.2 L/(m2.hr) when sucrose was used as an osmotic agent and

increased from 0.5 to 1.9L/(m2.hr) when glucose was used as an osmotic agent. Therefore, in

the experiment, sucrose and glucose based draw solutions exhibited different effects on the

water flux. Normally, water flux increases when the temperature is increased due to faster

movement of water molecules, decreasing viscosity of OA and the difference in osmotic

potential across the membrane surface. Therefore, the behaviour exhibited by the process

based on sucrose osmotic agent requires further investigation.



recovery percentage. It can be seen from the table that as the temperature increased from 20

to 35 oC, the water recovery percent increased from 1.0 to 7.0% when sucrose was used as an

osmotic agent and increased from 0.6 to 6.0 % when glucose was used as an osmotic agent.

Therefore, the results of the experiment showed that in general increasing feed water

temperature leads to an increase in water recovery percent. Sucrose-based process had higher

ending water recovery percentages and showed a sharp increase in this parameter over the

course of the experiment. This may be due to the difference in viscosity between the two

osmotic agents which showed the effect of feed water temperature on % R in FO process.


Table ‎4.14. Effect of Feed Water Temperature on Water Flux, Recovery Percentage, Water Permeability, Specific Energy Consumption and

Solute Flux

Note: the results were obtained at constant feed and draw solution flow rate (2 L/min) and osmotic agent concentration (200 and 250g/L) for


Temperature

(o C)

Water Flux

(Jw)[L/m2.hr)]

Osmotic

Pressure

(bar)

Osmotic

Pressure

(bar)

Recovery Percent

(% R)

Water Permeability

(Aw)[L/(m2.hr.bar)]

Specific Energy

Consumption

(kWhr/m3)

Solute Flux

(Js)[g/(m2.hr)]


20 0.5 0.7 12.3 10.7 0.6 1 0.05 0.08 2.7 3 2.6 2.6

25 1 1.2 21.3 12.6 1.6 2.2 0.15 0.18 1.77 2 2.6 2.6

30 1.5 1.7 25.0 17.9 3.2 4 0.25 0.28 1 1.2 2.6 2.6

35 1.9 2.2 28.1 27.5 6 7 0.4 0.48 0.5 0.66 2.6 2.6




permeability. It can be seen from the table that as the temperature increased from 20 to 35 oC,

the water permeability increased from 0.08 to 0.48 L/(m2.hr.bar) when sucrose was used as an

osmotic agent and increased from 0.05 to 0.4 L/(m2.hr.bar)when glucose was used as an

osmotic agent. Therefore, the results of the experiment showed that in general increasing feed

water temperature leads to a decrease in water permeability for the cases involving each of

the osmotic agent solutions. Sucrose based process by the end of experiment demonstrated a

higher degree of water permeability. The increasing trend in water permeability may be due

to increased concentration of solution with increasing temperature which resulted in high

osmotic potential difference between the active layer and the support layer of membrane thus

causing an increase in Aw. However, the difference in Aw between the sucrose and glucose

OAs may be attributed to the difference in their respective viscosity.


The experimental data in Table 4.14 show the effect of feed water temperature on the specific

energy consumption. It can be seen from the table that as the temperature increased from 20

to 35 oC, the energy consumption decreased from 3 to 0.66 kWhr/m

3 when sucrose was used

as an osmotic agent and decreased from 2.7 to 0.5 kWhr/m3 when glucose was used as an


feed water temperature leads to a decrease in specific energy consumption for both types of

osmotic agents. However, by the end of the experiment the final permeability rates were

higher for the processes involving glucose as an osmotic agent. This may be due to the

increasing concentration effect for osmotic agents when feed water temperature rises. This

leads to a lower energy requirement to sustain the osmotic process. The effect was

particularly noticeable for the glucose-based osmotic process.


The experimental data in Table 4.14 show the effect of feed water temperature on the solute

flux. It can be seen from the table that as the temperature increased from 20 to 35 oC, the

solute flux remained constant at 2.6 g/(hr.m2) for both types of osmotic agents. Therefore, the

results of the experiment showed that there was constant effect of feed water temperature on

the solute flux.


4.2.4.6. Discussion

It can be seen from Table 4.14 that water flux, recovery percentage and water permeability

increased, the specific energy consumption decreased, but the solute flux was not affected by

increasing the feed water temperature from 20 to 35oC. It is evident that increasing the feed

water temperature affects the different water characteristics such as concentration, speed and

movement of water molecules, rate of diffusion and the osmotic pressure (driving force, Δπ)

near the membrane surface. This phenomenon causes high fluxes thus requiring low SEC

requirements as the FO process is mainly governed by the osmotic potential difference across

the membrane surface. These findings agree with the research of Phuntsho et al. (2012) and

You et al. (2012) as already mentioned in section 4.1.4.6.

4.2.5.‎The‎Effects‎of‎Osmotic‎Agents’‎Temperatures

The effect of osmotic agent’s temperature was measured for water flux (Jw), water recovery

percent (%R), water permeability (Aw) and specific energy consumption (SEC). The results

of the experiment are shown in Table 4.15.


The experimental data in Table 4.15 show the effect of temperature of draw solutions on the


the water flux increased from 0.1 to 0.5 L/(m2.hr)when sucrose was used as an osmotic agent

and increased from 0.17 to 0.6 L/(m2.hr)when glucose was used as an osmotic agent.

Therefore, the results of the experiment showed different behaviour exhibited by the

processes involving different osmotic agents. This difference may be due to the difference in

viscosity: sucrose’s higher viscosity level seems to make it more susceptible to the

temperature increases.


Table ‎4.15. Effect of Osmotic Agent Temperature on Water Flux, Recovery Percentage, Water Permeability, Specific Energy Consumption and

Solute Flux

Note: the results were obtained at constant feed and draw solution flow rates (2 L/min) and osmotic agent concentration (200 and 250g/L) for


Temperature

(o C)

Water Flux

(Jw)[L/(m2.hr)]

Recovery Percent

(% R)

Water Permeability

(Aw)[L/(m2.hr.bar)]

Specific Energy

Consumption

(kWhr/m3)

Solute Flux

(Js)[g/(m2.hr)]


20 0.17 0.1 0.13 0.2 0.03 0.02 1 1.2 2.6 2.6

25 0.27 0.2 0.24 0.25 0.05 0.04 0.8 1 2.6 2.6

30 0.4 0.3 0.34 0.38 0.07 0.06 0.6 0.8 2.6 2.6

35 0.6 0.5 0.43 0.55 0.1 0.09 0.5 0.7 2.6 2.6


4.2.5.2. Effect on the Water Recovery Percent

The experimental data in Table 4.15 show the effect of temperature of the draw solution on

the water flux. It can be seen from the table that as the temperature increased from 20 to 35

oC, the water recovery percent increased from 0.2to 0.55 % when sucrose was used as an

osmotic agent and increased from 0.13 to 0.43% when glucose was used as an osmotic agent.

Therefore, the results of the experiment showed that increasing temperature of an osmotic

agent leads to an increase in water recovery percent. Notably, for the glucose based process,

the increase was particularly sharp. This may be attributed to the effect of temperature on

viscosity, movement of water molecules and the rate of diffusion of water molecules across

the membrane surface in FO process.



the water flux. It can be seen from the table that as the temperature increased from 20 to 35

oC, the water permeability increased from 0.02 to 0.09 L/(m

2.hr.bar)when sucrose was used

as an osmotic agent and increased from 0.03 to 0.1 L/(m2.hr.bar) when glucose was used as

an osmotic agent. Therefore, the results of the experiment showed different effects of sucrose

based and glucose based processes on water permeability. Difference in the agents’ viscosity

could have contributed to the aforementioned differences.



the specific energy consumption. It can be seen from the table that as the temperature

increased from20to35oC, the specific energy consumption rate decreased from 1.2 to0.7

kWhr/m3

when sucrose was used as an osmotic agent and decreased from 1.0 to 0.5 kWhr/m3

when glucose was used as an osmotic agent. Therefore, the results of the experiment showed

different behaviour exhibited by the processes involving sucrose and glucose as osmotic

agents. The decrease in the SEC for both osmotic agents may be due to temperature effect

thus requiring less energy in FO process.


4.2.5.5. Discussion

It can be illustrated from data in Table 4.15 that increasing the OAs temperature from 20 to

35 oC increased the water flux, recovery percentage and water permeability, but reduced the

requirement of specific energy consumption in FO process. A explained earlier, increasing

the OAs temperatures will affect the different properties of OAs such as the viscosity, speed

and movement of water molecules, rate of diffusion and the osmotic pressure (driving force,

Δπ) near the membrane surface. Consequently, it will produce high fluxes and minimize the

SEC requirements of FO process which mainly depends on the osmotic potential difference

across the membrane surface. These findings agree with the results of Phuntsho et al. (2012)

and You et al. (2012) as explained earlier in section 4.1.4.6.

4.2.6. RO Membrane Results Overview

Several conclusions were drawn from this experiment for RO membrane (Table 4.16).The

best results for the given process and the given membrane were obtained by manipulating

feed water and osmotic agent temperatures. Increasing each of these factors led to higher

water flux, water recovery rate, and water permeability, although in the case of osmotic agent

temperature, an increase was noted for glucose-based solution only. Decrease in specific

energy consumption was noted for glucose based solutions as well. Notably, sucrose solution

showed superior results over glucose based solutions when changing feed water flow rate,

while glucose proved the superior osmotic agent when osmotic agent temperature was

increased. However, overall neither of the considered osmotic agents demonstrated an

absolute superiority when used with RO membrane. Out of twenty experiment measures

(without solute flux effects which were not affected), sucrose showed superiority in eleven.

While sucrose could be a promising agent to use with increasing feed water and glucose at

higher temperatures of osmotic agent, there is no certainty about either of these agents as an

agent of choice for the FO process using a RO membrane.


Table ‎4.16. Results of Changes of Manipulated FO Factors on Key Measured Parameters of

the Process: RO Membrane

Parameter Water Flux Water Recovery

Rate

Water

Permeability

Specific Energy

Consumption

Solute Flux

Increase in OA

concentration

Decreased

Sucrose superior

Decreased

Sucrose superior

Decreased

Sucrose superior

Increased

Glucose superior

No change

Increase in FW

flow

Increased

Sucrose superior

Increased

Sucrose superior

Increased

Sucrose superior

Increased

Sucrose superior

No change

Increase in OA

flow

Decreased

Glucose superior

Decreased

Glucose superior

Decreased

Glucose superior

Increased

Glucose superior

No change

Increase in FW

temperature

Increased

Sucrose superior

Increased

Sucrose superior

Increased

Sucrose superior

Decreased

Glucose superior

No change

Increase in OA

temperature

Increased

Glucose superior

Increased

Sucrose superior

Increased

Glucose superior

Decreased

Glucose superior

No change

4.3. A Comparison Study of Membrane - Osmotic Agent Systems

This section provides a side by side comparison of the FO processes for NF and RO flat sheet

membranes using sucrose and glucose as osmotic agents. In line with the data collected for

each type of membrane and for each type of osmotic agent, the comparisons of FO process

performance are made on the bases of osmotic agents’ concentrations, feed water flow rates,

osmotic agents’ flow rates, feed water temperature and osmotic agents’ temperature. The

compared effects are on water flux, water recovery percent, water permeability and specific

energy consumption. The purpose of the comparisons was to determine the best membrane

and an osmotic agent that can be used in FO process.

4.3.1.‎Osmotic‎Agents’‎Concentration


Figure 4.2 provides a comparative graph of the water fluxes obtained for each combination of

a membrane (NF and RO Flat Sheet) and an osmotic agent (sucrose and glucose). From the

graph, it can be seen that the highest fluxes were achieved for the NF membrane-glucose OA

and RO membrane-sucrose OA combination despite the steep decline in water flux

performance during the experiment. Specifically, the obtained fluxes were 3.5 and 1.5

L/(m2.hr) for 90 and 250 g/L concentrations of glucose, respectively. The RO membrane -

glucose OA system also showed a decline in water flux from 2.0 to 0.3L/(m2.hr)as the OA

concentrations increased from 170 to 400 g/L. Sucrose based systems with RO membrane

showed high water flux, but relatively low water fluxes at the initial concentrations (3.7 and


1.2L/(m2.hr) for RO and NF membranes, respectively) and these fluxes declined to 1.5 and

0.9 L/m2.hr for RO and NF membranes, respectively. Therefore, based on the experiment

results, sucrose and glucose were the preferable OAs, and the best results were demonstrated

in combination with the NF membrane type. The highest water flux of 3.7 L/(m2.hr)was

achieved by the RO membrane-sucrose OA system at sucrose concentration of 170 g/L.

Figure ‎4.2. Effect of OA concentration on Water Flux with RO and NF Using RO and NF

Membrane at Constant Temperature (25 oC) and Feed and Draw Solution Flow Rates (2

L/min)

4.3.1.2. Effect of OAs Concentration on Percent Recovery

Figure 4.3 provides a comparative graph of the water recovery percentages obtained for each

combination of a membrane (NF and RO Flat Sheet) and an osmotic agent (sucrose and

glucose). From the graph, it can be seen that the highest percentages were obtained for the NF

membrane-sucrose OA system with 2.5% and 0.8% recoveries at 170 and 400 g/L osmotic

agent concentrations, respectively. However, the sucrose systems for both membrane types

demonstrated rapid declines in water recovery as the osmotic concentrations increased. In

contrast, the glucose based systems showed steady decreases in water recovery with the

increased concentrations. This is especially true with regards to the RO membrane – glucose

OA system. It showed a decrease in water recovery percent from 0.5 % to 0.13 % as glucose

concentrations grew from 90 to 250 g/L. The final recovery rate exceeded the one

0

0.5

1

1.5

2

2.5

3

3.5

4

50 100 150 200 250 300 350 400 450

J w

L (

hr.

m2)

OA concentration (g/L)

Sucrose RO Glucose RO

Sucrose NF Glucose NF


demonstrated by the RO membrane-sucrose OA system. It remained unclear whether further

increase in OA concentrations could lead to the convergence of water recovery rates for the

RO membrane-glucose OA and NF membrane-sucrose OA systems. At the same time, the

available data showed clear inverse trends for sucrose-based and glucose-based systems at

increasing OA concentrations. The highest water recovery percent of 2.5% was achieved by

the NF membrane – sucrose OA system at sucrose concentration of 170 g/L.

