A NANO FILTRATION (NF) MEMBRANE PRETREATMENT OF SWRO FEED AND MSF MAKE-UP

(PART I)

A. M. Hassan, M. AK. Al-Sofi, A. M. Farooque, A. G. I. Dalvi, A. T. M. Jamaluddin, N. M. Kither, A. S. Al-Amoudi, and I. A. R. Al-Tisan

Research and Development Center,

Saline Water Conversion Corporation P.O. Box #8328, Al-Jubail 31951, Kingdom of Saudi Arabia

For the first time ever, a new approach to seawater desalination processes was

developed at Saline Water Conversion Corporation (SWCC), R&D Center, by

integrating the nanofiltration (NF) membrane pretreatment process with one of the

conventional desalination processes to form, for example, an NF-SWRO or an NF-MSF

or a combination thereof, such as an NF-SWROreject-MSF. The process was successfully

applied to those cases on a pilot plant scale with remarkable results. The seawater

treatment first with the NF membrane removed from it turbidity and microorganism,

caused significant rejection of the scale forming hardness ions, e.g., SO4= by up to 98%,

reduced TDS in Gulf Seawater by up to 65 %, and produced a new, partially

desalinated seawater product, considerably different and superior to seawater in

qualities and without the problems normally associated with seawater of high

concentration of scale forming ions, high TDS, high turbidity and high bacteria count.

The said desalination arrangements led to a significant improvement in the seawater

desalination processes, for example by doubling the SWRO product water output and

recovery ratio and the production of high purity permeate (TDS < 200 ppm) from one

single stage SWRO. It also allowed for the successful operation at high recovery ratio

of the MSF as part of an NF-MSF unit or as part of a trihybrid NF-SWROreject-MSF

desalination system, where the reject from SWRO constituted the make-up to the MSF

unit, at a top brine temperature (TBT) of 120 oC, without the addition to the MSF make-

up of acid or antiscalant or antifoam and its operation under those conditions without

scale formation. Moreover, by this trihybrid process, up to 90% of the NF permeate was

converted to potable water. In addition to the above findings, the report describes the

* Issued as Technical report #3807/98008-I Part of the work titled A New Approach to Membrane and Thermal Seawater Desalination Processes Using Nanofiltration Membranes Part 2 was presented at 4th WSTA Conference held at Bahrain, Feb. 13-17, 1999.

effect of long term operation on the performance of those seawater hybrid desalination

processes.

As for the execution of the project, to be referred to hereafter as Part I, it was started on

22/3/1997 and after seven weeks of system modification and dual media cleaning to

remove the residual Fe3+ (from its media), the nanofiltration (NF) operation, using one

module with two NF membrane elements, was started on 10/5/1997 followed by the

installation and operation of 5 NF modules (10 elements) on 2/6/1997. The NF-SWRO,

NF-MSF and NF-SWROreject-MSF pilot plants were first operated on 3/6/1997,

21/9/1997 and 10/2/1998, respectively. So far the NF pilot plant unit has been in

operation for more than 9700 hours, while the SWRO and MSF pilot plants were

operated on NF product as feed to the former and make-up for the latter for over 5000

and 2300 hours, respectively. The hybrid system NF-SWROreject-MSF was operated for

a total of 170 hours (further operation shall be conducted at a later date). Continuous

non-interrupted operation of the NF unit without shutdown was maintained from the

start of the project until now. The execution of the project proceeded at a fast pace

ahead of schedule and all tasks were completed on time. Results were presented in a

paper to IDA World Congress, Madrid 97 [1]. A series of lectures on the same topic

were presented at SWCC. A two parts paper was published in the International IDA

D&WR Quarterly Magazine, May & September issues, 1998, on request of the editor

[2]. Two additional papers were presented at EDS Conference, Amsterdam, 1998, and

both papers were published in Desalination 118 (1998) 35-61 and 123-129 (3, 4). One

paper was accepted for presentation at the WSTA 4th Gulf Water Conference, Bahrain,

Feb. 1999. Three paper abstracts were submitted to organizers of IDA conferences on

desalination, San Diego, 1999, and the process is now in a patent pending status in

USA, Saudi Arabia and the International Patent Cooperation Treaty (PCT) Countries.

In view of the positive results obtained in this part of the R&D investigation, it is

recommended to continue work in this new field of seawater desalination. Summary of

an eight part proposed R&D work is covered under Recommendation, Section 9, Page

(26) of this report.

INTRODUCTION

With an increase in world population and rise in their living standard, there is an

increase in demand for good quality water. To meet this rise in demand, water treatment

in all its forms, is also on the rise. This tends to be the case also for water desalination

which has raised within the last two decades by over twenty folds and by the end of

1995 it stood at 20,300,000 m3/d [6]. Most of this world capacity is made of seawater

desalination by the MSF process (48.15%) and to a lesser extent by the RO desalination

process (35.9%). This latter number represents all forms of RO desalination: seawater

RO (SWRO), Brackish water RO (BWRO), industrial and membrane softening

processes.

By comparison to other forms of water desalination, seawater desalination is by far the

most complicated and complex process. It has the lowest water recovery ratio (30% to

35%). Operation is restricted to certain operation conditions. Moreover, it tends to

require extensive pretreatment, especially if the feed is taken from an open seawater

intake. The process is an energy intensive process, and for all the above factors the

seawater desalination is the most expensive among all desalination processes. The

major cause for the high expense and process complexity is the seawater itself which is

characterized by having: (1) high degree of hardness, (2) varying degrees of turbidity,

and (3) high TDS at pH of about 8.2. These seawater properties give rise to three major

problems in seawater desalination which exert severe limitation and have pronounced

effects on the performance and productivity of seawater desalination plants.

As the world leader and the major producer of desalinated water from the sea by both

the multi stage flash (MSF) and seawater reverse osmosis (SWRO) processes, SWCC

has a major interest in both the thermal and membrane technologies. SWCC has a keen

interest in solving the above three problems. For the last ten years, SWCC R&D Center

was engaged in finding solutions to those problems. First work dealt with studying the

existing SWCC SWRO plants by establishing their performance and identification of

major problems encountered in their operation [7]. This was followed by a meticulous,

systematic study of factors affecting the seawater feed pretreatment and investigation

of the various parameters affecting the conventional coagulation filtration process [8].

The

study led to the use of membranes in the pretreatment of seawater feed to SWRO plants

[9], first by using the ultrafiltration (UF) membrane process as a secondary pretreatment

to SWRO feed along with the evaluation of other microfiltration (MF) and

nanofiltration (NF) membranes for pretreatment of feed to SWRO plants. The UF

experiment proved the effectivity of this membrane process as a secondary

pretreatment. Like the MF membrane pretreatment, the UF membrane pretreatment

succeed only in the removal of turbidity and bacteria from the feed, but kept the ionic

composition of seawater the same. The NF membrane pretreatment, however, was

found to be successful in removal of turbidity, in significant removal of hardness and in

lowering of the seawater TDS [1-5]. With this pretreatment SWCC which introduced

during the late 1980s the new hybrid concept by the addition of the two 15 mgd Jeddah

SWRO plants to the existing MSF - Power facilities, introduced a new concept to

seawater desalination by combining the NF membrane process with one or more of the

conventional seawater desalination processes in one fully integrated process system. In this process, the NF product constitutes the feed to SWRO plants or the make-up to thermal,

e.g., MSF, VC, ME, desalination plants. Alternatively, in a trihybrid system,

the reject from SWRO of the NF-SWRO unit is used as the make-up to the MSF unit. This

concept was evaluated on an NF-SWRO, an NF-MSF and an NF-SWROreject -MSF pilot

plant units using Gulf seawater. From the results obtained so far together with

preliminary technoeconomic and cost analysis performed on the system, the process

was found to merit further consideration and evaluation on a demonstration plant, which

is currently ongoing, and on commercial SWRO plants as well as its application to

thermal seawater desalination processes. The report describes the long-term operation

and results obtained so far on pilot plants scale (capacity 20 m3/d), using 4" x 40"

commercial NF membranes with this new concept of NF-seawater desalination

processes, which is now in a patent pending status. 2-SOME MAJOR PROBLEMS IN SEAWATERDESALINATION &THERREMEDIES

As mentioned under Introduction, by comparison to other forms of water desalination,

seawater desalination is the most complicated, complex and expensive water

desalination process. The major cause for the high expense and process complexity is

the seawater itself which is characterized by having high degree of hardness due to

presence of Ca++, Mg++, SO4= and HCO3 ions at relatively high concentrations, high

TDS, varying degrees of turbidity, presence of micro particles, macro and

microorganisms at pH of about 8.2. For illustration, compositions of Gulf and normal

(ocean) seawater are given in Table 1. These extreme seawater characteristics give

rise to three major problems in seawater desalination, which are discussed along with

their remedies in the following sections.