Figure ‎4.3. Effect of OAs Concentration on Percent Recovery Using RO and NF Membrane at

Constant Temperature (25 oC) and Feed and Draw Solution Flow Rates (2 L/min)

4.3.1.3. Effect on of OAs Concentration on Water Permeability

Figure 4.4 provides a comparative graph of the water permeability obtained for each


glucose). From the graph, it can be seen that all systems showed similar pattern of water

permeability decline at the growing OA concentrations. Still, the best results for this specific

parameter were demonstrated by the RO membrane-sucrose OA system with 0.7 and 0.1

L/(m2.hr bar) at 170 g/L and 400 g/L concentrations, respectively. At the same time the NF

membrane-glucose OA system showed nearly the same level of water permeability at higher

concentrations: 0.08L/(m2.hr.bar)at 250 g/L. This system showed a slow rate of decline for

water permeability, although its initial permeability of 0.3 L/(m2.hr.bar)at 90 g/L glucose

0

0.5

1

1.5

2

2.5

3

50 100 150 200 250 300 350 400 450

% R

OA concentration (g/L)

Sucrose NF Glucose NFSucrose RO Glucose RO


concentration was lower than in the RO membrane-sucrose OA system (0.7L/(m2.hr bar)at

170 g/L sucrose concentration). Therefore, based on the experiment results, the RO

membrane-sucrose OA system can be considered superior based on water permeability

parameter, although at the higher concentration levels it is likely that the NF membrane-

glucose OA may become at least as much preferable. The highest water permeability value of

0.7 L/(m2.hr.bar)was obtained by the RO membrane-sucrose OA system at sucrose

concentration of 170 g/L.

Figure ‎4.4. Effect of OAs Concentration on Water Permeability Using RO and NF

Membranes at Constant Temperature (25 oC) and Feed and Draw Solution Flow Rates (2

L/min)


Figure 4.5 provides a comparative graph of the specific energy consumption obtained for

each combination of a membrane (NF and RO Flat Sheet) and an osmotic agent (sucrose and

glucose). From the graph, it can be seen that the sucrose based systems demonstrated the

higher energy requirements i.e.2.9 to 5.6 kWhr/m3for the RO system at 170 and 400 g/L,

respectively and 1.2 to 3.3 kWhr/m3for the NF system at the same concentrations. As was

discussed earlier, higher viscosity of sucrose and smaller pore structure of the RO membrane

make FO processes of such systems much more energy requiring. The process with the

lowest SEC rates was obtained by the NF membrane-glucose OA system with the increase

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

50 100 150 200 250 300 350 400 450

Aw (

L /

m2.h

r.b

ar)

OA concentartion (g/L)




from 0.3 to 0.9 kWhr/m3as glucose concentrations rose from 90 to 250 g/L. While the RO

membrane-glucose OA system showed an increase in SEC (from 1.0-2.0 kWhr/m3as glucose

concentrations rose from 90 to 250 g/L), the graph clearly indicates an upward trend for the

concentrations starting around 180-190 g/L. Therefore, based on the results of the experiment

the NF membrane-glucose OA system is preferable based on the SEC parameter. The lowest

SEC value of 0.3 kWhr/m3 was obtained by the NF membrane – glucose OA system at

glucose concentration of 90 g/L.

Figure ‎4.5. Comparison of Osmotic Agents’ Concentration Effect on Specific Energy

Consumption for NF and RO Membranes at Constant Temperature (25 oC) and Feed and

Draw Solution Flow Rates (2 L/min)

4.3.1.5. Discussion

The study of comparative graphs (Figures 4.2-4.5) show the effect of varying concentrations

of glucose and sucrose OAs on different FO parameters using NF and RO membrane. Based

on the experiment results, sucrose and glucose were the preferable OAs, and the best results

were demonstrated in combination with the NF membrane type. The highest water flux of

3.75 L/(m2.hr) was achieved by the RO membrane-sucrose OA system at glucose

concentration of 90 g/L. The available data on recovery percentage showed clear inverse

trends for sucrose-based and glucose-based systems at increasing OA concentrations. The

highest water recovery percent of 2.5% was achieved by the NF membrane-sucrose OA

0

1

2

3

4

5

6

50 100 150 200 250 300 350 400 450

SE

C (

kW

.hr/

m3)

OA cocnetration (g/L)




system at sucrose concentration of 170 g/L. With respect to water permeability, the RO

membrane-sucrose OA system can be considered superior, although at the higher

concentration levels it is likely that the NF membrane – glucose OA may become at least as

much preferable. The highest water permeability value of 0.7 L/(m2.hr.bar) was obtained by

the RO membrane – sucrose OA system at sucrose concentration of 170 g/L. Also, the data

showed that the NF membrane-glucose OA system is preferable based on the SEC parameter.

The lowest SEC value of 0.2 kWhr/m3 was obtained by the NF membrane – glucose OA

system at glucose concentration of 90 g/L. The study findings are closely related to those of

Peng (2004) and, Al-Sharif and Al-Mayahi, 2005) who concluded that though the NF

membranes are new entry in the field of RO membrane separation, but is quite different to

RO membrane due to having large membrane pore structure compared to RO membrane thus

allowing larger molecules through the membrane compared to RO membrane. The

advantages of using NF over RO membrane are many. Because NF membrane has a higher

water recovery than RO, thereby conserving total water use and energy. With the exception

of few results, the NF membrane with glucose and sucrose OAs seems better and preferable

on RO membrane with glucose and sucrose OAs in FO process.

4.3.2. Feed Water Flow Rate


Figure 4.6 provides a comparative graph of the water fluxes obtained for each combination of


graph, it can be seen that all four systems in the experiment showed a pattern of increasing

water flux as a result of an increase in feed water flow rate. Sucrose-based systems provided

the best results in terms of water flux. The best result overall was demonstrated by the RO

membrane – sucrose OA system with the fluxes ranging from 0.89 to 1.8 L/(m2.hr)as the feed

water flow rate increased from 2 to 4 L/min. The lowest water flux was demonstrated by the

NF membrane-glucose OA system with a range of 0.1 to 0.5 L/(m2.hr)at the same

corresponding feed flow rates. Therefore, based on the feed flow parameter, the preferable

FO system was with the RO membrane and sucrose as an osmotic agent. The highest water

flux value of 1.8 L/(m2.hr)was obtained by the RO membrane – sucrose OA system at feed

water flow of 4 L/min.


Figure 4.6.Comparison of Feed Water Flow Effect on Water Flux for NF and RO Membranes


and glucose respectively


Figure 4.7 provides a comparative graph of the water recovery percentage rates obtained for


glucose). From the graph, it can be seen that all systems showed increasing water recovery

rates with the increasing feed water flows. However, the RO membrane-sucrose OA system

showed clearly the best results which is seen from the graph. Specifically, the obtained water

recovery rates were 4.3% to 7.5% as feed water flow increased from 2 to 4 L/min. Therefore,

the RO membrane – sucrose OA system is preferable based on the water recovery percentage

parameter. The highest water recovery percent of 7.5% was obtained by the RO membrane –

sucrose OA system at feed water flow of 4 L/min.

0

0.2

0.4

0.6

0.8

1

1.2

1.4

1.6

1.8

2

1.75 2 2.25 2.5 2.75 3 3.25 3.5 3.75 4

Jw (

L/m

2.h

r)

Feed water flow rate (g/L)

Sucrose NF Glucose NF Sucrosee RO Glucose RO


Figure 4.7. Comparison of Feed Water Flow Effect on Water Recovery Percentage for NF

and RO Membranes at Constant Temperature (25 oC) andOAs Concentration (275 and 200

g/L) for both sucrose and glucose respectively




glucose). From the graph, it can be seen that all systems in the experiment demonstrated

increasing water permeability with the increased feed water flow. However, the RO

membrane-sucrose OA system demonstrated the best result with an increase of water

permeability from 0.1 to 0.5 L/(m2.hr.bar)as the feed water flow increased from 2 to 4 L/min.

Noticeably, the NF systems showed relatively poor improvement with the increased feed

water flow. The RO membrane-glucose OA system showed a progressively increasing water

permeability which increased from 0.02 to 0.2 L/(m2.hr.bar), although these values were still

much lower than the ones demonstrated by the RO membrane-sucrose OA system. Therefore,

based on water permeability parameter, the RO membrane-sucrose OA system is preferable

with the increasing feed water flow. The highest water permeability value of 0.5

L/(m2.hr.bar)was obtained for the RO membrane – sucrose OA system at feed water flow of 4

L/min.

0

1

2

3

4

5

6

7

8

1.75 2 2.25 2.5 2.75 3 3.25 3.5 3.75 4 4.25

% R

Feed water flow rate (L /min)

Sucrose NF Glucose NFSucrose RO Glucose RO


Figure 4.8. Comparison of Feed Water Flow Effect on Water Permeability for NF and RO

Membranes at Constant Temperature (25 oC) and OAs Concentration(275 and 200 g/L) for



Figure 4.9 provides a comparative graph of the specific energy consumption (SEC) obtained

for each combination of a membrane (NF and RO Flat Sheet) and an osmotic agent (sucrose

and glucose). From the graph, it can be seen that all systems except for the RO membrane –

sucrose OA in the experiment demonstrated declining SEC rates with the increasing feed

water flow. It is also noticeable that by the time that the feed water flow reached 4 L/min,

SEC demonstrated by all systems converged around the value of 0.6 –1.0 kWhr/m3.

However, at the lower flow levels, there is much difference in SEC. The NF membrane –

sucrose OA system showed the highest SEC at 12. kWhr/m3, while the RO membrane –

sucrose OA showed the lowest SEC at 0.4 kWhr/m3. Therefore, this particular system could

be preferable based on SEC parameter. The lowest SEC value of 0.4 kWhr/m3 was obtained

by the RO membrane-sucrose OA system at feed water flow of 4 L/min.

0

0.1

0.2

0.3

0.4

0.5

0.6

1.75 2 2.25 2.5 2.75 3 3.25 3.5 3.75 4 4.25

A

w

( L

/ m

2.h

r.b

ar )

Feed water flow rate ( L/ min)




Figure 4.9. Comparison of Feed Water Flow Effect on Specific Energy Consumption for NF

and RO Membranes at Constant Temperature and OAs Concentration (275 and 200 g/L) for


4.3.2.5. Discussion

The comparative graphs (Figures 4.6 to 4.9) represent the effect of feed water flow rate on

different parameters in FO process using NF and RO membrane with glucose and sucrose as

OAS. Based on the feed water flow rate, the highest water flux value of 1.8 L/(m2.hr)was

obtained by the RO membrane-sucrose OA system at feed water flow of 4 L/min.

Furthermore, based on the water recovery percentage parameter, the RO membrane-sucrose

OA system is preferable. The highest water recovery percent of 7.5% was obtained by the RO

membrane-sucrose OA system at feed water flow of 4 L/min. With respect to water

permeability parameter, the RO membrane-sucrose OA system is preferable with the

increasing feed water flow. The highest water permeability value of 0.5 L/(m2.hr.bar) was

obtained for the RO membrane-sucrose OA system at feed water flow of 4 L/min. The NF

membrane-sucrose OA system showed the highest SEC at 12 kWhr/m3, while the RO

membrane-sucrose OA showed the lowest SEC at 0.44 kWhr/m3. Therefore, this particular

system could be preferable based on SEC parameter. The lowest SEC value of 0.44 kW.hr/m3

was obtained by the RO membrane-sucrose OA system at feed water flow of 4 L/min.

Overall, it appears that RO membrane-sucrose OA is more preferable system for feed water

flow rate study to study the performance of FO process in water desalination. The study

0

2

4

6

8

10

12

14

1.75 2 2.25 2.5 2.75 3 3.25 3.5 3.75 4

SE

C (

kW

.hr/

m3 )

Feed water flow rate (L/min)




findings are closely related to those of Peng (2004) and, Al-Sharif and Al-Mayahi, 2005) who

concluded that though the NF membranes are new entry in the field of RO membrane

separation, but is quite different to RO membrane with respect to porosity, structure and the

diffusion coefficient thus allowing more passage of water molecules through the membrane

compared to RO membrane. The advantages of using NF over RO membrane are many.

Because NF membrane has a higher water recovery than RO, thereby conserving total water

use and energy. With the exception of few results, the NF membrane with glucose and

sucrose OAs seems better and preferable on RO membrane with glucose and sucrose OAs in

FO process.