Problems in Seawater Desalination

2.1.1 Problem 1: Seawater Hardness and its Effect on Seawater Desalination

The high degree of hardness constitutes a problem inherent to all forms of desalination,

be it thermal (MSF, ME, VC) or membrane type (SWRO, seawater ED). The seawater

desalination processes are separation processes in which fresh water is extracted from

saline water. This way the salts and hardness ions are left behind in the brine with the

effect that both the brine TDS and hardness concentrations are increased. Because

hardness ions are sparingly soluble in seawater, the increase in their concentration and

under certain operation conditions could lead to their precipitation on the desalination

equipment, e.g., tubes and membranes causing them to scale.

Depending on the desalination operating conditions, two types of scale could form: an

alkaline soft scale made of CaCO3 and Mg(OH)2 and a non-alkaline hard scale

consisting of CaSO4, or CaSO4.H2O or CaSO4.2H2O. The formation of non-alkaline

scale becomes exaggerated at temperature above 120 oC, since the CaSO4 solubility

decreases as the solution temperature is increased. To prevent and avoid alkaline scale

formation certain antiscalant additives are added to the feed. For example, in thermal

desalination processes such as MSF a scale control chemical is added to the make up

seawater to allow for MSF operation without scaling. There are three classes of such

scale control chemicals. These chemicals are: (i) polyphosphate (ii) polymeric type

chemicals, e.g., polyphosphonates or polycarboxylic acid or (iii) mineral acids (H2SO4

or HCl). In the same order as above they are dosed at the rate of 3-5, 1-4 or (110-120)

parts per million (ppm) for operation at top brine temperature (TBT) of 90 oC, 115 oC

and 120 oC, respectively. In spite of this, the product water recovery as a fraction of

product to make-up remains low, 30% to 35%. Operation at higher TBT requires

removal from seawater of sulfate or Ca++, which was done in the past by the ion

exchange process [10]. This way higher water recovery and more importantly at a

reduced unit production cost could be achieved. In SWRO operation, acid and in some

cases SHMP or other antiscalants are normally added to prevent membrane or SWRO

plant scaling. Again, water recovery tends to be limited, e.g., for Gulf seawater to about

35% or less.

2.1.2 Problem 2: Impurities and marine organisms and their Effect on Seawater Desalination

Another problem in seawater desalination is the impurities in seawater feed to the

desalination plants. Their presence in seawater feed could and many cases cause the

desalination system to foul. Thus, the presence of macroparticles and macroorganisms

(mussels, barnacles, algae) requires their removal from feed to both SWRO and thermal

desalination plants. In presence of bacteria in the feed, disinfection of the feed tends to

be a requirement also for both the SWRO and the thermal processes. Removal of

turbidity and fine particulates, measured as total suspended solids (TSS), from feed

destined to SWRO plants is an essential process but is not required for the thermal

processes. Although, as illustrated in this study, removal of particulates from the make-

up to MSF plant reduces foaming and thereby eliminates the addition of antifoam

chemicals. Removal of Cl2 from feed to chlorine sensitive SWRO membrane is also a

must.

2.1.3 Problem 3: Seawater TDS and its Effect on Osmotic Pressure and SWRO Desalination

A third problem in seawater desalination is the seawater feed high TDS. This

constitutes a major problem to the SWRO process but not as much to the thermal

processes. However, an increase in seawater TDS is always accompanied by an increase

in hardness which as mentioned above in Section 2.1 constitutes a problem to thermal

desalination processes i.e., by limiting distillate recovery. Per se, the TDS without or

with very low concentration of hardness ions has a minimum effect on distillate

recovery in the thermal desalination process. This is demonstrated later in the text,

section 5.3 where high degree of distillate recovery was achieved from high TDS feed

from which a significant portion of the hardness ions up to 90% were removed.

Nevertheless, increased TDS hence elevated brine recycle salinity would influence

distillation processes. The increased brine recycle salinity would give rise to solution

specific gravity (SP) and boiling point elevation (BPE). These in turn could marginally

give rise and in the same order as they appeared above to pumping power plus heat

input and distillate production due to increases in required differential temperatures and

pressures between flash zones and condensing zones.

The feed osmotic pressure increases as the ionic molar concentration and, therefore,

TDS in seawater are increased. The feed osmotic pressure f in bar is calculated from

the following equation:

f = 8.308 x 0.9 (Tf + 273.15) Mi / 100.0195 (1)

where Tf is the feed temperature and Mi is the sum of ionic concentration in moles

while the factor 100.0195 is to convert kilo pascal to bar [11]. For example f (of feed)

is 31.08 compared to 27.03, 24.94 and 14.19 bars for Gulf, Mediterranean, Open Ocean

(normal) and Baltic seawater with TDS of 43800, 38,088, 35145 and 20,000 ppm,

respectively. The average decline in f is 0.70968 bar per 1000 ppm drop in seawater

TDS. The ratios of f value for seawater from Baltic sea to that from normal sea and to

f of Gulf seawater are: 0.4566: 0.8024: 1.0, respectively.

From the principles of SWRO, the applied pressure Pappl is necessarily used to overcome

the osmotic pressure and the remaining pressure is the net water driving pressure

through the membrane (Pnet). Thus, the net driving pressure can be represented simply

as:

Pnet = Pappl - (2)

Where Pappl and are the differential applied and osmotic pressures across the

membrane, respectively. Both Pappl and are calculated from the following equation

[].

Pappl = Pf - (Pfb / 2) - Pp (3)

= fb - p (4)

Thus, Eq. 2 can be rewritten as:

Pnet = [Pf - (Pfb / 2) - Pp] - [fb - P] (5)

where the subscripts f, b and p denote the feed, brine and product, respectively.

The fb in bar is given by the following equation:

fb = [0.2654 Cfb [T + 273.15] / [1000 - Cfb / 1000] [100.0195] (6)

The concentration Cfb is the average concentration of feed and brine:

Cfb = (Cf + Cb) / 2 (7)

While Cb in terms of the recovery ratio is given as:

Cb = Cf / [1 - ] (8)

The product water quantity (Qp) or the rate of permeate passage through the membrane

is given by:

Qp = Kw (P - ) (A / ) Tc Mf = Kw Pnet (A / )Tc Mf (9)

where

- Kw is the membrane permeability coefficient

- A is the membrane surface area

- is the membrane thickness

- Tc and Mf are the correction factor effects for temperature and membrane

flux, respectively.

The salt quantity in permeate (Qs) is proportional to membrane salt permeation

coefficient (Ks) and the salt concentration differential across the membrane (C) and is

defined by the following equation:

Qs = Ks (Cfb) (A/) (10)

From Eqs. (2) And (5), the less is the osmotic pressure , the greater is the Pnet and the

greater is the amount of pressure available to drive the permeate water through the

membrane and the greater is the quantity of product. This is apparent in Eq. 9 as Qp and

Pnet are directly proportional. Likewise, the process energy is directly related to

pressure as shown later by Eq. 11. Unlike permeate flow, the salt passage through the

membrane (Qw) is independent of pressure and depends on the salt concentration

differential across the membrane. Therefore, raising the RO feed pressure increases

water flow and normally without increasing salt passage which leads to a decrease in

permeate TDS, i.e., it leads to salt dilution in permeate.

The effect of varying feed TDS on fb and Pnet in the SWRO process at an applied

pressure of 60 bars and final brine TDS of 66,615 ppm is shown in Figure 1. The

available useful pressure to drive the water though the membrane (Pnet) marked by the

shaded area increases as the feed TDS and therefore fb is decreased and vise-versa.

The fraction of the Pappl which equals to fb is considered to be a wasted energy

although it is necessary in the SWRO process. Since the permeate flow through the

membrane is directly proportional to the Pnet, therefore, any process that lowers the feed

TDS not only reduces the wasted energy but it increases the fresh water permeation (Qp)

through the membrane (Eq. 9) and decrease the permeate TDS. As shown later, this

case of increasing Pnet by lowering TDS of feed by NF pretreatment of feed is being

exploited with advantage in the combined NF-SWRO process by increasing at an

applied low pressure the permeate flow and reducing their TDS (see Section 5.2).