4.3.3.‎Osmotic‎Agents’‎Flow‎Rate


Figure 4.10 provides a comparative graph of the water flux obtained for each combination of


graph, it can be seen that the NF based systems showed an increasing water flux with the

increased osmotic agents’ flow, while the RO systems showed a decreasing water flux

pattern. This means that in general NF systems could be preferable at the higher osmotic

agents’ flow rates while RO systems-at the lower rates. For both increasing and decreasing

patterns, however, the systems with glucose as an osmotic agent showed superior values over

the sucrose based systems. At the 2 L/min osmotic agent flow, the best performance was

demonstrated by the RO membrane – glucose OA system at 1.3 L/(m2.hr) level. At the 4

L/min osmotic agent flow, the best performance was demonstrated by the NF membrane –

glucose OA at 0.9 L/(m2.hr).Therefore, the results of the experiment demonstrated superiority

of the glucose OA on the basis of water flux parameter when osmotic agent flow rate is the

manipulated measure. The highest water flux of 1.3 L/(m2.hr) was obtained by the RO

membrane-glucose OA system at glucose solution flow rate of 2 L/min.


Figure 4.10. Comparison of Osmotic Agent Flow Effect on Water Flux for NF and RO

membranes at Constant Temperature (25 oC) and OAs Concentration (275 and 200 g/L) for



Figure 4.11 provides a comparative graph of the water recovery percent obtained for each


glucose). From the graph, it can be seen that the glucose based systems showed an increasing

water recovery percent with the increased osmotic agents’ flow, while the sucrose systems

showed a decreasing water recovery percent pattern. This means that in general glucose

systems could be preferable at the higher osmotic agents’ flow rates while sucrose systems –

at the lower rates. At the 2 L/min osmotic agent flow, the best performance was demonstrated

by the RO membrane-glucose OA system at 1.0% recovery rate. At the 4 L/min osmotic

agent flow, the best performance was demonstrated by the NF membrane-sucrose OA at

0.9%.

0

0.2

0.4

0.6

0.8

1

1.2

1.4

1.75 2 2.25 2.5 2.75 3 3.25 3.5 3.75 4 4.25

J w

( L

/m

2.h

r )

OA flow rate (L/min)




Figure 4.11. Comparison of Osmotic Agent Flow Effect on Water Recovery Percent for NF

and RO membranes at Constant Temperature (25 oC) and OAs Concentration (275 and 200

g/L) for both sucrose and glucose respectively




glucose). From the graph, it can be seen that the NF based systems showed increasingly

growing water permeability at growing osmotic agent flow rates. The NF membrane –

glucose OA system showed a permeability of 0.35 L/(m2.hr.bar) at 2 L/min flow rate, which

was the best result in the experiment for the high level of osmotic agent flow. The RO

membrane-sucrose OA system showed the highest permeability rate at the lower osmotic

agent flow: 0.35 L/(m2.hr.bar) at 2 L/min flow rate, although at the high levels of flow it

showed the worst permeability at just 0.06 L/(m2.hr.bar). At the same time, the RO

membrane-glucose system never showed water permeability rates close to the ones

demonstrated by the sucrose-based systems. Overall, the results of the experiment showed

that the NF systems showed the best performance in terms of water permeability, especially

the NF membrane-glucose OA system. The RO membrane-sucrose OA system could be

considered at low levels of osmotic agent flow, although it demonstrates steep decline in

permeability rates as the osmotic agent flow increases. The lowest water permeability rate of

0.1 L/(m2.hr.bar) was obtained by the NF membrane-glucose OA system at glucose solution

flow rate of 2 L/min.

0

0.2

0.4

0.6

0.8

1

1.2

1.75 2 2.25 2.5 2.75 3 3.25 3.5 3.75 4 4.25

% R

OA flow rate (L /min)




Figure 4.12. Effect of OAs Flow Rate on Water Permeability Using RO And NF Membranes

at Constant Temperature (25 oC) and OAs Concentration (275 and 200 g/L) for both sucrose

and glucose respectively

4.3.3.4. Effect of OAs Flow Rate on Specific Energy Consumption (SEC)



and glucose). From the graph, it can be seen that nearly all systems showed declining SEC

values as the osmotic agents’ flow increased. The RO membrane-sucrose OA was the only

system with an increasing SEC: from 0.9 to 1.3 kWhr/m3as the osmotic agents’ flow rate

increased from 2 to 4 L/min. However, this system also demonstrated the lowest SEC up to

the flow rate of around 3.5 L/min where its SEC value of 0.9 kWhr/m3was the same as for the

NF membrane-glucose OA system. At the higher flow rates, the SEC values for all four

systems converged at around 1.04 to 1.50 kWhr/m3which means that there could be particular

preference for a system at this stage. However, there were substantial differences at the lower

osmotic agent flow rates. The NF membrane-sucrose OA system showed the highest SEC

starting at 12 kWhr/m3 at the flow rate of 2 L/min. Although this system showed the steepest

decline in SEC as the flow rate increased, its performance was still poorer than for the other

systems. The results of the experiment showed that despite a decrease in SEC as a result of

increasing osmotic agent flow rate, the RO membrane-sucrose OA could be the preferred

system based on SEC parameter, at least for the lower levels of flow rates. The lowest SEC

0

0.05

0.1

0.15

0.2

0.25

0.3

0.35

0.4

1.75 2 2.25 2.5 2.75 3 3.25 3.5 3.75 4 4.25

Aw

( L

/m2.h

r.b

ar)

OA flow rate (L/min)




value of 0.7 kWhr/m3 was obtained by the RO membrane – glucose OA system at sucrose’s

solution flow rate of 2 L/min.

Figure 4.13. Effect of OAs flow rate on SEC using NF and RO membrane at Constant

Temperature (25 oC) and OAs Concentration (275 and 200 g/L) for both sucrose and glucose

respectively

4.3.3.5. Discussion

The results for the effect of OAs flow rate on different FO parameters are presented in

Figures 4.10 to 4.13 using NF and RO membranes. The NF membrane-glucose OA showed

the best performance of water flux at 0.9 L/(m2.hr). Therefore, the results of the experiment

demonstrated superiority of the glucose OA based on this parameter when osmotic agent flow

rate is the manipulated measure. The highest water permeability rate of 0.35L/(m2.hr.bar) was

obtained by the NF membrane – glucose OA system at glucose solution flow rate of 4 L/min.

The recovery percentage was highest (0.9 %) at the 4 L/min osmotic agent flow by the NF

membrane-sucrose OA. Although, a decrease in SEC was found as a result of increasing

osmotic agent flow rate, the RO membrane-sucrose OA could be the preferred system based

on SEC parameter, at least for the lower levels of flow rates. The lowest SEC value of 0.6

kWhr/m3 was obtained by the RO membrane-glucose OA system at sucrose’s solution flow

rate of 2 L/min. The results agree with those of Peng (2004) and, Al-Sharif and Al-Mayahi,

0

1

2

3

4

5

6

7

8

9

1.75 2 2.25 2.5 2.75 3 3.25 3.5 3.75 4

SE

C (

kW

.hr/

m3)

DS flow rate (L/min)




2005) who reported that the difference in characteristics of the two membranes (NF and

RO)namely porosity, structure and the diffusion coefficient play a major role in membrane

selection depending on the experimental conditions such as flow rate, temperature and the

concentration of OAs.

4.3.4. Feed Water Temperature

4.3.4.1. Effect of Feed Water Temperature on Water Flux



graph, it can be seen that the systems demonstrated different responses to the increase in feed

water temperature. As the feed water temperature rose from 20to 35oC, the RO membrane-

sucrose OA system demonstrated a water flux increase from 0.7 to 2.2 L/(m2.hr); the RO

membrane-glucose OA system demonstrated a moderate increase from 0.5 to 1.9 L/(m2.hr);

and both NF systems remaining at approximately same level of water flux at 0.5-1.1 and 0.2-

0.8 L/(m2.hr). Therefore, the experiment results demonstrated superiority of the RO

membrane systems. The system with by far the best water flux performance was the RO

membrane-sucrose OA, which can, therefore, be considered as preferable based on the

experiment results. The highest water flux of 2.2 was obtained by the RO membrane-sucrose

OA system at feed water temperature of 35oC.


Figure ‎4.14. Effect of Feed Water Temperature on Water Flux Using NF and RO Membranes


250 g/L) for both sucrose and glucose respectivly


Figure 4.15 provides a comparative graph of the water recovery percentages obtained for


glucose). From the graph, it can be seen that both RO and NF systems showed an apparent

increase in water recovery percentages as the feed water temperature grew. This makes the

RO systems preferable based on water recovery percent. The system with by far the highest

values for water recovery percent was the RO membrane-sucrose OA with the water recovery

percentage increased from 1.0 % to 7.0 as the temperature of feed water increased from 20 to

35oC. Also, similar trend of increase in recovery percentage was followed by RO membrane-

glucose OA where it ranged between 0.5 to 6.0 % as the temperature of feed water increased

from 20 to 35oC. This makes both these systems preferable based on the water recovery

percent parameter. The highest water recovery percent of 7.0 % was obtained by the RO

membrane-sucrose OA system at feed water temperature of 35oC.

0

0.5

1

1.5

2

2.5

17.5 20 22.5 25 27.5 30 32.5 35

J w (

L/m

2.h

r )

feed water teprature (°C)




Figure ‎4.15. Effect of Feed Water Temperature on Percent Recovery Using NF and RO


(200 and 250 g/L) for both sucrose and glucose respectivly




glucose). From the graph, it can be seen that the RO systems showed gradual increase in

water permeability as a result of feed water temperature increase while the NF systems

showed a gradual decline in this parameter. At the lower feed water temperature of 20 oC, the

NF membrane-glucose OA system showed the best water permeability value of 0.6

L/(m2.hr.bar).The system showed the best water permeability starting at the point of around

23-23.5 oC, and at the highest temperature of 35

oC, it showed the best water permeability

value of just 0.6 L/(m2.hr.bar). Therefore, the preferable system based on water permeability

is the RO membrane-glucose OA, which showed the second best water permeability value of

0.05 L/(m2.hr.bar) when the feed water temperature was at 20

oC, the same best value as the

NF membrane-sucrose OA system at the point of around 23-23.5 oC and by far the best value

thereafter, peaking at 0.6 L/(m2.hr.bar) at feed water temperature of 35

oC.

0

1

2

3

4

5

6

7

8

17.5 20 22.5 25 27.5 30 32.5 35

% R

Feed water temprature (°C )




Figure ‎4.16. Effect of Feed Water Temperature on Water Permeability Using NF and RO


(200 and 250 g/L) for both sucrose and glucose respectivly




and glucose). From the graph, it can be seen that both the NF and RO membrane systems

showed a rapid decline in the SEC value as the feed water temperature increased. The NF

membrane-sucrose OA system showed the lowest SEC for most of the experiment which

declined from 3.0 to 0.66 kWhr/m3 as the feed water temperature grew from 20 to 35

oC.

However, at around 32 oC, the RO membrane-sucrose OA system started demonstrating

better SEC values. Eventually, at the feed water temperature of 35oC, this system showed the

most efficient energy use with the SEC value of just 1.1 kWhr/m3. Therefore, the results of

the experiment demonstrated that the NF membrane-sucrose OA system could be preferable

except when feed water temperatures achieve higher points; from there, the RO membrane –

sucrose OA system could be preferable. The initial lowest SEC value of 2.7 kWhr/m3was

obtained by the NF membrane-sucrose OA system at feed water temperature of 20 oC.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

17.5 20 22.5 25 27.5 30 32.5 35

Aw

( L

/m2.h

r.b

ar)

Feed water tempreture (°C)




Figure ‎4.17. Effect of Feed Water Temperature on Specific Energy Consumption Using NF

and RO Membranes at Constant Feed and Draw Solution Flow Rate (2 L/min) and OAs

concentration (200 and 250 g/L) for both sucrose and glucose respectivly

4.3.4.5. Discussion

The data in Figures 4.14 to 4.17 illustrate the effect of feed water temperature on various FO

parameters. The best water flux performance was obtained by the RO membrane-sucrose OA,

which can, therefore, be considered as preferable. The highest water flux of 2.37 was

obtained by the RO membrane-sucrose OA system at feed water temperature of 35 oC.

Whereas, the highest water recovery percent of 7.0 % was obtained by the RO membrane-

sucrose OA system at feed water temperature of 35oC. Also, the preferable system based on

water permeability is the RO membrane-glucose OA., which showed the second best water

permeability value of 0.10 L/(m2.hr.bar) when the feed water temperature was at 20

oC. Also,

The initial lowest SEC value of 2.5 kWhr/m3 was obtained by the NF membrane-sucrose OA

system at feed water temperature of 20 oC and is considered as the preferable over RO with

OAs combinations. These results are supported by Peng (2004) and, Al-Sharif and Al-

Mayahi, 2005) as presented in details is section 4.3.3.5.

0

1

2

3

4

5

6

7

8

9

17.5 20 22.5 25 27.5 30 32.5 35

SE

C (

kW

.hr/

m3 )

Feed water tempreture (°C)




4.3.5 Osmotic Agent Temperature

4.3.5.1. Effect of Osmotic Agent Temperature on Water Flux



graph, it can be seen that the RO membrane-glucose OA outperformed the others virtually at

every point of the graph. The system showed an increase of water flux from 0.17 to 0.6

L/(m2.hr)as the osmotic agent temperature rose from 20to 35

oC. This system is the preferable

system based on the water flux parameter. The highest water flux value of 0.6 L/(m2.hr)was

achieved by RO membrane-glucose system at glucose temperature of 35 oC.

Figure ‎4.18. Effect of OA Temperature on Water Flux Using RO and NF Membranes at




Figure 4.19 provides a comparative graph of the water recovery percent obtained for each


glucose). From the graph, it can be seen that both the systems showed an increase in water

recovery percent as a result of osmotic agent temperature increase was the NF membrane –

glucose OA. The highest recovery percent at lower temperatures, specifically, 0.4 % at 20 oC

of an osmotic agent. However, it showed a sharp increase in the parameter thereafter: at

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

17.5 20 22.5 25 27.5 30 32.5 35

J w (

L /

m2.h

r )

OA temprature (°C)




around 24.5 oC it no longer showed the best recovery rate, and at 35

oC of the sucrose

osmotic agent it showed the lowest recovery rate at just 0.2%. The best water recovery

percentages among the remaining systems were demonstrated by the NF membrane-sucrose

OA system. Its water recovery percentage showed the steepest increase and grew from 0.4%

to 1.7 % as the osmotic agent temperature increased from 20 to 35o C. This is, therefore, the

preferable system based on the water recovery percentage parameter. The highest water

recovery percent of 1.7 % was obtained by the NF membrane-sucrose OA system at sucrose

solution temperature of 35 oC.

Figure ‎4.19. Effect of OA Temperature on Percent Recovery Using NF and RO Membranes at






glucose). From the graph, it can be seen that the only system to demonstrate increasing water

permeability corresponding to increasing osmotic agent temperature was the RO membrane-

glucose OA. Specifically, it showed an increase of water permeability from 0.03 to 0.1

L/(m2.hr.bar) as the osmotic agent’s temperature grew from 20 to 35

o C. However, at the

lower temperatures (up until 27 oC) of an osmotic agent, the RO membrane-glucose OA

0

0.5

1

1.5

2

2.5

17.5 20 22.5 25 27.5 30 32.5 35

% R

OA temprature (°C)




system demonstrated the superior results. Specifically, it showed the highest rate of water

permeability as 0.03 L/(m2.hr.bar) at osmotic agent’s temperature of 20

oC. Therefore, the

results of the experiment showed that based on water permeability parameter, the preferable

system at the lower temperatures of osmotic agent is the RO membrane-glucose OA system

while the NF membrane-glucose OA is preferable at the higher temperatures of osmotic

agent. The highest permeability rate in the course of the experiment was 0.1 L/(m2.hr.bar)for

the RO membrane-glucose OA system at the osmotic agent’s temperature of 35 oC.

Figure ‎4.20.Effect of OA Temperature on Water Permeability Using NF and RO Membranes



4.3.5.4. Effect of OA Temperature on Specific Energy Consumption (SEC)



and glucose). From the graph, it can be seen that the RO membrane-glucose OA system was

the only system to demonstrate the declining SEC value as the osmotic agent temperature

grew from 20 to 35oC. Specifically, the system’s SEC declined from 1.1 to 0.6 kWhr/m

3. At

the lower temperatures, the RO membrane -glucose OA system showed the lowest SEC: at

the osmotic agent’s temperature of 20 oC, the system’s SEC was 1.0kWhr/m

3. However, this

system also demonstrated the highest decline in SEC as the osmotic agent’s temperature

grew: it reached 0.6kWhr/m3 at glucose solution temperature of 35

oC. Therefore, the results

0

0.02

0.04

0.06

0.08

0.1

0.12

17.5 20 22.5 25 27.5 30 32.5 35

Aw (

L/m

2.h

r.b

ar)

OA temprature (°C )

Sucrose NF Glucose NF Sucrose RO Glucose RO


of the experiment showed that the NF membrane-glucose OA system could be preferable at

the lower temperatures of osmotic agent; however, the RO membrane-glucose OA system

would be the choice for a wider range of osmotic agent’s temperatures.

Figure ‎4.21. Effect of OA Temperature on Specific Energy Consumption Using NF and RO

Membranes at Constant Feed and Draw Solution Flow Rate (2 L/min) and and OAs

concentration (200 and 250 g/L) for both sucrose and glucose respectivly

4.3.5.5. Discussion

The comparison for the effect of temperature of the draw solution on water flux, recovery

percentage, water permeability and specific energy consumption is presented in Figures 4.18

to 4.21 using NF and RO membranes. The RO membrane-glucose OA showed an increase of

water flux from 0.17 to 0.6 L/(m2.hr) as the osmotic agent temperature was increased from 20

to 35 oC. This system is the preferable system based on the water flux parameter as it

outperformed the other NF membrane system. The highest water flux value of 0.6 L/(m2.hr)

was achieved by RO membrane-glucose system at glucose temperature of 35 oC. Therefore,

based on the water recovery percentage, the NF membrane-sucrose OA system is the most

suitable system at sucrose solution temperature of 35 oC with the highest water recovery

percent of 1.75 %. Furthermore, depending on the temperature and to achieve the highest

water permeability, the RO membrane-glucose system is more suitable at lower temperature.

While the NF membrane-glucose OA is preferable at the higher temperatures of osmotic

0

0.2

0.4

0.6

0.8

1

1.2

1.4

1.6

17.5 20 22.5 25 27.5 30 32.5 35

SE

C (

kW

.hr/

m3 )

OA temprature (°C )




agent. With respect to SEC, the NF membrane-glucose OA system could be preferable at the

lower temperatures of osmotic agent to obtain higher SEC, while the RO membrane- glucose

OA system could be a better preposition for a wider range of osmotic agent’s temperatures.

4.3.6. Generalised Results for the Membrane-Osmotic Agent Systems

Table 4.17 summarises the outcomes of the experiments which allowed the best combinations

of membranes and osmotic agents to be determined based on manipulated factors and the

experiment parameters. Based on the experiment outcomes it is difficult to distinguish the

best system. The RO membrane-sucrose osmotic agent combination most frequently

accounted for the best results across different measured parameters, although the

combinations of NF-sucrose as an osmotic agent and NF-glucose as an osmotic agent were

not far behind. Because of this, it cannot be concluded with utmost confidence that any

combination is superior to the others within the outcomes of the conducted experiments.

Different systems could turn out better for specific parameter measures because of different

membrane and osmotic agent characteristics. Such characteristics could be, for example,

membrane compactness (NF membranes are more loose) and viscosity of the osmotic agent

materials. Accordingly, different membrane-osmotic agent systems would demonstrate

different sensitivity degrees to changes in various factors considered in this study. Therefore,

it is difficult to make a judgement about any system being superior in increasing efficiency of

the FO process.

Table ‎4.17. The Best Combinations of Membranes and Omsotic Agents Based on Experiment

Factors and Parameters

Water Flux Water Recovery

Rate

Water

Permeability

Specific Energy

Consumption

OA

Concentration NF-gl NF-su RO-su NF-gl

FW Flow Rate RO-su RO-su RO-su RO-su at lower flow rates

NF-gl at higher flow rates

OA Flow Rate

RO-gl at lower

flow rates

NF-gl at higher

flow rates

RO-gl at lower flow

rates

NF-su at higher flow

rates

NF-gl RO-gl

FW Temperature RO-su RO-su NF-su or RO-su RO-gl

OA Temperature RO-gl NF-su RO-gl RO-gl

*gl stands for glucose

**su stands for sucrose


4.4. Binary and Ternary Osmotic Systems

4.4.1. Binary System with Deionised and Brackish Water as the Feeds

A series of experiments were conducted for the binary systems involving an osmotic agent

(sucrose or glucose) and feed water (deionised water or sea water) to determine the optimal

combinations for the most effective FO process performance. The experiments were run at

varying osmotic pressure between 10 and 60 bar using a hollow fine fibre membrane

(HFFM). The systems’ performances were measured on the bases of water flux, water

recovery percent, water permeability, specific energy consumption and solute flux verses

OPD.


Figure 4.22 presents a comparative graph of the water flux values obtained for binary sucrose

and glucose systems against deionised and brackish waters as the feeds. The water flux was a

function of the osmotic pressure differential (OPD) which was varied between 10 to 60 bar

for the deionised water (DW) and brackish water (BW) experiments. As is seen from the

graph, the highest fluxes were obtained by the systems using deionised water (DW) as the

feed water. For the binary glucose-DW system, the water flux increased from 0.4 to 0.88

L/(m2.hr), and for the binary sucrose-DW system, the water flux increased from 0.35 to 0.83

L/(m2.hr). These values were substantially larger than the ones obtained by the systems using

brackish water (BW) as the feed water: the binary sucrose-BW system showed an increase

from 0.21 to 0.70 L/(m2.hr)and the binary glucose-BW system showed an increase from 0.29

to 0.76 L/(m2.hr)as the osmotic pressure differential increased from 10 to 60 bar. The

variation in the water flux between the binary systems may be due to the internal

concentration polarization (ICP) and the external concentration polarization (ECP) effects in

DW as compared to BW. Since the osmotic pressure of DW is zero, only ECP affects water

flux, thereby making it stronger than for BW feed where both ICP and ECP effects are

present. The binary glucose-DW system showed the highest water flux value of 0.70

L/(m2.hr) at the 60 bar OPD.


Figure ‎4.22. Osmotic Pressure Differential Effect on Water Flux in a Binary System


Figure 4.23 presents a comparative graph of the water recovery percentage values obtained

for binary sucrose and glucose systems with deionised and sea waters as the feed water. The

water recovery percentage was a function of the osmotic pressure differential (OPD) which

was manually varied between 10 to 60 bar for the deionised water (DW) and brackish water

(BW) experiments. As is seen from the graph, all systems demonstrated higher water

recovery percentages as a result of increasing OPD. The binary sucrose-DW system,

however, demonstrated the best recovery percentages over the entire course of the experiment

with the values increasing from 35 % to 38 % as the OPD grew from 10 to 60 bar. Therefore,

based on the results of the experiment, the sucrose-DW binary system is preferable based on

water recovery percentage criterion.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 10 20 30 40 50 60 70

J W (

L /

m2.h

r)

Osmotic pressure differential ( Δπ in bar)

Binary sucrose vs DW Bianry glucose vs DW

Binary glucose vs BW Binary sucrose vs BW


Figure ‎4.23. Osmotic Pressure Differential on Percent Recovery in Binary Sucrose and

Glucose System


Figure 4.24 presents a comparative graph of the water permeability values obtained for binary

sucrose and glucose systems with deionised and brackish waters as the feed water. The water

permeability was a function of the osmotic pressure differential (OPD) which was manually

varied between 10 to 60 bar. As is seen from the graph, the systems involving deionised

water (DW) as feed water demonstrated a steady decline in permeability as the OPD

increased from 10 to 60 bar. Specifically, the binary glucose-DW system demonstrated an

increase in water permeability from 0.14 to 0.43 L/(m2.hr.bar), and the binary sucrose-DW

system demonstrated an increase from 0.11 to 0.40 L/(m2.hr.bar). In contrast, the systems

using brackish water (BW) as feed water showed an increase in water permeability as the

OPD increased from 35 to 55 bar: the binary sucrose-BW system demonstrated an increase in

water permeability from 0.04 to 0.32 L/(m2.hr.bar), and the binary glucose-BW system

demonstrated an increase in water permeability from 0.075 to 0.36 L/(m2.hr.bar).

33.5

34

34.5

35

35.5

36

36.5

37

37.5

38

38.5

0 10 20 30 40 50 60 70

% R


Binary glucose vs DW Binary surose vs DW

Binary sucrose vs BW Binary sucrose vs BW


Figure ‎4.24. Osmotic Pressure Differential on Water Permeability in Binary Sucrose and

Glucose System

4.4.1.4. Effect of OPD on Specific Energy Consumption (SEC)

Figure 4.25 presents a comparative graph of the specific energy consumption (SEC) values

obtained for binary sucrose and glucose systems with deionised and brackish waters as the

feed water. The SEC was a function of the osmotic pressure differential (OPD) which was

manually varied between 10 to 60 bar for the experiments involving deionised water (DW)

and brackish water (BW). As is seen from the graph, all systems demonstrated a steady

increase in SEC with the increased OPD levels. The higher SEC values were measured for

the binary sucrose systems. Specifically, the sucrose-DW system showed an increase from

0.10 to 0.166 kWhr/m3 as the OPD increased from 10 to 60 bar, while the glucose-DW

system showed an increase from 0.09 to 0.15 kWhr/m3as the OPD increased from 10 to 60

bar. For the binary glucose systems the measured SEC values were substantially lower: from

0.08 to 0.15 kWhr/m3

for the glucose-DW system as the OPD grew from 10 to 60 bar, and

from 0.07 to 0.135 for the glucose-BW as the OPD grew from 10 to 60 bar. Based on the

results of the experiment, binary glucose systems showed better performance in terms of

SEC.

0

0.05

0.1

0.15

0.2

0.25

0.3

0.35

0.4

0.45

0.5

0 10 20 30 40 50 60 70

AW

(L

/m

2.h

r.b

ar)


Binary glucose vs DW Binary sucrose vs DW

Binary glucose vs BW Binary sucrose vs BW


Figure ‎4.25. Osmotic Pressure Differential Effect on SEC of Sucrose and Glucose OAs in a

Binary System

4.4.1.5. Effect of OPD on Solute Flux in sucrose and glucose Binary System

Figure 4.26 presents a comparative graph of the solute flux values obtained for binary sucrose

and glucose systems with deionised and sea waters as the feed water. The sugar flux was a

function of the osmotic pressure differential (OPD) which was varied between 10 to 60 bar

for the experiments involving deionised water (DW) and brackish water (BW). As is seen

from the graph, all systems in the experiment demonstrated an increase in sugar flux values

as a result of growing OPD. However, the rates of this increase were different. The systems

with deionised water as the feed water showed a gradual increase: from 2.5 to 3.75

g/(m2.hr)for the sucrose-DW system and from 2.0 to 3.58 g/(m

2.hr)for the glucose-DW

system. In contrast, the systems with brackish water as the feed water demonstrated sharp

increases in solute fluxes: from 0.13 to 3.0 L/(m2.hr)for glucose-BW and 1.6 to 3.35

L/(m2.hr) for sucrose-BW binary systems.