Remedies of Problems in Seawater Desalination

2.2.1 feed pretreatment in conventional seawater desalination plants as means of improving feed quality

The above problems in seawater desalination caused by the seawater properties and

measures used to alleviate them are summarized in Table 2 along with the quality

requirements of feed to SWRO plant and make-up feed to MSF plants when the feed is

taken from an open sea (surface) intake. In the past it was demonstrated that use of

beachwell can effectively remove turbidity and improve SWRO plant performance. The

same effect can be achieved by using ultrafiltration (UF) membranes as a secondary

pretreatment to the SWRO coagulation filtration process [9]. Use of antiscalants proved

effective in preventing scale formation but failed to increase significantly water

recovery in both the membrane and the thermal processes. Simple water filtration,

which is employed in some brackish or drinking water treatments or coagulation -

filtration processes are employed in the removal of fine particles. This process,

however, does not remove the very fine particles with sizes of less than 1 to 2

micrometer.

2.2.2 Application Of Membrane Filtration In Treatment Of Feed To Brackish And Low Salinity Water (Drinking & Waste Water)

For the removal from the feed of fine particles with sizes less than 1 micrometer,

microfiltration (MF), UF, NF and hyperfiltration/reverse osmosis membrane filtration

are employed. The MF is used for particle separation having sizes in the range of 2.0

down to about 0.08. The UF membrane process is suited for the separation of finer

particles having sizes in the range of 0.1 down to 0.02 and of molecular weight

(MW) in the range of 10,000 g/mole and above. On the other hand, the RO process,

which is based on the well-known RO desalination principle, deals with separation of

ionic size particles in the range of 0.001 or less, molecular weight 200 g/mole or less.

The NF Membrane falls in-between the RO and UF separation range, and is suited for

the separation of particle sizes in the range of 0.01 to 0.001, MW of 200 and above.

The rejection by NF, however, is based on three principles. The rejection of neutral

particles is done according to their sizes, as is the case with the MF and UF membranes.

Rejection of part of the ions and the very fine particles is done by the reverse osmosis

process. Like the RO membranes, the permeate flow through the NF membranes

increases, as illustrated in Section 5.1.2, as the Pnet is increased. Also, by a third

process, the ionic rejection of inorganic matter is achieved by their electrostatic

interaction with the negatively charged membrane [12]. The negatively charged

membrane repulses the anion increasing their rejection by the membrane. To maintain

electro-neutrality in the solution the cations are also rejected to the same degree as the

anion. For example, in certain membranes the rejection of sulfate ions is in the order of

80-90% or better with a similar rejection for the cations Mg++ and Ca++. This compares

to 10-60% rejection for the Na+ and Cl- ions. Thus, the degree of rejection by the NF

membrane is lesser for mono valent ions such as Cl- and Na+ than that for the divalent

SO4= and Ca++. The above ion selectivity allowed for the use of NF in the removal of

hardness from low salinity water. For this reason, the process is gaining acceptance as a

major water softening treatment replacing the conventional lime softening treatment.

This is most evident in the State of Florida where many of the ground waters in coastal

areas are classified as hard water. Presently the total NF plants capacity in Florida alone

is over 60 mgd and is on the rise [13]. In some of the plants both NF & RO (hybrid) are

included to allow for greater removal of dissolved salts. Moreover, use of NF serves in

the removal of hardness allowing for higher water recovery by the RO plant. Not only

the NF process removes hardness but also the product water quality was found to meet

or to exceed the drinking water standards.

Recently there is an increased interest in the application of NF membrane filtration in

brackish water and drinking water softening, removal of color, turbidy, removal of

dissolved organic which are precursors to disinfection by-products (THM) [14&15].

The water utilities justify the slightly higher cost of membrane softening over the lime

process by the superior water quality it produces especially when the feed has high color

and other impurities. To meet the various Drinking Water Acts in USA, UF and MF

membrane processes have been used also in the treatment of drinking water [16,17&18].

The NF has been used in other applications to treat salt solution and landfill Leachate

[19], demineralization of whey, removal of sulfate from seawater to be injected in off-

shore oil well reservoirs [20,21 &22] oil water separation [23], removal of natural

organic matter and precursors of disinfection by-products from a highly colored ground

water. Cross-flow microfiltration have been used in the treatment of industrial waste

water for the removal of toxic heavy metals and other suspended particles [24&25]. It

was also used to achieve 95% water recovery, which is essential in the treatment of

waste and industrial waters [26]. Use of NF filtration in treatment of various low

salinity waters is described in the literature.

2.2.3 Membrane Pretreatment Of Feed To Conventional Seawater Desalination Plants

Use of UF and MF membrane in the pretreatment of SWRO feed was reported in earlier

work [9,27&28]. The MF pretreatment, as the only primary treatment to seawater feed,

however, required membrane washing at a frequency of 30 to 40 minutes. It was

concluded that the membrane filtration pretreatment approach could be suitable as an

alternative to the conventional coagulation-filtration method, but recommended further

investigation [27& 28]. Use of UF as secondary pretreatment at SWCC, however, gave

good results reducing the SDI of feed from the primary pretreatment from 2.5 0.2 to

less than SDI of 1 for the feed receiving UF secondary pretreatment. Moreover, the

differential pressure across the SWRO membranes remained steady at less than 4 psi

during the 12,000 hours of SWRO pilot plant operation [26]. However, Use of NF

membrane as part of a combined NF-seawater desalination plant to pretreat feed to

seawater desalination plants, e.g., SWRO, MSF, etc., or as a secondary pretreatment to

remove hardness, to lower TDS, to remove turbidity, has not been reported prior to this

work. This has been tried here for the first time in the pretreatment of feed to SWRO

pilot plant and make-up to MSF pilot plant unit in the hybridization process of NF-

SWRO, NF-MSF and NF-SWROreject -MSF [1-5].

A combination of simple dual media filtration (without coagulation) with NF membrane

was found adequate not only in removal of turbidity and improvement of seawater feed

quality but also performed the important role of reducing seawater hardness and

lowering of its TDS.

Investigation of a new process utilizing NF in the pretreatment of feed to SWRO

plant and make up to MSF.

Study the effect of NF pretreatment on removal of turbidity and reduction of

seawater hardness.

Study the effect of NF pretreatment in improving the performance of : (1) SWRO

and (2) MSF (in particular reduction of hardness and antiscalant consumption)

Technoeconomical evaluation of the NF-SWRO Desalination Process.

All experimental work were done on a pilot plant scale. A schematic flow diagram of

the NF-SWRO pilot plant is given in Figure 2, while Figure 3 shows the integration

of the NF-SWRO with an MSF pilot plant distiller comprising 2 and 4 stages of heat

rejection and recovery, respectively. Moreover, this arrangement allows for utilization,

especially in winter season, of the seawater from MSF heat rejection section as feed to

the NF unit. The NF-SWRO pilot plant consists of seawater supply system, dual media

filter followed by fine sand filter, 5-micron cartridge filter, feed tank, the NF unit and

the SWRO unit. The SWRO and MSF pilot plant set-up excluding the NF unit were

described in earlier work [8&29]. The NF unit consists of the high-pressure pump and

NF modules each containing two membrane elements (size 4"x40") which were bought

about four years back. The arrangement of the modules is as shown in Figure 2 where

the feed is supplied to the first two modules arranged in parallel and the reject of each is

fed to its following module which is connected to it in series. Reject from the latter two

modules constitutes the feed for the final fifth module. The SWRO unit is made of a

high pressure pump followed by six SWRO modules, each contains one spiral wound

membrane element (size 2.5"x40"), all arranged in series as shown in the figure.

After its filtration without coagulation the filtrate was passed to the NF membrane under

pressure. This was followed by passing the product from the NF unit to the SWRO unit

where it is separated under pressure of 40 to 60 bar into product (permeate) and reject.

Alternatively, the NF permeate or SWRO reject can be fed to the MSF unit as shown in

the Figure 3.

During the entire experiment, no coagulant was added to the NF feed, which was

nonchlorinated seawater except in a very few occasions when the seawater, due to

intake maintenance, was received from intake to Al-Jubail MSF Phase-I plant. Chlorine

when present in the feed was removed prior to feed entry into the NF membranes.

During the first 2200 hours of NF pilot plant operation no acid was dosed in the feed.