0

0.02

0.04

0.06

0.08

0.1

0.12

0.14

0.16

0.18

0.2

5 10 15 20 25 30 35 40 45 50 55 60 65

SE

C (

kW

.hr/

m3

)

Osmotic Pressue differntial (Δπ in bar )

Binary sucrose vs DW Binary glucose vs DW

Bianry glucose vs BW Binary sucrose vs BW


Figure4.26. Osmotic Pressure Differential Effect on Solute Fluxin Binary Sucrose and Binary

Glucose Systems

4.4.1.6. Discussion

Overall, the binary study involving osmotic agents (glucose and sucrose) and two feed waters

DW and BW showed the effect of OPD on different FO parameters such Jw, % R, Aw, SEC

and Js is presented in Figures 4.22 to 4.26. It was noticed that the values of all these

parameters increased substantially with increasing the OPD. But the main effect on these

parameters in a binary system was the use of using either brackish as feed water as well as

two different osmotic agents varying in viscosity (glucose being less viscous than sucrose.

Therefore, the varying response of different parameters is due to the OA and the feed

solution. The binary glucose-DW system showed the highest water flux value of 0.70

L/(m2.hr) at the 60 bar OPD compared to other binary systems. Whereas, the sucrose-DW

binary system is preferable based on water recovery percentage criterion. On the other hand,

the binary sucrose-BW system demonstrated an increase in water permeability from 0.03 to

0.30 L/(m2.hr.bar), and the binary glucose-BW system demonstrated an increase in water

permeability from 0.07 to 0.35 L/(m2.hr.bar). While, the binary glucose systems showed

better performance in terms of SEC. The systems with deionised water as the feed water

showed a gradual increase from 2.5 to 3.5 g/(m2.hr) in solute flux or the sucrose-DW system

and from 2.0 to 3.2 g/(m2.hr)for the glucose-DW system. In contrast, the systems with

0

0.5

1

1.5

2

2.5

3

3.5

4

5 10 15 20 25 30 35 40 45 50 55 60

J s (

g /

m2.h

r)

Osmotic pressure differential (Δπ in bar )

Bianry sucrose vs DW Binary glucose vs DW

Binary sucrose vs BW Binary glucose vs BW


brackish water as the feed water demonstrated sharp increases in solute fluxes: from 0.13 to

2.8 kW.hr/m3 for both sucrose and glucose binary systems.

The study results agree with those of Achilli et al. (2010) and Jeffery et al. (2006) who

reported that when using deionized water as feed water, then the differences in water fluxes

achieved by the OA (DS) may be due to the ICP in the system. They further stated that with

brackish water as feed water both the ICP and ECP play a major role in reducing the effective

osmotic pressure difference across the membrane surface thus affecting the different FO

parameters in a binary system. In an other study, Zhao and Zho (2011) reported that the

differences in the magnitude of ICP between OAs at the same osmotic pressure can be

reasonably explained by the difference in the viscosity of the osmotic agents in the study.

Jeffrey et al. (2006) concluded that the penetration of an OA with higher viscosity is much

less through the membrane layer, thereby, increasing the presence of ICP. Also, the diffusion

of solute trough the membrane porous structure is much less for a higher viscous OA than a

less viscous OA. Therefore, the flux performance of more viscous OAs will be less than the

OA with lower viscosity. In another study, Yuan Xu et al. (2010) stated that besides viscosity

of OAs, the type of interaction between the solute and the membrane may be another factor

affecting the OAs flux performance.

4.4.2. Binary System with Seawater as the Feed water

A series of experiments were conducted with sucrose and glucose as osmotic agents solution

and seawater as a feed water. The experiments were run at varying osmotic pressure between

13 and 33 bar for a hollow fine fibre membrane (HFFM). The systems’ performances were

measured on the bases of water flux and water recovery percent.


Figure 4.27 presents a comparative graph of the water flux values obtained for binary sucrose

and glucose systems when brackish was used as the feed water. The water flux was a function

of the osmotic pressure differential (OPD) which was varied between 13 and 33 bar during

the experiment. As is seen from the graph, the sucrose based binary system showed a decline

in the water flux values from 0.19 to 0.11 L/(m2.hr) as the OPD increased. In contrast, the

glucose based system demonstrated a gradual increase in water flux: from .08 to 0.124

L/(m2.hr). Since both the seawater and the osmotic agents exert osmotic pressure across the

membrane surface, the water flux is affected by the magnitude of ICP and ECP across the


membrane surface (active layer and the support layer sides). This may explain the different

behaviour of the systems with the different osmotic agents. Based on the results of the

experiment, binary sucrose system has a better water flux performance under lower OPDs,

while glucose based binary systems- under higher OPDs.

Figure ‎4.27. Comparison of Osmotic Pressure Differential Effect on Water Flux in Binary

Osmotic Systems and Seawater as the Feed Water


Figure 4.28 presents a comparative graph of the water recovery percent values obtained for

binary sucrose and glucose systems when seawater was used as the feed water. The water

recovery percent was a function of the osmotic pressure differential (OPD) which was

manually varied between 13 and 33 bar during the experiment. As is seen from the graph, the

sucrose based binary system showed a decline in the water recovery percentage values: from

16% to 13% over the course of the experiment. In contrast, the glucose based system

demonstrated a gradual increase in water recovery percentage: from 12.5 % to 16 % at the

same OPD levels. The increase in % R in glucose OA plus seawater as feed water in the

binary system may due to ICP and ECP in the binary system as well as on the level of

ionization of glucose as compared to sucrose OA and seawater as feed solution in the binary

system. Also this variation in % R may be due to the difference in the viscosity of the two

organic osmotic agents being more for sucrose than glucose OA. In general, based on the

water recovery percent parameter, binary sucrose systems may be preferable under lower

0

0.02

0.04

0.06

0.08

0.1

0.12

0.14

0.16

0.18

0.2

10 15 20 25 30 35

J w (

L /

m2.h

r )

Osmotic pressure diffrentical (∆𝜋 in bar)

Sucrose OA vs SW Glucose OA vs SW


OPD levels (in this experiment, under 25 bar), while glucose binary systems-for the higher

OPD levels (in this experiment, over 25 bar).

Figure ‎4.28. Comparison of Osmotic Pressure Differential Effect on Water Recovery Percent

in Binary Osmotic Systems and Seawater as The Feed Water

4.4.2.3. Discussion

It follows from Figures 4.27 and 4.28 that water flux and the recovery percentage were not

appreciably affected in a binary system using seawater as feed along with glucose and

sucrose as OAs. There was a mix response for the effect of OPD on these two parameters

using seawater as feed water. Based on the results of the experiment, binary sucrose system

using seawater as feed water has a better water flux performance under lower OPDs, while

glucose based binary systems-under higher OPDs. In general, based on the water recovery

percent parameter, binary sucrose systems using seawater as feed may be preferable under

lower OPD levels (in this experiment, under 25 bar), while glucose binary systems – for the

higher OPD levels (in this experiment, over 25 bar). Similar findings were reported by Achilli

et al. (2010) and Jeffery et al. (2006) who found that when using seawater or brackish water

as feed water both the ICP and ECP play a major role in reducing the effective osmotic

pressure difference across the membrane surface thus affecting the different FO parameters in

a binary system.

0

2

4

6

8

10

12

14

16

18

10 15 20 25 30 35

% R

Osmotic pressure diffrentical (∆𝜋 in bar)

Sucrose OA vs SW Glucose OA vs SW


4.4.3. Ternary Systems

Several ternary osmotic systems in FO were experimentally examined in this study in order to

selection optimal osmotic agent by considering its transport behaviour. The considered

ternary systems included an osmotic agent (glucose or sucrose), salt (NaCl) and a feed water

(deionised water or brackish water). All experiments were carried out at osmotic pressure

differential from 20 to 60 bar.

Four different ternary sucrose systems were tested during the experiment:

- A mixture with the constant sucrose concentration of 134.20g/L (equivalent to 10 bar

osmotic pressure), varying NaCl concentration (between 10 and 60 bar and deionised

water as a feed water;

- A mixture with the constant NaCl concentration of 12.73 g/L (equivalent to 10 bar

osmotic pressure), varying sucrose concentration (between 10 and 60 bar) and

deionised water as a feed water;

- A mixture with the constant sucrose concentration of 134.20 g/L, varying NaCl

concentration (between 10 and 60 bar) and brackish water as a feed water;

- A mixture with the constant NaCl concentration of 12.73g/L, varying sucrose

concentration (between 10 and 60 bar) and brackish water as a feed water.

Correspondingly, four different glucose systems were tested during the experiment:

- A mixture with the constant glucose concentration of osmotic pressure bar, varying

NaCl concentration (between 10 and 60bar) and deionised water as a feed water;

- A mixture with the constant NaCl concentration of 12.73g/L, varying glucose

concentration (between 10 and 60 bar) and deionised water as a feed water;

- A mixture with the constant glucose concentration of 134.20 g/L, varying NaCl

concentration (between 10 and 60 bar) and brackish water as a feed water;

- A mixture with the constant NaCl concentration of 12.73g/L, varying glucose

concentration (between 10 and 60 bar) and brackish water as a feed water.

4.4.3.1. Sucrose Ternary Systems and Water Recovery Percent


four different ternary sucrose systems. The water recovery percent was a function of the

osmotic pressure differential (OPD) which was varied between 20 and 60 bar during the

experiment. As is seen from the graph, all systems demonstrated increasing water recovery

patterns with the increasing OPD values. However, the highest recovery values throughout

the experiment were obtained by the sucrose (fix) + NaCl (var) / DW system. It showed an

increase from 12.3 % to 24%. The 24% recovery rate obtained by the system at 60 bar OPD


was also the highest water recovery percent obtained during the experiment. Notably, the

other sucrose system involving DW as feed water also demonstrated higher water recovery

percentage in comparison to the BW-based systems. Specifically, the sucrose (var) + NaCl

(fix) / DW system showed an increase in water recovery percentage from 11.8 % to 15%

whereas the sucrose (fix) + NaCl (var) / BW showed an increase from 6.63 % to 11.7 % and

the sucrose (var) + NaCl (fix) / BW showed an increase from 6.43 % to 8% only. This higher

% R in a ternary system verses DW may due to the stronger ionisation of the sucrose and salt

(NaCl) mixture when compared to the ternary system against BW as the feed water. Also, the

differential osmotic pressure may be much lower in a ternary system using DW than using

the BW across the membrane surface resulting in a higher % R in the experiment.

Figure ‎4.29. Effect of Osmotic Pressure on Water Recovery Percent in a Ternary System

(sucrose + NaCl /H2O )

4.4.3.2. Sucrose Ternary Systems and Water Flux

Figure 4.30 presents a comparative graph of the water flux values obtained for four different

ternary sucrose systems. The water flux was a function of the osmotic pressure differential

(OPD) which was varied between 20 and 60 bar during the experiment. As is seen from the

graph, all systems in the experiment demonstrated increasing water fluxes as a result of OPD

0

5

10

15

20

25

30

17.5 20 22.5 25 27.5 30 32.5 35 37.5 40 42.5 45 47.5 50 52.5 55 57.5 60

% R

Total Osmotic pressure (bar)

Sucrose fixed with variable NaCl vs DW NaCl fixed with variable sucrose vs DW

Sucrose fixed with variable NaCl vs BW NaCl fixed with variable sucrose vs BW


increase. However, the highest recovery values throughout the experiment were obtained by

the sucrose (fix) + NaCl (var) / DW system. It showed an increase in water flux from 3.87

L/(m2.hr)to 7.3L/(m

2.hr). The 7.3 L/(m

2.hr) water flux obtained by the system at 60 bar OPD

was also the highest water flux obtained during the experiment. Notably, the other sucrose

system involving DW as a feed water also demonstrated higher water flux in comparison to

the BW-based systems. Specifically, the sucrose (var) + NaCl (fix) / DW system showed an

increase in water flux from 3.7 to 4.21 L/(m2.hr)whereas the sucrose (fix) + NaCl (var) / BW

showed an increase from 2.08 to 3.77 L/(m2.hr)and the sucrose (var) + NaCl (fix) / BW

showed an increase from 2.03 to 2.36 L/(m2.hr)only.

The increasingly higher Jw in a ternary system using DW may be due to the absence of

external concentration polarization as DW pressure is zero than the NaCl salt mixture as

compared to the ternary system using BW as the feed water. The use of different feed water

(DW and BW) might have affected the resulting osmotic potential of the mixture in a ternary

system which may be higher in the BW mixture than the DW mixture. This difference in the

osmotic potential may have caused higher water flux in the DW mixture than BW mixture.

Overall, the solute type and the ratio of the solutes in the mixture proved to be important

factors affecting the water flux in a ternary system. Overall systems involving deionised

water as a dissolving agent showed stronger water fluxes during the experiment.


Figure ‎4.30. Effect of Osmotic Pressure on Water Flux in a Ternary System (sucrose + NaCl /

H2O )

4.4.3.3. Glucose Ternary Systems and Water Recovery Percent


four different ternary glucose systems. The water recovery percent was a function of the

osmotic pressure differential (OPD) which was manually varied between 20 and 60 bar

during the experiment. As is seen from the graph, all systems demonstrated increasing water

recovery patterns with the increasing OPD levels. However, the highest recovery values

throughout the experiment were obtained by the glucose (fix) + NaCl (var) / BW system. It

showed an increase from 8.5 % to 14.5%. The 14.5% recovery rate obtained by the system at

60 bar OPD was also the highest water recovery percent obtained during the experiment.