Acid, however, was added to the feed, at the rate of about 25 ppm, to bring the feed pH

to 6.7 when the NF permeate recovery was raised above 50%. No antiscal or acid,

however, were added to the feed to SWRO or make-up to MSF pilot plants when they

were operated in the dual desalination system of NF-SWRO or NF-MSF or the trihybrid

desalination system of NF-SWROreject-MSF. For comparison purposes, however, acid

was added in some trials to MSF make-up.

Chemical and biological analyses were done according to latest standard analysis

methods, which are already fully established for them at RDC, chemical and biological

laboratories. These analysis were performed on a routine basis and as required, for the

seawater feed, permeate, reject from NF and SWRO units as well as for feed, make-up,

brine recycling and blow down in the MSF pilot plant.

The work was done in five steps: first by passing the feed to the NF unit and only after

establishing its performance the product from the NF unit was passed to the SWRO unit

(second step). The NF unit was operated continuously without adding acid except

occasionally and for short time to reduce the feed pH to 7.0. During the winter months,

to offset the effect of decline in NF feed temperature on NF product flow, the NF feed

pressure was raised to reach 31 bars. The SWRO unit was operated initially at pressure

from 56 bars reduced thereafter to 40 bars when the NF filtrate TDS was reduced to

about 16,000 ppm by increasing the NF feed pressure. After completing this phase, the

NF permeate from the NF desalination unit was fed to the MSF unit (third step). The

MSF pilot plant was operated at 120oC, which is the temperature operation limit of the

present MSF pilot plant, without addition of antiscalant or acid or antifoam, which are

generally used to overcome the scaling problems and to prevent product quality

degradation due to excessive foaming. The MSF pilot plant was also operated on

SWRO reject from the NF-SWRO unit in an integrated trihybrid NF-SWROreject -MSF

system (step 4). The effect of varying feed temperature & pressure on NF performance

was investigated in step 5.

5.1 NF- Trials

Figure 4 shows the reduction in hardness ions in seawater receiving NF pretreatment

utilizing FilmTec NF-70 membranes, while Figure 5 shows the percentage reduction

for the same plus other ions (Cl-, Na+) TDS and conductivity in NF permeate. Table 3

lists the concentration of the various seawater ions in Gulf seawater before and after the

NF treatment along with their percent salt rejection. When the seawater feed is passed

at a pressure of 18 bar through 5 NF modules, the average ion concentration of Ca++,

Mg++, SO4= and HCO3 was 93, 193, 206 and 46 ppm, respectively, while their average

salt rejection was to 80.7, 87.7 93.3 and. 63.3% and again in the same order. Total

hardness was reduced by 86.5%. The SO4= ion concentration, however, decreased with

operation time to about 65 ppm (rejection of 98%). By comparison at the same applied

NF feed pressure, the concentration of the hardness ions of Ca++, Mg++, SO4= and HCO3

in NF permeate when using one NF module was 63,105,55 and 37 ppm, respectively, as

compared to their concentration in seawater of: 481, 1608, 3200 and 128 ppm in the

same order. The rejections of those ions Ca++, Mg++, SO4= and HCO3 from the feed

was: 87, 92, 98 and 71%, respectively. The M-Alkalinity (of NF permeate) as CaCO3

was reduced from 45 ppm at a pressure of 18 bars without acid dosing to the feed to less

than 25 ppm with acid dosing to pH of about 6.7.

In addition to the reduction of hardness ions by the NF pretreatment, the Cl ions are

also reduced from 22,780 ppm in seawater feed to an average of abut 16,692 ppm in NF

permeate or a reduction of about 26.7%. Similar reduction is expected for the Na+ and

K+ ions. The net effect of this reduction by the NF treatment in Cl , Na+ and K+ ions

together with the reduction in hardness ions caused a reduction in TDS from 44,046

ppm in seawater to an average of 27,720 ppm for the NF pretreated feed, i.e., a

reduction of 37.3%. Due to hardness ions reduction, the pH of the feed of 8.2 is also

reduced to an average of 7.85 in the NF permeate.

Raising the NF feed pressure to 22 bar reduced the Ca++, Mg++, SO4= and HCO3-

concentration in the filtrate from the 5 modules to 50, 96, 72 and 30 ppm, respectively,

for a remarkable rejection of 89.6, 94, 97.8 and 76.6%. Total hardness was reduced by

93.3%. The Cl-, Na+ and TDS were also reduced, the latter from 44046 to 20230 for a

remarkable rejection of 54%, (Table 3). Moreover, reduction in seawater hardness and

TDS the latter value by up to 63% was achieved by raising the NF feed pressure to 31

bar.

Typical bacteria count, expressed as colony forming units/ml (CFU), was as shown in

Figure 6. The CFU number in the NF permeate of 1.3 x 103 which was observed on

26.7.97 is, unusually high and is not expected, since the average size of bacteria is in the

order of 1 m or over and is much larger than the NF membrane pore size of less than

0.01m. So, how can this relatively large size bacteria (i.e. 1m) pass through such tiny

NF membrane of pore size less than 0.01m ? Moreover, the differential pressure

across the SWRO membranes (P) remained steady and low in value at less than 1 bar,

during the entire NF-SWRO experiment indicating no biofouling. Also the NF

membrane autopsy and analysis, after 12000 hours of operation, showed no biofouling.

The bacteria count on membrane surface was within the acceptable limit, i.e., CFU

Figure 9 illustrates the NF permeate flow versus operation time, for about 9700 hours

of operation. Figure 9a shows the operation conditions, while Figure 9b, c and d

shows the actual and normalized permeate flow for feed flow, temperature and pressure.

Data in Figure 9c are normalized only for temperature effect. Figure 9a shows that

during the cold seasons, operation hours 6000 to 9000 hours, adjustment of feed

pressure allowed for raising the product flow by off-setting the decline in permeate flow

due to the seasonal lowering of feed temperature. Full normalization of the data,

however, shows that a decline in flow occurred after the cleaning of the NF membranes

by a commercial phosphate - based type detergent, DMCA-14/BIZ, made by Cheyma

Inc., Monterial, Canada. Use of various cleaning and flushing procedures failed to

improve the permeate flow to its level prior to cleaning. Use of this detergent in NF

membrane cleaning should be avoided. However, after cleaning the NF permeate flow

remained nearly steady at its level before cleaning (Figure 9b). Prior to this cleaning the

membrane retention coefficient (MRC) calculated from Qt/Qi was about 80%, where Qt

and Qi are the quantity of permeate flow at time t = 5000 hour and initial time over the

first 100 hours.

5.1.2 Effect Of Operating Conditions On NF Membrane Performance

The effect of feed applied pressure, flow and temperature at different operating

conditions on NF permeate flow, recovery and conductivity are shown by a family of

curves, one for each operating case, in Figures 10-13 Both permeate flow and recovery

are noticed to increase as either or both the feed applied pressure or temperature are

increased. On the average, this increase was about 6% and 3.4% for a rise in applied

pressure by 1 bar and for a rise in temperature by 1C, respectively. Increasing feed

flow had lesser effect on increasing permeate flow than that occurred when increasing

either feed applied pressure or temperature, but it is the recovery, which is markedly

increased with decrease in feed flow. For example, under same operating conditions of

applied pressure of 37.5 bars, the permeate recovery at the reduced feed flow rate of 20

l/min is 60% compared to only 43% when the feed flow rate was increased to 33 l/min

(Figure 10). On the other hand, permeate conductivity tends to decrease as the feed flow

rate or applied pressure or both are increased, while it tends to increase as the feed

temperature is increased (Figures 10, 12 and 13).

In the NF membrane process, permeation of the NF permeate through the membrane

occurs by two means, either by passage through the fine pores or by the RO permeation

process. In both cases the permeate flow is dependent on the applied feed pressure,

where up to the membrane flux limitation, the flow increases as the pressure increases.

This is illustrated in Figure 11, where both the permeate flow and recovery increased

while permeate TDS decreased as the applied pressure is increased. As shown in

Figure 7, more specifically, the observed increase in permeates flow and recovery is

due to the net driving pressure (Pnet). As illustrated by Eq. 9 shown also in Figure 11,

a good portion of the applied pressure is lost in overcoming the osmotic pressure ().

For example, at the applied pressure of 37.9 bars, the Pnet, the net driving pressure, is

only 12.6 bars for a loss in the applied pressure of 25.3 bars (Figure 11).

From the above trials it can be concluded that the three operating variables: feed

pressure, feed temperature and feed flow exert different influence on NF recovery and

product quality. Increasing feed pressure increases both permeate flow as well as

recovery and improves its quality. Improvement in permeates flow and recovery can be

achieved also by increasing feed temperature which leads to a moderate decline in

permeate quality. Increasing feed flow improves both permeate flow and quality but it

has a marked influence on lowering permeate recovery. For proper plant operation, by

optimizing NF permeate yield and quality, a balancing act of operating the NF plant at

best values of feed: flow, temperature and pressure are to be identified and selected.

Those operation criteria are being further investigated thoroughly for the proper

operation of large NF plants by this NF-seawater desalination process.

5.2 NF-SWRO Trials

During the early operation, from the start up to 2200 hours at the applied pressure of 60

kg/cm2 and feed temperature of 33 1 oC, the conductivity of the NF product, which

constituted the feed to the SWRO unit in the NF-SWRO hybrid system, was 41000

s/cm. The SWRO permeate flow and recovery ratio were maintained steady at 5 l/min

and about 50%, respectively. This is illustrated in Figure 14, which also shows the

SWRO permeate flow, product recovery and conductivity plotted versus operation time.

During the final phase of SWRO operation, because of a drop in NF permeate

conductivity to about 25000 s/cm, the same above SWRO permeate value and

recovery ratio were maintained at an applied pressure of only 40 bars instead of 60 bars

as shown in Figure 14. Operation at the normally applied pressure of 60 bars increased

both the permeate flow and permeate recovery each by about 40 and 25%, respectively,

Figure 14. Most important, the SWRO membranes maintained, a steady high

performance, which did not decline with operation time when operation was done at

same operating conditions.

The effect of applied pressure on SWRO permeate flow, recovery and conductivity are

illustrated in Figure 15, which for reasons of comparison shows also the same for the

conventional operation of same SWRO membranes under identical conditions but

without the NF pretreatment. Passage of the NF permeate to SWRO unit under pressure

gave satisfying results with P remaining steady and constant at 2 bars during the entire

operation. As shown in Figure 15, because of the low hardness of SWRO and TDS of

feed (see Table 3 for ions conc. at 31 bar) it is possible to obtain a recovery of up to

80% when the pressure is raised to about 70 bars. This high product recovery was

achieved with NF permeate compared to one half this value or less for normally

pretreated seawater feed. The product flow (Qp) and recovery ratio are also increased

directly with the applied pressure Figure 15. The SWRO permeate from the combined

NF-SWRO desalination system were much greater than those for the SWRO alone

when the two systems are operated at the same pressure and temperature. For example,

at 40 bars the permeate flow and recovery from the conventional SWRO are 1 l/min and

16.7%, respectively, as compared to a much higher flow and recovery ratio of 4.8 l/min

and 48% from the new process of NF-SWRO, i.e., for an increase of 480% in flow and

by 3 folds for permeate recovery. Even at the frequently employed pressure of 56 bars

(800 psi) the SWRO product flow and recovery for NF-SWRO: SWRO alone are in the

ratio of 2.43:1. Moreover, the quality of the permeate product from SWRO process is

4500 s/cm at an applied pressure 40 kg/cm2 and drops to 2300 s/cm at 60 kg/cm2 as

compared to less than 500 s/cm for SWRO permeate from the NF-SWRO process

obtained from the same membrane at the same pressure range. The SWRO membranes

used in this test, Figure 15, are old membranes and the NF pretreatment revives their

low performance. This process (NF-SWRO) is expected to extend the life of an

otherwise ready to be replaced membranes.

5.3 NF-MSF Trials

In this trial NF permeate was used as make-up to MSF at a flow rate of 1.5 m3/hour

replacing normal seawater at low concentration of scale forming ions of alkaline and

non-alkaline types in the NF permeate to the MSF pilot unit as shown in Table 3,

while Table 4 shows for comparison these concentrations in the brine recycle stream

of the MSF operated with NF or SWRO reject from an NF-SWRO unit as make-up or

seawater as make-up (conventional MSF). Also, listed in Table 4 are pH and

conductivity values of the brine recycle streams. From this table, it can be seen that

scaling potential in the MSF system have been significantly reduced and it was safe to

operate the MSF plant with NF make-up for over 2320 hours, or with reject of SWRO

from the NF-SWRO unit for 270 hours at high temperature of 120 oC without addition

of antiscalant or antifoam chemicals. At the same operating conditions, the

concentration of the scale forming ions of Ca++, and SO4= of 168 and 410 ppm in the

NF-MSF case, Table 4, and 232 and 1020 in the SWRO reject make-up case, Table

4, are low when compared to 882, and 5830 ppm in the brine recycle stream of

conventional MSF. These observation, especially the drastic reduction in SO4=, Ca++

and Mg++, are encouraging to project MSF operation at higher TBT in the range of 120 oC to 160 oC without inducing scaling [30] thus improving plant production and hence

water cost. Operation of MSF plants at higher temperature should increase the gain

output ratio (GOR) in Kgproduct/Kgsteam and the performance ratio (PR) in Kgproduct/1000

Kj, while decreasing the energy consumption in Kj/Kg product (Figure 16). Finally, the

MSF plant operation on NF make-up remained steady with operation time.

The result obtained so far from the NF-SWRO pilot plant were encouraging to consider

the application of the NF membrane process first in a demonstration NF-SWRO plant,

which is being done now, to be followed by its actual application to an existing

conventionally operated SWRO plant, which is now under consideration. For the

present report the performance of SWCC Jeddah SWRO plant with and without NF

unit was conceptually evaluated using the results obtained from the pilot plant studies

described above [1-3].

When NF feed pressure was set to 31 bar the hardness ion concentration in the NF

permeate, which constitutes the feed to SWRO plant was as shown in Table 3. At an

applied pressure of 40 bar and SWRO product recovery of 50%, the reject from the

SWRO contains low concentration of hardness ions of 96, 253, 410 and 42 ppm for

Ca++, Mg++, SO4= and HCO3-, respectively. The TDS of reject brine of 30,640 ppm in

the same Table 3 is also low when compared for example to the reject from Jeddah

SWRO plant of about 66,615 ppm at the applied pressure of about 60 to 65 bar

[31&32]. This suggests that higher recovery of more than 35% can be achieved from

the Jeddah SWRO plant if it is modified to operate with an NF pretreatment in a

combined NF-SWRO system. This is illustrated in Figure 17 which is a schematic

flow diagram of the desalination part of Jeddah SWRO plant with and without NF unit.

Figure 17a represents the actual Jeddah SWRO plant feed, product and reject flows

along with the product water recovery percentage (%), the brine flow/modules and the

energy required for the desalination part alone. Energy was calculated from the

equation [33]:

Energy (KWH/m3) = [Qf . Hf / 366 Qpe] (11)

where:

- Qf and Qp are the quantity of feed and product in m3/hr, respectively,

- H is the pressure head in (m),

- density of seawater (1.03), and

- e pump efficiency ( 0.85).

Figure 17b, and c simulates the results of operation of Jeddah SWRO in a combined

NF-SWRO system utilizing the present SWRO desalination set-up as it is now (case b)

and with reject staging (case c), respectively. The Jeddah SWRO plant receives feed,

with TDS of 43,300 ppm from a conventional coagulation filtration unit at the rate of

6760 m3/hr and produces from 1480 modules at an applied pressure of about 60-65 bar

2370 m3/hr of fresh water for a product recovery of 35%. The total quantity of reject is

4390 m3/hr with TDS of about 66,678 ppm. The product and reject flow per

module/hour is 1.6 and 2.97 m3/hr, respectively. The energy requirement for the SWRO

desalination part alone is 6.14 kwh/m3 of product, assuming plant operation at 60 bars

and rises to 6.65 kwh/m3 at an operational pressure of 65 bar.

Each of the hollow fine fiber membrane modules used at Jeddah contains two SWRO

membrane elements arranged in series with brine staging where the feed is first passed

to the first set of elements (1st desalination step) and the remaining feed after extraction

of a fraction of it as product is passed as the SWRO reject to the second set of elements

constituting the 2nd desalination step which in turn extracts a second fraction of product

(Figure 17). The ratio of product quantities extracted by the first set of elements and the

second set of elements can be computed through Eq. 9 and the following ratio equation:

Q setQ set

K P A T MK P A T M

pst

pnd

w appl fb c f

w appl fb c f

=

12

1

2

( ) ( / )( ) /

=

( )( )

PP

PP

appl fb

appl fb

net

net

=

1

2

1

2 (12)

where Pnet-1 and Pnet-2 are the net driving pressure force for all the first set and second

set of elements (1 and 2), respectively. Also as demonstrated in Figure 15, the product

flow (Qp) should increase as Pnet is increased.