Notably, the other glucose system involving BW as feed water also demonstrated higher

water recovery percentage in comparison to the DW-based systems. Specifically, the glucose

(var) + NaCl (fix) / BW system showed an increase in water recovery percentage from 8.15

% to 11% whereas the glucose (fix) + NaCl (var) / DW showed an increase from 4.7% to 8.6

% and the glucose (var) + NaCl (fix) / DW showed an increase from 4% to 5.3 % only. This

higher % R in a ternary system using BW may due to the stronger ionisation of the glucose

and salt (NaCl) mixture when compared to the ternary system using DW as the feed water.

0

1

2

3

4

5

6

7

8

17.5 20 22.5 25 27.5 30 32.5 35 37.5 40 42.5 45 47.5 50 52.5 55 57.5 60

J w

( L

/m

2. h

r )

Total Osmotic pressure (bar)

Sucrose fixed with variable NaCl vs DW NaCl fixed with variable sucrose vs DW

Sucrose fixed with vaiable NaCl vs BW NaCl fixed with variable sucrose vs BW


Because glucose has a lower viscosity than sucrose, higher degree of ionisation took place in

the ternary system using BW as feed water.


(glucose + NaCl / H2O )

4.4.3.4. Glucose Ternary Systems and Water Flux

Figure 4.32 presents a comparative graph of the water flux values obtained for four different

ternary glucose systems. The water flux was a function of the osmotic pressure differential

(OPD) which was manually varied between 20 and 60 bar during the experiment. As is seen

from the graph, all systems in the experiment demonstrated increasing water fluxes as a result

of OPD increase. However, the highest recovery values throughout the experiment were

obtained by the glucose (fix) + NaCl (var) / BW system. It showed an increase in water flux

from 3.0 L/(m2.hr) to 5.5 L/(m

2.hr). The 5.5 L/(m

2.hr) water flux obtained by the system at 60

bar OPD was also the highest water flux obtained during the experiment. Notably, the other

glucose system involving BW as a feed water also demonstrated higher water flux in

comparison to the DW-based systems. Specifically, the glucose (var) + NaCl (fix) / BW

system showed an increase in water flux from 2.85 to 4 L/(m2.hr)whereas the glucose (fix) +

NaCl (var) / DW showed an increase from 1.9 to 3.0 L/(m2.hr)and the glucose (var) + NaCl

(fix) / DW showed an increase from 1.69 to 1.9 L/(m2.hr)only. This higher water flux in a

ternary system using BW may due to the stronger ionisation of the glucose and salt (NaCl)

0

2

4

6

8

10

12

14

16

17.5 20 22.5 25 27.5 30 32.5 35 37.5 40 42.5 45 47.5 50 52.5 55 57.5 60

% R

Total Osmotoc pressure (Bar)

Glucose fixed with avirable NaCl vs DW NaCl fixed with virable Glucose vs DW

Glucose fixed with variable NaCl vs BW NaCl fixed with variable glucose vs BW


mixture when compared to the ternary system using DW as the feed water. Because glucose

has a lower viscosity than sucrose, higher degree of ionisation took place in the ternary

system using BW as a feed water.

Figure ‎4.32. Effect of Osmotic Pressure on Water Recovery Percent in a Ternary System

(glucose + NaCl /H2O )

4.4.3.5. Discussion

It can be seen from Figures 4.29 to 4.32 that water flux and recovery percentage increased in

the ternary systems for both OAs (sucrose and glucose) using deionized water (DW) and

brackish water (BW) as feed waters. It was noticed that the increase in the values these

parameters was considerably higher with increasing the osmotic pressure from 20 bar to 60

bar. It was also observed that the mole ratio of NaCl to any of these OAs plus DW seemed to

play a role in various fluxes due to varying osmotic pressure and ionization resulting from

addition of NaCl salt compared to the OAs (sucrose and glucose). Similar views were

expressed by many researchers who described the role of NaCl and sugar ionization on the

performance of the FO process (Narebska et al. 1987; Jiang et al. 2003; Holloway et al.

2015).But these fluxes were low when the mole ratio of OAs was wider with NaCl salt due to

less ionization. Similarly, in a ternary system using BW, the level of these fluxes is slightly

higher than a ternary system using DW. Because BW play an important role in the osmotic

pressure of a ternary system thus affecting the water flux and the recovery percentage.

0

1

2

3

4

5

6

17.5 20 22.5 25 27.5 30 32.5 35 37.5 40 42.5 45 47.5 50 52.5 55 57.5 60

Jw (

L /

m2.h

r )

Total Osmotic pressue (Bar )

Glucose fixed with variable NaCl vs DW NaCl fixed with variable glucose vs DW

Glucose fixed with variable NaCl vs BW NaCl fixed with variable Glucose vs BW


Similar findings were reported by Achilli et al. (2010) and Jeffery et al. (2006) who found

that when using seawater or brackish water as feed water both the ICP and ECP play a major

role in reducing the effective osmotic pressure difference across the membrane surface thus

affecting the different FO parameters in a ternary system. In another study, Yuan Xu et al.

(2010) stated that besides viscosity of OAs, the type of interaction between the solute and the

membrane may be another factor affecting the OAs flux performance.


5. CHAPTER FIVE: CONCLUSIONS

5.1. Introduction

The purpose of the study was to investigate the performance efficiency of various types of

membranes in the FO process using specific organic osmotic agents such as sucrose and

glucose versus deionised and salty feed water. The major aims and objectives of the study

were successfully met. Specifically, the study investigated the appropriateness of using a

number of selected organic osmotic agents (sucrose, glucose, NaCl) in the Forward Osmosis

process using different types of membranes (NF, RO, HFFM Flat Sheet) under varying

conditions. This was determined over the course of experiments where osmotic agent

concentration, flow rate and temperature were manipulated to determine the optimal

conditions and performance of each membrane-osmotic agent system under investigation.

The efficacy of the forward osmosis process was successfully measured in terms of water

flux, recovery percentage, water permeability, specific energy consumption, and solute flux.

Based on the experimental results, the best systems were identified for the FO part of a MOD

process. These systems provided acceptable water flux and good quality water.

This chapter provides the major conclusions of the conducted research based on the reviewed

literature as well as collected and analysed experimental data. The first set of conclusions is

drawn for the study of the forward osmosis process involving nano-filtration (NF) membrane

and reverse osmosis (RO) membranes with osmotic agent and feed water concentrations,

flow and temperatures. Accordingly, the conclusions with regards to efficacy of draw

solutions are based on the measurements of FO parameters such as water flux, water recovery

percentage, water permeability, energy consumption and solute flux. The second set of

conclusions is drawn for the binary and ternary systems for the experiments involving hollow

fine fibre (HFFM) membrane that uses an osmotic agent (sucrose or glucose) and feed water

(deionised water or brackish water) to determine the optimal combinations for the most

effective FO process performance. The system’s performance was measured based on water

flux, water recovery percentage, water permeability, specific energy consumption and solute

flux.


5.1. NF Membrane Study Outcomes

5.1.1. Effect of Osmotic Agent Concentration Change

An increase in osmotic agent concentration adversely affected water flux and led to a decline

in water recovery and water permeability rates with glucose solution, demonstrating higher

permeability rates. Furthermore, increasing sucrose osmotic agent concentration increased the

specific energy consumption rates, but glucose solution showed low energy consumption.

The solute flux was not affected by increasing solution concentration.

In this study, the effect was stronger for sucrose-based solution because of its higher ionic

charge, higher concentrations required and higher rate of membrane rejection. Finally, the

rate of membrane rejection for sucrose was higher than for glucose.

5.1.2. The Effect of Feed Water Flow Rate Change

An increase in feed water flow rate resulted in higher water flux, water recovery and water

permeability for both glucose and sucrose solutions. The degree of change in water recovery

rates was similar for glucose and sucrose solutions. However, the glucose solution

demonstrated better energy consumption than sucrose solution. Use of neither of the solutions

showed changes in solute flux with changing feed water flow rates.

5.1.3. The Effect of Draw Solution Flow Rate Change

An increase in draw solution concentrations for the NF membrane resulted in higher water

flux and water permeability, but showed decreases in the specific energy consumption. Use

of neither of the solutions showed changes in solute flux with changing draw solution flow

rates.

5.1.4. The Effect of Feed Water Temperature Change

An increase in feed water temperature for the NF membrane resulted in higher water flux,

water recovery percentage and water permeability, but caused a decrease in specific energy

consumption. Use of neither of the solutions showed changes in solute flux with changing

draw solution flow rates.

5.1.5. The Effect of Draw Solution Temperature Change

It was observed that increasing the temperature of the draw solutions resulted in higher water

flux and water recovery percentage but with low water permeability and specific energy


consumption. However, the use of either sucrose or glucose solution did not show any change

in solute flux with changing draw solution flow rates.

5.2. RO Membrane Study Outcomes

5.2.1. The Effect of Draw Solution Concentration Change

An increase in osmotic agent concentration showed a negative effect on water flux, water

recovery and water permeability, but demonstrated an increase in the specific energy

consumption rates. Finally, neither of the osmotic agents showed any influence on solute flux

during the experiment.

5.2.2. The Effect of Feed Water Flow Rate Change

An increase in feed water flow rate resulted in higher water flux, water recovery rates, water

permeability and energy consumption rates throughout the experiment. However, the solute

flux was not affected by using any of the osmotic agent with change in the flow rate during

the experiment.

5.2.3. The Effect of Draw Solution Flow Rate Change

An increase in osmotic agent flow rate resulted in lower water flux and water recovery rate,

but showed higher water recovery and specific energy consumption rates. Finally, neither of

the osmotic agents showed influence on solute flux during the experiment.

5.2.4. The Effect of Feed Water Temperature Change

An increase in feed water temperature caused an increase in water flux, water recovery, water

permeability and the specific energy consumption with glucose performing better than

sucrose osmotic agent throughout the experiment. Finally, neither of the osmotic agents

showed any influence on solute flux during the experiment.

5.2.5. The Effect of Draw Solution Temperature on Flow Rate Change

An increase in the temperature of the draw solution showed an increase in water flux, water

recovery and water permeability, but showed a downward trend in the specific energy


consumption for both solutions. Finally, neither of the osmotic agents showed influence on

solute flux during the experiment.

5.3. Comparative Performance of NF and RO Membranes

The following conclusions are drawn from the experiments involving NF and RO

membranes:

The NF membrane proved better for multivariate osmotic agents at lower

concentration due to its large pore size compared to RO membrane, which showed

better results at high concentration due to its compact nature. This study found

sucrose to be more suitable for use with RO membrane at higher concentrations.

In terms of water flux rates, better results were achieved by using NF membrane for

both types of osmotic agents with changing osmotic agent concentration rates. The

best result was demonstrated by the combination of NF membrane and glucose as an

osmotic agent at lower concentrations and by the combination of NF membrane and

sucrose as an osmotic agent at higher concentrations.

The highest water recovery percentage was obtained with the pair of NF membrane

and sucrose, while better water permeability rates were obtained by the combination

of RO membrane and sucrose by changing the concentration.

All the combinations of membrane-osmotic agent demonstrated increases in specific

energy consumption with increasing osmotic agent concentration rates. But the lowest

energy requirements were noted for the combination of NF membrane-glucose as an

osmotic agent.

A combination of NF-membrane and sucrose gave the best water flux with increasing

feed water flow rate. On the other hand, with the same factor change, pairing of RO

membrane and sucrose osmotic agent gave the best water recovery percentage and

water permeability.

RO membrane with sucrose as an osmotic agent demonstrated the lowest energy

consumption rates at lower feed water flow rates (up to 3.5 L/min).

While different membranes show better results at different osmotic agent flow rates,

glucose proved to be the superior osmotic agent.

A combination of RO membrane-glucose as an osmotic agent demonstrated the best

rates at lower osmotic agent flow rate (up to 3.25 L/min), while at the higher rates the


best water recovery was achieved by the combination NF membrane – sucrose as an

osmotic agent.

A combination of RO membrane and sucrose as an osmotic agent showed the best

water permeability rates at lower osmotic agent flow rate (up to 2.6 L/min).

At the osmotic agent flow rates above 2.6 L/min the best permeability rates were

observed for NF membrane and glucose as an osmotic agent.

The RO membrane with sucrose as an osmotic agent demonstrated the lowest energy

consumption rates at lower osmotic agent flow rates (up to 3.5 L/min).

With an increase in feed water temperature, the highest water flux rates were achieved

for the combination RO membrane-sucrose as an osmotic agent. The highest water

recovery rates during the experiment were achieved by the combination of NF

membrane and glucose as an osmotic agent.

The highest water permeability was observed for NF membrane and sucrose as an

osmotic agent at low feed water temperature (up to 22.5 oC).

The best permeability rates were observed for the combination RO membrane and

glucose as an osmotic agent with feed water temperature above 22.5 oC.

The RO membrane with sucrose as an osmotic agent showed lower energy

consumption rates with feed water temperature above 31 oC.

With an increase in osmotic agent’s temperature rate, the best water flux rates were

achieved by the combination of RO membrane and glucose as an osmotic agent. The

highest water recovery rates at lower temperatures of osmotic agent (up to 24 oC)

were observed for the combination of NF membrane and sucrose as an osmotic agent.

The best water recovery rates were obtained by the combination RO membrane –

glucose as an osmotic agent at temperatures above 24 oC.

The highest permeability rates were observed for the combination NF membrane –

glucose as an osmotic agent at temperature up to 27.5 oC.