To establish the value of Pnet-1 and Pnet-2 the recovery ratio for the first set of elements is

to be established. For first approximation the recovery of first set of elements is

estimated at 20% making the second set recovery equals 15% of the 6760 m3/hr feed for

a total system recovery of 35%. With this first approximation of recovery for the first

set of elements the water product recovery ratio of 1st set: 2nd set of elements is 0.5705:

0.4295 as compared to different ratio of Pnet-1 : Pnet-2 of 0.6175 : 0.3825. The Pnet values

are calculated from Eqs. 5 and 6. Regression analysis lead to the values shown in

Figure 13, where the product water ratio of 1st set: 2nd set of elements is 0.62: 0.38 or

in the ratio of 1:0.61 which is not only identical to the same ratio of Pnet-1 : Pnet-2 but also

the same, with some deviation, as the observed ratio values which were established

experimentally at our pilot plant by using the same SWRO membrane as used at Jeddah

SWRO desalination plant. The product recovery ratio of 1st set: 2nd set of elements

when compared to their feed is 21.75%: 17.01% and with product ratio of 1470: 900

m3/hr, respectively. But the recovery (%) of each of 1st set: 2nd set of elements

compared to feed of 6760 m3/hr is 21.75%: 13.31% for a total product recovery of

35.059%.

The same data treatment used in case "a" for Jeddah SWRO was employed in

establishing the potential performance of Jeddah SWRO when combined with NF (NF-

SWRO), cases "b" and "c" in Figure 17. Again, the product flow, recovery and energy

were as shown in Figure 17b, and c. As in the previous case "a" the SWRO

desalination is assumed to occur in two steps. In case "b" of Figure 17, the first set of

elements, step 1, is assumed to treat the NF-product 27,300 ppm to yield reject with

comparable salinity (TDS) as that for the actual feed to Jeddah SWRO plant, which

constitutes the feed to the second stage elements. The second stage set of elements

allows for extraction of product from this feed to yield brine with TDS of 68258 ppm.

This arrangement yields 2724 m3/hr of product at a recovery of about 41% compared to

a product flow of 1282 m3/hr and recovery of 32% of the feed to the second step. The

overall recovery is 60%. But in this latter case the brine flow per module of 1.83 m3/hr

will be less than the minimum brine flow requirement of 2.0 m3/hr per module. This

could be tolerated since the hardness content of the SWRO reject is very low (Table 3).

However, SWRO feed has an SDI

turn is expected to improve the yield, recovery and product water quality of Jeddah

SWRO plant when combined with an NF pretreatment.

The above results obtained with the NF pretreatment of seawater feed in the removal of

hardness, lowering of TDS, pH when added to the improvement of NF permeate as feed

to SWRO and the observed gain in SWRO product water recovery ratio are remarkable

ones and should allow for the overcoming of the three problems in seawater

desalination described earlier, Section 2, which are due to the composition of seawater

and what it contains such macro and micro particles and organisms. Moreover, because

of the high purity of product from the SWRO unit (TDS 200 ppm) the NF

pretreatment should make it quite easy to produce fresh water from the sea by the

SWRO process in one single stage SWRO, thus eliminating the second brackish RO

stage with savings in each of capital investment and O&M cost by over 10%. This is in

addition to increasing the plant output at least by 15% since the elimination of the

second stage allows for the recovery of all products from the first stage SWRO unit.

The above treatment of adding the NF to SWRO plant can be extended to NF addition

to MSF plants. Earlier work showed that operation of MSF plant at 135 oC and 150 oC

without scale formation was possible when the sulfate in the feed (Mediterranean sea)

was reduced, using ion exchange, from 2900 to 1200 ppm [10]. With the NF

pretreatment the sulfate ions in Gulf seawater is reduced from 3200 ppm to less than

206 ppm when using pressure to the NF feed of 18 bars and to only 72 ppmwhen using

a pressure of 21 bars. Further reduction in level of sulfate to less than above values is

expected from the NF treatment of seawater in other seas, e.g., Ocean, Mediterranean.

This should allow for the use of the NF permeate as make-up to MSF plant in a

combined NF product - MSF unit permitting the operation at TBT of 120 oC to 150 oC or

higher with an increased gain in distillate output. Furthermore, the sulfate content in

SWRO reject from the NF-SWRO pilot plant is less than 420 ppm and Ca++ less than

100 ppm when their concentration in the NF product is 230 and 52 ppm, respectively.

Lower values are expected when theCa++ and SO4= levels in the NF product are less

than the above values. Again, this should allow for its use as make-up to MSFplant in

a hybrid NF-SWROreject/MSF system. This last desalination system arrangement should

allow, as already demonstrated in this study, for the recovery of up to 90% of the NF

product as potable fresh water, where 60 to 70% of the NF product is converted to

potable water by the SWRO unit and with the remaining 20 to 24% derived from the

conversion of 80% of the SWRO reject also into potable (distillate) water by the MSF

Unit. Work is in progress on further evaluation of the two concepts of NFproduct-MSF

and NF-SWROreject-MSF system by operation at TBT of 120 oC. Future work will

explore TBT elevation to as high as 160 oC in the MSF pilot plant at SWCC RDC, Al-

Jubail. Results of work in progress as well as future work are to be described in

separate reports.

(Cost of Water from Conventional SWRO and NF-SWRO Processes)

The cost of water production in SR/m3 is calculated for Jeddah 1, Jeddah 2 and Yanbu

SWRO plants with and without NF pretreatment unit. In all cases, the production for

each SWRO plant when operated alone without NF pretreatment is set to the actual

Jeddah 1 and 2 SWRO plants capacities at 56880 m3/d, while the installed SWRO plant

cost is adjusted to reflect this limitation in plant size for the Yanbu SWRO plant. The

power and chemical consumption, the costs of spare parts, membrane replacement,

micron cartridge filters, other consumables and O&M including labor as well as plant

availability of 90% are set to the actual values established for Jeddah -1 SWRO plant

[34]. However, no coagulant is used in the NF-SWRO process and the H2SO4 is added

at the reduced rate of about 30 ppm to lower feed pH to about 7 in order to allow for

product recovery 60% from the NF unit without alkaline scaling. For the combined

NF-SWRO case 'b' in Figure 13 the energy is calculated from Eq. 11 based on 60%

recovery of both the NF and SWRO units. The total energy E(NF-SWRO) kwh/m3 was

computed from the equation:

ET(NF-SWRO) = Ef + Eb1 + ENF + Eb2 + ESWRO + Ep+ Eothers (13)

Where E is the electrical energy (kwh/m3) delivered by the various pumps. The letters f,

b, and p denote the feed, booster, and product, respectively, ENF and ESWRO are the

energy delivered by the high-pressure pump(s) to the desalination part of the NF and

SWRO units, respectively. The figure 8.35 kwh/m3 is the reported actual energy for the

total SWRO process including intake, pretreatment and post treatment steps in Jeddah

SWRO plant by [34], i.e., equals the sum of energies as given by Eq. 13. The energy

required for the desalination parts of the NF & SWRO process, i.e., ENF and ESWRO, in

Eq. 13 computed from Eq. 11 at recovery ratio of 60% for each process is 1.655 and

3.31 kwh/m3, respectively, for a total of 4.97 kwh/m3. This value compares to 6.14

kwh/m3 for the conventional SWRO process or in the ratio of the latter to the former

case of 1.0: 0.81. As established from Eq. 13 the total energy for the NF-SWRO

process is 6.28 kwh/m3 and the total energy ratio of ESWRO: ENF-SWRO is 8.35: 6.28

kwh/m3 or in the ratio of 1.0: 0.75. Thus the non-desalination part of the total process

energy in Eq. 13, i.e., Ef + Eb1 + Eb2 + Ep + Eothers, for the conventional ESWRO : ENF-

SWRO is 2.21: 1.32 kwh/m3 or 1.0: 0.597 and product ratio for ESWRO: ESWRO is /.32 :

2.21 kwh /m3 or 0.584 : 1.0 which explains the reduction in energy of the NF-SWRO to

that of the conventional SWRO process (see Eq. 11). The reduction in energy for the

NF-SWRO as compared to conventional SWRO is due to the expected increase in plant

productivity.

Operation of the combined NF-SWRO requires the following modifications: (1) the

addition of NF unit to the existing SWRO plant and (2) an increase in feed quantity by

introducing additional feed line with its necessary electrical, mechanical and civil

works. The latter works were computed from actual prices at about 18% of total

contract value for the Jeddah-2 SWRO plant and Yanbu plant and at about 25% for the

Jeddah-1 SWRO plant. As for the cost of the NF unit it was estimated from a quotation

for its desalination part alone (high pressure pump, NF modules, pipe and control)

without other auxiliary equipment.

Table 5 lists the cost in SR/m3 product for six cases: Jeddah-1, Jeddah-2 and Yanbu

SWRO plants with and without NF pretreatment unit for two energy prices. In case "a"

the cost of power was set at SR 0.05/ kwh as was done in reference [34] and in case "b"

the power cost was increased to SR 0.375 ($0.1)/ kwh. The annual fixed charge rate

was assumed at 10% per year of the installed cost. The interest on investment is zero.

This case is assumed to equal plant depreciation over 20 years with 7% annual interest.

In all cases, the cost of product from the NF-SWRO plants was lower than that when the

conventional SWRO plant is operated alone without the NF pretreatment. A drop in

cost of product by about 28% can be realized when the plant is operated with NF

pretreatment. The product water ratio of SWRO: NF-SWRO is 18,685,080: 31,977,504

m3/yr or 1: 1.71. The product cost SR/m3 is in the ratio of 1: 0.72 for SWRO: NF-

SWRO for Jeddah-2 and Yanbu SWRO plants. More details of the results are shown in

Table 5. The techno-economic results are in conformity with results of the simulated

model in that the operation of the SWRO plants with NF unit is superior to SWRO plant

operation alone without the NF pretreatment. Moreover, improvement in plant

productivity accompanied with reduction in water cost as well as reduction in energy

consumption are expected when NF feed pressure and NF and SWRO recovery ratios

are raised.

1. The NF membrane treatment of noncoagulated dual media filtered seawater feed

to desalination plants removes from it (1) very fine turbidity, (2) residual

bacteria, (3) scale forming hardness ions, in some case by up to 98% and (4)

lowers its TDS, depending on operation conditions by 35 to 60%.

2. With this NF feed treatment the otherwise complex conventional seawater

desalination process, are simplified since the effects on seawater desalination by

the above four factors, which constitute the major problems in seawater

desalination by the conventional processes, are eliminated.

3. Feeding this low turbidity new NF product to SWRO membrane resulted in

remarkable increase in membrane product (permeate) quantity and improved its

quality. Recovery ratio increased by over 100%.

4. With this new NF SWRO process addition of a second stage (brackish RO) is

not required.

5. Use of this new NF product as MSF make-up increased distillate output up to

80%. Operation at the same condition of MSF with make-up made of seawater

was not possible due to increased potential of non-alkaline scaling.

6. With this new NF source of make-up to the MSF process it becomes feasible to

operate at TBT above 120 oC due to its low potential of non-alkaline scaling.

Table 1. Composition of Gulf Seawater, Al-Jubail and Normal Seawater

Constituents Gulf Seawater, Al-Jubail

Normal Seawater

Cataions (ppm) Sodium, Na+ Potassium, K+

13440 483

10780 388

Calcium, Ca2+ * Magnesium, Mg2+ *

508 1618

408 1297

Copper, Cu2+ Iron, Fe3+ Stronsium, Sr2+ Boron, B3+

0.004 0.008

1 3

-- -- 1 --

Anions (ppm) Chloride, Cl-

24090

19360 Sulfate, SO4= * Bircarbonate, HCO3- *

3384 176

2702 143

Carbonate, CO3= Bromide, Br- Flurried, F- Silica, SiO2

--- 83 1

0.09

-- 66 1.3 --

Other Parameters Conductivity (S) H

Dissloved Oxygen (ppm) CO2 (ppm) Total Suspended Solids (ppm)

62800 8.1 7

2.1 20

-- 8.1 6.6 2 --

Total Dissolved Solids (ppm)

43800 35146

* Hardness scale forming ions

Table 2. Pretreatment and Quality Requirements of Feed Taken from an Open Sea (Surface) Intake

Problems in Seawater Desalination Due to

Seawater Characteristics

Pretreatment and Quality Requirement of Feed to

SWRO Thermal High Degree of Hardness (Ca++, Mg++, SO42-, HCO3-)

Requires Removal or Inhibition of precipitation

by addition of antiscalant, and by

Operation at correct conditions

Requires: Removal or Inhibition of precipitation by

adding antiscalant Operation at correct conditions

High TDS Requires lowering of TDS which in turn Lowers Waste due to Increases Recovery Ratio Lowers Energy /m3 Lowers Cost /m3

Lowering of TDS beneficial by reducing concentration of hardness ions

High Turbidity (TSS, Bacteria, etc.).

Requires complete removal Requires Partial Removal Complete removal of

turbidity, however, reduces foaming and, therefore, eliminates need for addition to make-up of antifoam

Table 3. Chemical Composition and Physical Properties of Seawater, NF Filtrate, and NF and SWRO rejects at Different NF Feed pressure

Element / Parameter Seawater NF Filtrate (5 modules)*

SWRO Reject

NF Reject

Ion. Conc. Ion Conc. Rejection (%) Ion Con. Rejection (%) Ion. Con. Rejected % Ion Con. Ion Con. NF Feed Pressure (Bar) 18 18 22 22 31 31 31 31 A. Hardness Ca++ (ppm)

481

92

80.9

50

89.6

52

89.2

96

701

Mg++ (ppm) 1608 192 88% 96 94.0 143 91.1 253 2200 Total Hardness (ppm) 7800 1014 87 520 93.3 720 90.8 1280 10800 SO4-- (ppm) 3200 206 93.3 72 97.8 230 92.8 414 4950 HCO3- (ppm) 128 46 63.3 30 76.6 24 81.3 42 133 B. Other Ions Cl- (ppm)

22780

16,692

26.7

12320

46.3

9640

57.7

19570

29350

Na+ (ppm) (12860) (9426) 26.7 (6904) 46.3 (5442) 57.7 C. Total Dissolved Solids TDS (ppm)

44046

27,720

37.3

20230

54.1

16400

62.8

30,640

63640

pH 8.2 7.85 7.92 6.38 7.08 7.46 Conductivity (s/cm) 60,000 40,470 31100 24600 43300 68600 Module arrangement 2:2:1 , each module contains 2 NF elements, arrangement equals two parallel lines each having 5 elements in series.

Table 4. Operation and Performance Parameters of NF-MSF and NF-SWROreject-MSF Desalination Hybrid Systems Vs Conventional Seawater MSF Desalination System (Operation of MSF unit with NF product or SWRO Reject from NF- SWRO unit was done without addition to make-up of antiscalant or antifoam)

Trials Brine Recycle Data* Performance Parameters TBT

(oC) Ca++

(ppm) SO4=

(ppm) M- alkalinity

(ppm) Conductivity

(s/cm) pH Make-up

(m3/hr) Product (m3/hr)

Recovery ratio (%)

A. NF-MSF With acid 120 160 390 26 59000 8.19 1.5 0.97 65

Without acid** 120 168 410 65 62000 8.63 1.5 0.97 65 B. NF-SWRO reject-MSF ***

Without acid 120 232 1020 72 87400 8.50 1.5 0.97 65 C. Seawater MSF

With acid 120 882 5830 30 92000 7.99 1.5 0.97 65 With acid 90 661 4460 20 74000 7.57 1.5 0.59 39 With acid 120 561 3330 14 61000 7.50 4.0 0.92 23 With acid 90 581 4000 18 66000 7.51 2.1 0.94 45

* Brine recycle flow rate was maintained between 6.5 to 6.8 m3/h. ** Make-up is NF product having Ca++ 52, Mg++ 143, SO4--230, M-Alkalinity 24 and TDS 16400 ppm *** Make-up is SWRO reject from NF-SWRO unit with composition : Ca++96, Mg++ 253, SO4-- 414, M-Alkalinity 42 and TDS 30640 ppm

Table 5. Cost (SR) of Product Water from SWRO Plants with and Without NF Pretreatment (a). Cost of Power SR. 0.05 KWH