The best permeability results were demonstrated by the combination RO membrane –

glucose as an osmotic agent with temperature above 27.5o C.

5.4. Outcomes for Binary and Ternary Systems for HFFM Membrane

The following conclusions are drawn from the experiments involving binary and ternary

systems for HFFM membrane:


The DW-sucrose system showed the highest water recovery rates by increasing the

osmotic pressure differential.

The DW-glucose system demonstrated the highest permeability rates increase at

pressure levels below 40 bar.

The BW-glucose system showed the best permeability rates among the considered

systems at pressure levels above 40 bar.

At higher osmotic pressure differential levels, water permeability showed

improvements.

The BW-glucose system showed the best results for in terms of specific energy

consumption

In general, all the systems tested showed increases in energy requirements with

increasing osmotic pressure differential across the membrane surface.

With respect to the solute flux, the DW-sucrose system showed the highest values

Overall, deionised water binary systems demonstrated higher FO process efficacy

across most of the considered parameters. However, energy consumption of DW

binary systems was higher, which does not make it possible to safely conclude that

they are superior.

ECP effects were the most likely source of an increase of major parameter values for

both osmotic agents, while ICP effect was the likely reason for higher water flux in

DW binary systems.

In view of all the results obtained, neither sucrose nor glucose showed superiority as

an osmotic agent in binary systems.

The binary sucrose system showed better water flux and water recovery performances

under lower osmotic pressure differential, while glucose based binary systems showed

better performance under higher osmotic pressure differentials with the border line

around 25 bar. Neither of the systems, therefore, could be concluded as being superior

with seawater used as feed water.

Water recovery rate and water flux were considerably higher in a ternary system

involving sucrose + NaCl + DW when compared to the ternary system using brackish

water as dissolving agent in FO process.

Water recovery and water flux were considerably lower in a ternary system with

glucose + NaCl / BW compared to the system when deionised water was used as feed

solution.


Overall, it appears that in a ternary mixture of glucose osmotic agent + NaCl salt +

BW, the FO efficiency is higher with respect to water recovery and water flux than

other ternary mixture with DW as dissolving agent.

The ratio of the solute and the solvent in the ternary mixture proved crucial, because

DW has no ECP. Therefore, both these parameters seem to be affected by the ratio of

the solutes in the system. Accordingly, increases in the ratio, lead to higher values for

water recovery and water flux.


6. CHAPTER SIX: FUTURE WORK

Specific recommendations for future works are listed below:

- Manipulated Osmosis Desalination should be used in future to explore the factors

affecting the process’ efficacy. Specifically, MOD can be applied to conduct

experiments manipulating various membrane parameters such as thickness, porosity

and different pore diameters. Similarly, manipulating the parameters of the existing

and new osmotic agents to uncover the best conditions for FO process is a viable

direction for future research. For example, draw solution diffusivity and viscosity

parameters could have an influence on ICP and therefore, on outcomes of the FO

process. These and other physicochemical properties should be appropriately

investigated experimentally.

- Development of new membranes that are better suited for enhancing the FO process

performance is another promising research direction.

- Further investigations should be conducted for binary and ternary systems involving

different types of membranes and solutes for refining and optimising the process of

selection.

- To avoid the obstacles of the close loop circulation system, future studies should

adopt no-circulation systems for better FO performance.

- Another important limitation of the present study was the lack of simulations and

comparisons of the experimental results against mathematically modelled data. It is,

therefore, recommended that experimental results of osmotic agents in FO are

compared with the developed mathematical models.

- Future studies should focus on identifying optimal regeneration approaches.

Adsorption by granular activated carbon (GAC) could be one possible specific

direction for research. Selection of osmotic agents based on their regeneration

economics along with other important FO parameters could be undertaken by future

research.


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http://www.sciencedirect.com/science/journal/13858947





8. Appendix‎A1:‎Osmotic‎Agents‎Samples’‎Analysis

Chromatography Laboratory

Department of Chemistry

University of Surrey

4 July 2013

Analysis of Glucose Samples

1. Samples (151 supplied) analysed by a Varian 920-LC HPLC instrument with a Varian 385-

LC light scattering detector.

2. All samples were analysed as supplied.

3. Figure 8.1 shows a calibration curve for glucose – this was used to calculate the sucrose

content in the submitted samples.

4. Table 8.1 shows the results of the analyses in mg/ml glucose.

If you have any further queries, please do not hesitate to contact us.

Figure ‎8.1. Calibration Curve for Glucose

y = 102.54x - 10.869 R² = 0.9908

-200

0

200

400

600

800

1000

1200

0 2 4 6 8 10 12

Pe

ak h

eig

ht

(mV

)

Glucose conc. (mg/ml)


Table ‎8.1. Results of the Analyses in mg/ml Glucose

Sample

Peak

height

(mV)

Glucose

conc.

(mg/ml) Sample

Peak

height

(mV)

Glucose

conc.

(mg/ml) Sample

Peak

height

(mV)

Glucose

conc.

(mg/ml)

1 167 1.7 21 1 0.1 41 586 5.8

2 147 1.5 22 543 5.4 42 2 0.1

3 1 0.1 23 484 4.8 43 512 5.1

4 285 2.9 24 2 0.1 44 486 4.8

5 212 2.2 25 589 5.9 45 2 0.1

6 1 0.1 26 526 5.2 46 619 6.1

7 188 1.9 27 1 0.1 47 596 5.9

8 116 1.2 28 487 4.9 48 3 0.1

9 1 0.1 29 477 4.8 49 510 5.1

10 348 3.5 30 2 0.1 50 471 4.7

11 491 4.9 31 513 5.1 51 2 0.1

12 1 0.1 32 514 5.1 52 522 5.2

13 361 3.6 33 1 0.1 53 453 4.5

14 529 5.3 34 483 4.8 54 2 0.1

15 1 0.1 35 461 4.6 55 488 4.9

16 478 4.8 36 2 0.1 56 510 5.1

17 530 5.3 37 481 4.8 57 1 0.1

18 1 0.1 38 478 4.8 58 673 6.7

19 488 4.9 39 3 0.1 59 532 5.3

20 492 4.9 40 508 5.1 60 3 0.1


Chromatography Laboratory

Department of Chemistry

University of Surrey

29 June 2013

Analysis of Sucrose Samples

1. Samples (65 supplied) analysed by a Varian 920-LC HPLC instrument with a Varian 385-

LC light scattering detector.

2. All samples were analysed as supplied.

3. Figure 1 shows a calibration curve for sucrose – this was used to calculate the sucrose

content in the submitted samples.

4. Table 1 shows the results of the analyses in mg/ml sucrose. Please note that “not detected”

means any sucrose present is below the limit of detection which is approximately 0.2 mg/ml.

If you have any further queries, please do not hesitate to contact us.

Figure ‎8.2. Calibration Curve for Sucrose

y = 102.38x - 17.452 R² = 0.9934

-200

0

200

400

600

800

1000

1200

0 2 4 6 8 10 12

Pe

ak h

eig

ht

(mV

)

Sucrose conc. (mg/ml)


Table ‎8.2. Results of the Analyses in mg/ml Sucrose

Sample

Peak

height

(mV)

Sucrose

conc.

(mg/ml) Sample

Peak

height

(mV)

Sucrose

conc.

(mg/ml) Sample

Peak

height

(mV)

Sucrose

conc.

(mg/ml)

1 597 6.0 11 438 4.4 21 1 0.2

2 506 5.1 12 1 0.2 22 466 4.7

3 1 0.2 13 ND ND 23 458 4.6

4 377 3.9 14 384 3.9 24 5 0.2

5 376 3.8 15 3 0.2 25 478 4.8

6 2 0.2 16 392 4.0 26 484 4.9

7 450 4.6 17 429 4.4 27 2 0.2

8 482 4.9 18 5 0.2 28 456 4.6

9 1 0.2 19 474 4.8 29 454 4.6

10 ND ND 20 444 4.5 30 6 0.2

Sample

Peak

height

(mV)

Sucrose

conc.

(mg/ml) Sample

Peak

height

(mV)

Sucrose

conc.

(mg/ml) Sample

Peak

height

(mV)

Sucrose

conc.

(mg/ml)

31 272 2.8 41 658 6.6 51 20 0.4

32 251 2.6 42 26 0.4 52 309 3.2

33 2 0.2 43 301 3.1 53 338 3.5

34 316 3.3 44 314 3.2 54 5 0.2

35 313 3.2 45 14 0.3 55 325 3.3

36 NO SAMPLE 46 317 3.3 56 306 3.2

37 452 4.6 47 311 3.2 57 3 0.2

38 417 4.2 48 9 0.3 58 338 3.5

39 4 0.2 49 334 3.4 59 344 3.5

40 694 6.9 50 322 3.3 60 22 0.4

Sample

Peak

height

(mV)

Sucrose

conc.

(mg/ml)

61 353 3.6

62 366 3.7

63 7 0.2

64 351 3.6

65 313 3.2

66 5 0.2


9. Appendix A2: Viscosity Meter TC-502

Figure ‎9.1. Viscosity Meter TC-502


10. Appendix A3: The Concentration of Mixture of Organic Osmotic Agent (Sucrose) with

Inorganic Osmotic Agent Sodium Chloride VersusDW/BW.

Table ‎10.1. The Concentration of Sucrose as OA with NaCl vs DW/BW as Feedwater

Draw solution [bar]

Osmotic Pressure

Mixture Feed Water Draw solution [g/Lw]

Volume

Draw S

Quantity chemicals

[Kg] Volume Feed W

Date Run Sucrose NaCl [atm] [bar] [ppm] Sucrose NaCl [L] Sucrose NaCl [L]

1.1 10 10 19.78 20.04 DW 0 134.200 12.730 20 2.7 0.3 80

1.2 15 10 24.73 25.06 DW 0 198.290 12.730 20 4.0 0.3 80

1.3 20 10 29.69 30.08 DW 0 260.660 12.730 20 5.2 0.3 80

1.4 30 10 39.59 40.12 DW 0 380.660 12.730 20 7.6 0.3 80

1.5 40 10 49.50 50.15 DW 0 495.100 12.730 20 9.9 0.3 80

1.6 50 10 59.40 60.19 DW 0 604.650 12.730 20 12.1 0.3 80

2.1 10 10 19.78 20.04 DW 0 134.200 12.730 20 2.7 0.3 80

2.2 10 15 24.73 25.06 DW 0 134.200 19.083 20 2.7 0.4 80

2.3 10 20 29.68 30.07 DW 0 134.200 25.380 20 2.7 0.5 80

2.4 10 30 39.58 40.11 DW 0 134.200 37.750 20 2.7 0.8 80

2.5 10 40 49.48 50.14 DW 0 134.200 49.780 20 2.7 1.0 80

2.6 10 50 59.38 60.17 DW 0 134.200 61.460 20 2.7 1.2 80

3.1 10 10 19.78 20.04 BW 5000 134.200 12.730 20 2.7 0.3 80

3.2 15 10 24.73 25.06 BW 5000 198.290 12.730 20 4.0 0.3 80

3.3 20 10 29.69 30.08 BW 5000 260.660 12.730 20 5.2 0.3 80


3.4 30 10 39.59 40.12 BW 5000 380.660 12.730 20 7.6 0.3 80

3.5 40 10 49.50 50.15 BW 5000 495.100 12.730 20 9.9 0.3 80

3.6 50 10 59.40 60.19 BW 5000 604.650 12.730 20 12.1 0.3 80

4.1 10 10 19.78 20.04 BW 5000 134.200 12.730 20 2.7 0.3 80

4.2 10 15 24.73 25.06 BW 5000 134.200 19.083 20 2.7 0.4 80

4.3 10 20 29.68 30.07 BW 5000 134.200 25.380 20 2.7 0.5 80

4.4 10 30 39.58 40.11 BW 5000 134.200 37.750 20 2.7 0.8 80

4.5 10 40 49.48 50.14 BW 5000 134.200 49.780 20 2.7 1.0 80

4.6 10 50 59.38 60.17 BW 5000 134.200 61.460 20 2.7 1.2 80

Osmotic

Pressure

OLI [atm] OLI [bar] Sucrose

10 9.8723 10.003 134.200

15 14.8044 15.001 198.290

20 19.7389 20.000 260.660

30 29.6074 30.000 380.660

40 39.4774 40.000 495.100

50 49.3453 49.999 604.650

Osmotic

Pressure OLI [atm] OLI [bar] NaCl

9.87056 10.001 12.730

10 14.7995 14.996 19.083

20 19.7333 19.995 25.380

30 29.6056 29.998 37.750

40 39.4775 40.001 49.780

50 49.3483 50.002 61.460


11. Appendix A4:Mixture of Sucrose and NaCl versus DW/BW

as Feed Water

Table ‎11.1. Sucrose and NaCl vs DW/BW as Feedwater

Draw solution [bar] Feed Water

Date Sucrose NaCl [ppm]

1

10 10 DW 0

15 10 DW 0

20 10 DW 0

30 10 DW 0

2

10 10 DW 0

10 15 DW 0

10 20 DW 0

10 30 DW 0

3

10 10 BW 5000

15 10 BW 5000

20 10 BW 5000

30 10 BW 5000

4

10 10 BW 5000

10 15 BW 5000

10 20 BW 5000

10 30 BW 5000


12. Appendix A5:Glucose and NaCl versus DW/BW as Feed

Water

Table ‎12.1. Glucose and NaCl vs DW/BW as Feedwater

Draw solution [bar]

Feed Water

Date Glucose NaCl [ppm]

1

10 10 DW 0

15 10 DW 0

20 10 DW 0

30 10 DW 0

40 10 DW

50 10 DW 0

2

10 10 DW 0

10 15 DW 0

10 20 DW 0

10 30 DW 0

10 40 DW 0

10 50 DW 0

3

10 10 BW 5000

15 10 BW 5000

20 10 BW 5000

30 10 BW 5000

40 10 BW 5000

50 10 BW 5000

4

10 10 BW 5000

10 15 BW 5000

10 20 BW 5000

10 30 BW 5000

10 40 BW 5000

10 50 BW 5000


OLI

Glucose NaCl atm bar

71.15 12.73 19.8488 20.11

105.4 12.73 24.8238 25.15

139 12.73 29.8134 30.21

203.9 12.73 39.7583 40.29

266.5 12.73 49.7363 50.40

326.65 12.73 59.6828 60.47

12.73 71.15 19.8488 20.11

19.083 71.15 24.827 25.16

25.38 71.15 29.8089 30.20

37.75 71.15 39.7738 40.30

49.78 71.15 49.7339 50.39

61.46 71.15 59.6888 60.48


13. Appendix A6: Flat Sheet Unit Experimental Work for Osmotic Agent Sucrose andRO

Membrane

Table ‎13.1. Results for The System of Flat Sheet RO Membrane and Sucrose as OA

Osmotic agent Feed solution

Concentration Osmotic Pressure Flow

rate Temperature Concentration

Osmotic

Pressure Flow rate Temperature Membrane

EXP Molar bar L/min °C Molar bar L/min °C

1 0.5M - 0.8M 12.7-22 2 25 Deionised water 0 2 25 98.35% NF TFC-SR2

2 0.5M - 0.8M 12.7 2 25 170000 ppm Sucrose 12.7 2 25 98.35% NF TFC-SR2

3 0.5M - 0.8M 15 2 25 200000 ppm Sucrose 15 2 25 98.35% NF TFC-SR2




7 0.5M - 0.8M 20 1.5 25 275000 ppm Sucrose 20 2 25 98.35% NF TFC-SR2

8 0.5M - 0.8M 20 1.5 25 275000 ppm Sucrose 20 2.5 25 98.35% NF TFC-SR2




12 0.5M - 0.8M 12.7-22 2 25 Deionised water 0 1.5 25 98.35% NF TFC-SR2

13 0.5M - 0.8M 20 2 25 275000 ppm Sucrose 20 1.5 25 98.35% NF TFC-SR2







19 0.5M - 0.8M 15 2 20 200000ppm Sucrose 15 2 25 98.35% NF TFC-SR2










14. Appendix A7: Flat Sheet Unit Experimental Work for Osmotic Agent Glucose andRO

Membrane

Table ‎14.1. Results for The System of Flat Sheet RO Membrane and Glucose as OA

Osmotic agent (OA) Feed solution (FS)

Concentration Osmotic Pressure Flow rate Temperature Concentration Osmotic Pressure Flow

rate Temperature Membrane


1 0.5M - 0.8M 12.7-22 2 25 Deionised water 0 2 25 98.35% RO 90

2 0.5M - 0.8M 12.7 2 25 90000 ppm Glucose 12.7 2 25 98.35% RO 90

3 0.5M - 0.8M 15 2 25 120000 ppm

Glucose 15 2 25 98.35% RO 90

2 0.5M - 0.8M 25 2 25 180000 ppm

Sucrose 20 2 25 98.35% RO 90

5 0.5M - 0.8M 31 2 25 250000 ppm

Glucose 22 2 25 98.35% RO 90

6 0.5M - 0.8M 12.7-22 2 25 Deionised water 0 2 25 98.35% RO 90

7 0.5M - 0.8M 20 1.5 25 250000 ppm

Glucose 20 2 25 98.35% RO 90

8 0.5M - 0.8M 20 1.5 25 250000 ppm

Glucose 20 2.5 25 98.35% RO 90

9 0.5M - 0.8M 20 1.5 25 250000 ppm

Glucose 20 3 25 98.35% RO 90

10 0.5M - 0.8M 20 1.5 25 250000 ppm

Glucose 20 3.5 15 98.35% RO 90

11 0.5M - 0.8M 20 1.5 25 250000 ppm

Glucose 20 4 35 98.35% RO 90

12 0.5M - 0.8M 12.7-22 2 25 Deionised water 0 1.5 25 98.35% RO 90

13 0.5M - 0.8M 20 2 25 250000 ppm

Glucose 20 1.5 25 98.35% RO 90

14 0.5M - 0.8M 20 2.5 25 250000 ppm

Glucose 20 1.5 25 98.35% RO 90


15 0.5M - 0.8M 20 3 25 250000 ppm

Glucose 20 1.5 25 98.35% RO 90

16 0.5M - 0.8M 20 3.5 25 250000 ppm

Glucose 20 1.5 25 98.35% RO 90

17 0.5M - 0.8M 20 4 25 250000 ppm

Glucose 20 1.5 25 98.35% RO 90


19 0.5M - 0.8M 15 2 20 250000 ppm

Glucose 15 2 25 98.35% RO 90

20 0.5M - 0.8M 15 2 25 250000 ppm

Glucose 15 2 25 98.35% RO 90

21 0.5M - 0.8M 15 2 30 250000 ppm

Glucose 15 2 25 98.35% RO 90

22 0.5M - 0.8M 15 2 35 250000 ppm

Glucose 15 2 25 98.35% RO 90


24 0.5M - 0.8M 15 2 25 250000 ppm

Glucose 15 2 20 98.35% RO 90

25 0.5M - 0.8M 15 2 25 250000 ppm

Glucose 15 2 25 98.35% RO 90

26 0.5M - 0.8M 15 2 25 250000 ppm

Glucose 15 2 30 98.35% RO 90

27 0.5M - 0.8M 15 2 25 250000 ppm

Glucose 15 2 35 98.35% RO 90


15. Appendix A8: Flat Sheet Unit Experimental Work for Osmotic Agent Sucrose and NF

Membrane

Table ‎15.1. Results for The System of NF Membrane and Sucrose as OA

Osmotic agent (OA) Feed solution(FS)

Concentration Osmotic Pressure Flow

rate Temperature Concentration

Osmotic

Pressure

Flow




2 0.5M - 0.8M 12.7 2 25 170000 ppm Sucrose 12.7 2 25 98.35% NF TFC-SR2



5 0.5M - 0.8M 31.2 2 25 400000 ppm Sucrose 22 2 25 98.35% NF TFC-SR2
















16. Appendix A9: Flat Sheet Unit Experimental Work for Osmotic Agent Glucose and NF

Membrane

Table ‎16.1. Results for The System of NF Membrane and Glucose as OA

Osmotic agent(OA) Feed solution(FS)

Concentration Osmotic

Pressure Flow rate Temperature Concentration

Osmotic

Pressure

Flow




2 0.5M - 0.8M 12.7 2 25 90000 ppm Glucose 12.7 2 25 98.35% NF TFC-SR2

3 0.5M - 0.8M 15 2 25 120000 ppm Glucose 15 2 25 98.35% NF TFC-SR2

2 0.5M - 0.8M 22 2 25 180000 ppm Glucose 20 2 25 98.35% NF TFC-SR2

5 0.5M - 0.8M 31.2 2 25 250000 ppm Glucose 22 2 25 98.35% NF TFC-SR2


7 0.5M - 0.8M 31.2 1.5 25 250000 ppm Glucose 22 2 25 98.35% NF TFC-SR2

8 0.5M - 0.8M 31.2 1.5 25 250000 ppm Glucose 22 2.5 25 98.35% NF TFC-SR2


10 0.5M - 0.8M 31.2 1.5 25 250000 ppm Glucose 22 3.5 15 98.35% NF TFC-SR2











17. Appendix A10: Sample Calculations for the Data Parameters (Table 4.1)

NF flat sheet membrane by using OAs GLUCOSE.

Table ‎16.1. Sample Calculations for the Data Parameters in Table 4.1. with Glucose OA

Draw solution concentration study Run 1 Run2 Run3 Run4

Draw solution in L/min 2 2 2 2

FW flow rate in L/min 2 2 2 2

membrane area in m2 0.0621 0.0621 0.0621 0.0621

Wt f at t=0 5.266 5.524 5.386 5.4

Wt f at t=2hr 5.24 5.507 5.375 5.392

WtDr at t=0 5.119 5.92 5.84 5.977

WtDr at t=2hr 4.684 5.485 5.463 5.787

Dr at t=0 Con 90 120 180 250

Dr at t=2hr Con 77 112.2 121.2 250

FW at t=0 Con g/L 0 0 0 0

FW at t=2 hr Con g/L 5.1 5.1 5.1 5.1

Dr at t=0 ,OP in bar 14.3 21 35 40.6

Dr att=2hr OP,in bar 11 15 20 36.5

Fw at t=0 O.P 0 0 0 0

Fw at t=2hr,O.P 0.738 0.738 0.738 0.738

NDP in bar 10.15 15.45 24.95 36

Jw in kg/hr.m2 3.5024155 3.5024155 3.0354267 1.5297907

%R FW 0.4937334 0.307748 0.2042332 0.1481481

Js g/hr.m2 2.6 2.6 2.6 2.6

Aw kg/hr.m2.bar 0.3450656 0.2 0.1216604 0.09


SEC kWhr/m3 0.3 0.4 0.5 0.9

Glucose concentration study

OA glucose at 90 g/L

( )

3.5 L(m

2.hr)


( )

3.5 L(m

2.hr)


( )

3 L(m

2.hr)


( )

1.5 L(m

2.hr)

Glucose Recovery Present % (Recovery Percentage)

Recovery percentage (% R) of the feed water was estimated using the following equation:

(

)


: Overall volume of permeate

: Initial volume of feed solution

At the end of each experiment, the recovery of the membrane was calculated by dividing the overall volume of permeate (calculate from the total

weight decrease of the feed solution) by the initial volume of feed solution.

(

)

(

)

(

)

(

)

Salt Flux

For all Glucose concentrations (90, 120, 180, 250 g/L) and Sucrose concentrations (170,200,275,400 g/L) the Js is constant at 2.6 g (m2.hr)

(

) g (m

2.hr)

Feed water concentration at time 2 hr to 10.2 g/L

feed water volumetric flow rate at time 2 h equal to 1.9 g/L


SEC Sucrose NF membrane

( )

( ) ( )

( ) kWhr/m

3

SEC for sucrose at concentration 200 g/L

( ) ( )

( ) kWhr/m

3


( ) ( )

( ) kWhr/m

3


( ) ( )

( ) kWhr/m

3

Net Driving Pressure

Glucose concentration 90 g/L

As the deionised water was used as feed water and its P is equal to zero, the NDP was calculated from the modified equation:


( )

( )

( )

( )

( )

( )

( )

( )

( )

( )

Water Permeability

=

L (m

2.hr.bar)

=

L (m

2.hr.bar)

=

L (m

2.hr.bar)

=

L (m

2.hr.bar)


NF flat sheet membrane by using OA SUCROSE:

Table ‎16.2. Sample Calculations for the Data Parameters in Table 4.1. with Sucrose OA

Draw solution concentration study Run 1 Run 2 Run 3 Run 4

Draw water flow rate 2L/min 2 2 2 2

FW flow rate in L/min 2 2 2 2

membrane arae in m2 0.0621 0.0621 0.0621 0.0621

Wt f at t=0 4.966 4.758 6.189 5.184

Wt f at t=2hr 4.84 4.663 6.109 5.143

WtDr at t=0 5.685 5.655 5.455 5.251

WtDr at t=2hr 5.557 5.557 5.393 5.285

Dr at t=0 Con 170 200 275 400

Dr at t=2hrCon 133 163.32 214 337

FW at t=0 Con g/l 0 0 0 0

FW at t=2 hrCon g/l 10.2 10.2 10.2 10.2

Dr at t=0 ,OP in bar 32.6 15 19 17

Dr at t=2hr OP,in bar 19.6 15 17 15.5

Fw at t=0 O.P 0 0 0 0

Fw at t=2hr,O.P 0.738 0.738 0.738 0.738

NDP in bar 21 9.9 12.9 11.15

Jw in L(/hr.m2) 1.0144928 0.7649 0.49919 0.27375

%R FW 2.5372533 0.78905 1.29262 0.7909

Js g (hr.m2) 2.6006441 1.99664 2.60064 2.60064

Aw in L(hr.m2.bar) 0.0483092 2.60064 0.0387 0.02455

SEC in kWhr/m3 1.2 1.3 2.2 3.3


Sucrose concentration study

( )

1 L (m

2.hr)

( )

0.8 L (m

2.hr)

( )

0.5 L (m

2.hr)

( )

0.3 L(m

2.hr)

Sucrose Recovery percentage R%

(

)

(

)

(

)

(

)


Net Driving Pressure

As the deionised water was used as feed water and its P is equal to zero, the NDP was calculated from the modified equation:

( )

( )

( )

( )

( )

( )

( )

( )

( )

( )


Water Permeability

L (m

2.hr.bar)

( )

( )

( )

SEC Sucrose NF membrane

( )

Where

PDin is the absolute osmotic pressure of the draw solution (osmotic agent) entering the membrane

QDin is the volumetric flow rate (L) of draw solution (osmotic agent) entering the membrane

PFin is the pressure (bars) of the feed water entering the membrane

QFin is the volume flow rate (L) of feed water entering the membrane and

QFoutis the volume flow rate (L) of feed water leaving the membrane during the FO process.



( ) ( )

( )

( ) ( )

( )

( ) ( )

( )

( ) ( )

( )

investigation of organic osmotic agents forward …epubs.surrey.ac.uk/808886/1/thesis _oct-...

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