Cost Component Jeddah -1 SWRO

Jeddah -2 SWRO

Yanbu SWRO

Jeddah -1 NF-SWRO

Jeddah-2 NF-SWRO

Yanbu NF-SWRO

Annual fixed charge rate 16,545,566 36,293,785 37,938,113 16,545,566 36,293,785 37,938,113 Cost of NF unit - - - 5,158,188 5,158,188 5,158,188 Cost of additional Feed - - - 4,376,353 4,376,353 4,376,353 Chemicals 9,32,940 9,32,940 9,32,940 574,500 574,500 574,500 Electrical 7,798,599 7,798,599 7,798,599 9,278,700 9,278,700 9,278,700 Spare parts 2,249,100 2,249,100 2,249,100 2,500,000 2,500,000 2,500,000 Membrane replacement 2,330,835 2,330,835 2,330,835 3,000,000 3,000,000 3,000,000 Micro cartridge filter 316,800 316,800 316,800 400,000 400,000 400,000 O&M - labor 3,500,000 3,500,000 3,500,000 3,600,000 3,600,000 3,600,000

Total Cost (SR) 33,673,840 53,422,659 55,066,387 45,433,307 65,181,526 66,825,854

Product cost (SQ/m3) 1.786 2.833 2.92 1.421 2.038 2.089 Ratio SWRO : NF-SWRO 1 1 1 0.791 0.719 0.716

(b). Cost of Power SR 0.3754 / KWH ($ 0.1 KWH)

Cost Component Jeddah -1 SWRO

Jeddah -2 SWRO

Yanbu SWRO

Jeddah -1 NF-SWRO

Jeddah-2 NF-SWRO

Yanbu NF-SWRO

Total cost without electric 25,875,241 45,623,460 47,267,788 36,154,607 55,902,826 57,547,154 Energy cost 58,551,881 58,551,881 58,551,881 69,664,480 69,664,480 69,664,480 Total cost 84,427,122 104,175,341 105,819,669 105,819,087 125,567,306 127,211,634

Cost (SR/m3) 4.478 5.525 5.612 3.309 3.928 3.978

Ratio 1 1 1 0.739 0.711 0.709

Total Product (m3/d) 18,685,080 18,685,080 18,685,080 31,977,504 31,977,504 31,977,504

0

10

20

30

40

50

60

70

80

20000 25000 30000 35000 40000 45000TDS (PPM)

Fee

d-b

rin

e O

smo

tic

Pre

ssu

re (

bar

)

Seawater

M S FTO MSF

HPP

SWRO SECTIONSWRO REJECT TO MSF

PERMEATE

NF REJECT

ROFEED

NF PRODUCT TO RO FEED TANK

CF FEEDTANK

NF SECTION

M S F

Dualmedia Fine Sand

Pump

Fig2Schematic Flow Diagram of NF-SWRO Desalination Pilot Plant

Seawater

CF FEEDTANK

Duel Media Fine Sand Media

HIGH PRESSUREPUMP

R e j e c t t o M S F

PERMEATE

SWRO UNIT

A/A

MSF UNIT

4 H . R . C S t a g e s

B r i n e h e a t e r

BOOSTER PUMPFEED TANK

H. RJ

SW

D

B.B

B.R

N F U n i t

PRODUCT

N F R E J E C T

Pump

S e a w a t e r f r o m M S F H e a t R e j e c t i o n S e c t i o nS W R O U n i t

Fig3Schematic Flow Diagram of NF, SWRO and MSF Pilot Plant

Fig 1 Effect of Seawater Feed TDS on Osmotic Pressure of Feed-brine (fb) Keeping SWRO Brine Concentration at 66615 ppm

Pnet= Pappl-

Pappl

Ca++ Mg++ SO4-- HCO3-- Total Hardness

48192

1608

192

3200

206 128 128

7800

1014

0

1000

2000

3000

4000

5000

6000

7000

8000

Con

cent

ratio

n (P

PM)

Ca++ Mg++ SO4-- HCO3-- Total Hardness

SeawaterNF Filtrate (AV )

Fig4 Effect of New Process on Removal of Hardness Ions (Ca++, Mg++, SO4--, HCO3--) From Gulf Seawater (Using 5 Modules, 10 NF Elements)

Ca+

+

Mg+

+

SO4-

-

HC

O3-

-

Tot

al H

ardn

ess

Cl-

Na+

TD

S

Con

duct

ivity

8188

93

63

86

27 27

3732

0

10

20

30

40

50

60

70

80

90

100

Rej

ectio

n %

Ca+

+

Mg+

+

SO4-

-

HC

O3-

-

Tot

al H

ardn

ess

Cl-

Na+

TD

S

Con

duct

ivity

Fig 5 NF Percent Rejection of Ions, TDS and Conductivity

26 / 07/ 97 12 / 10 / 97

0

1

2

3

4

5

thou

sand

BDMFADMFAMCFNFP

Fig 6 Bacterial Count Colony Forming Unit /ml (CFU) from Different Sampling Points in NF Pilot Plant at 0 hr.

CFU

BD

MF

AD

MF

AM

CF

NF

B

NF

P

MAY

JUN.

JUL.

AUG.

SEPT.

OCT.

Nov

DEC

1.00E+01

5.01E+03

1.00E+04

1.50E+04

2.00E+04

Mon BDMF ADMF AMC F NFB NFPMAY 1.0E+3 2 .8E+3 2.8E+3 6.2E+ 3 6.2E+1JUN 2.6E+3 2.4E+3 2.1E+3 4 .3E+3 2.1E+2JUL 2.9E+3 8.5E+3 1.4E+3 2.4E+3 2.1E+2 AUG 1.2E+4 7.9E+3 7.9E+3 1 .06E+4 3 .1E+2S EP 1.7E+4 3.2E+3 7.2E+3 1.9E+3 1.5E+2O C T 1.1E+2 1.9E+2 6.0E+1 6.0E+2 3 .0E+1 NO V 1.2E+3 9 .5E+1 1.0E+2 8.5E+2 7.1E+1DEC 2.7E+3 9.5E+1 7.5E+1 1.2E+3 5.7E+1

CFU

Fig 7 M onthly AVG Bacteria Count in Colony Forming Unit /ml (CFU) from Different Sampling Points in NF Pilot Plant at 0 hrs (BDM F& ADM F are Before and After Dual M edia Filter respectively, AM CF is After M icron Cartridge Filter, , NFB& NFP are Nanofiltration Brine and Permeate)

0

10

20

30

40

0 200 400 600 800 1000 1200

Feed

Flow(l/mi

Temp (

0C)

feed flow temperature(oC) pressure (bar)

y = 19.383x-

05

1015202530

0 200 400 600 800 1000 1200

Flow (l/

iCleaning

05

1015202530

0 200 400 600 800 1000 1200

Flow (l/

iCleaning

and Feed Flow

y = 39.456x-

05

1015202530

0 200 400 600 800 1000 1200

Flow (l/

i

20.7 24.1 27.6 31.0 34.5 37.9

y = 3.5199xR2 = 0.9987

y = 1.0519xR2 = 0.9989

0

2

4

6

8

10

12

14

16

18

20

0.0 2.0 4.0 6.0 8.0 10.0 12.0 14.0 16.0 18.0Pnet (bar)

Flo

w (

l/min

.)

0

5

10

15

20

25

30

35

40

45

50

Rec

ove

ry (%

)

Pf = Pappl. (bar)

Pnet = [Pf - (Pfb/2) - Pp] - [fb - p]

f = Feed; b = Brine; p = Product

Recovery

Flow

80 100 120 140 160

Q F = 11266 m3/H

Q F = 6760 m3/h

2774 + 1282 = 4056m 3/h

Q f = 6760 m3/H

1470 + 900 = 2370 m 3/h

Q f = 11266 m3/H

3057 + 933 = 3990m 3/h

=C f 43300ppm

C f 43300= ppm

C f 43300= ppm

66600 ppm

2771 m 3/h

68250 ppm

2704 m 3/h

66678 ppm

4390 m 3/h

a. Jeddah SW R O (A ctual)

b. N F-Jeddah SW R O

C F = 27300 ppm 3986 m 3/h

46300 m 3/hSWRO SWRO SWRO

3.33

SW R O SW R O SW R ON F

Q F = 6760 m3/h

C F = 27300 ppmC f = 67300 ppm

C f = 67300 ppm

c. N F-Jeddah SW R O W ith Brine Staging

SW RO unit

SW RO unit

N F

SW RO unit

1st Stage2nd Stage

IDA

Desalination and Water Reuse Quarterly

EDS Conference

Desalination

Water Science and Technology Association

IDA World Congress

IDA World Congress

IDA World

Congress

Desalination

ASTM Desalination, Desalination, Desalination

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Desalination Water Supply Aqua, Desalination Desalination Filtration and Separation The 1995 Thirteen Membrane Technology Conference

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