analysis and assessment of advanced hydrogen …

ANALYSIS AND ASSESSMENT OF ADVANCED HYDROGEN

LIQUEFACTION SYSTEMS

by

Anwar Hammad

A Thesis Submitted in Partial Fulfillment

of the Requirement for Doctor of Philosophy Degree

in

Mechanical Engineering

Faculty of Engineering and Applied Science

University of Ontario Institute of Technology

Oshawa, Ontario, Canada

September 2019

© Anwar Hammad, 2019

i

THESIS EXAMINATION INFORMATION

Submitted by: Anwar Hammad

Ph.D. in Mechanical Engineering

Thesis title: ANALYSIS AND ASSESSMENT OF ADVANCED HYDROGEN LIQUEFACTION

SYSTEMS

An oral defense of this thesis took place on August 8, 2019 in front of the following examining committee:

Examining Committee:

Chair of Examining Committee Dr. Sayyed Hosseini

Research Supervisor Dr. Ibrahim Dincer

Examining Committee Member Dr. Martin Agelin-Chaab

Examining Committee Member

University Examiner

Dr. Haoxiang Lang

Dr. Mustafa El-Gindy

Eternal Examiner Dr. Ziad Sagir

The above committee determined that the thesis is acceptable in form and content and that a satisfactory knowledge of the field covered by the thesis was demonstrated by the candidate during an oral examination. A signed copy of the Certificate of Approval is available from the School of Graduate and Postdoctoral Studies.

ii

ABSTRACT

As global warming and energy crisis issues continue to increase, it becomes critical to

investigate new sources of clean and affordable energy. Liquid hydrogen, is an

attractive energy alternative as the byproduct of hydrogen combustion is non-pollution

and useful water vapor. High hydrogen liquefaction work represents the most important

obstacle to achieving feasibility in the hydrogen economy.

In this thesis, a hydrogen liquefaction system is analyzed by using a hydrogen

liquefaction method both with and without catalyst infused heat exchangers. The goal

is to assess, modify and improve the proposed systems with the ultimate goal of

achieving sustainable and environment friendly hydrogen production.

The main objective of this work is to present detailed thermodynamic,

environmental, and economic analyses of the proposed multi-generation energy

systems. The study shows that when compared to the primary (main system), significant

improvements in energy and exergy efficiencies can be made by modifying the system

by employing vortex tubes, Organic Rankine Cycle (ORC), and the aid of a catalyst. In

fact, at 25oC the overall exergy efficiency of a configuration employing ORC is 42%

as opposed to 12% for the main system. This system also has the highest energy

efficiency of 76% as oppose to 10% for the main base system.

KEYWORDS: Hydrogen Liquefaction, Exergy Analysis, Energy Analysis, System

Optimization

iii

STATEMENT OF CONTRIBUTIONS

I hereby certify that I am the sole author of this thesis and that no part of this thesis has been published or submitted for publication. I have used standard referencing practices to acknowledge ideas, research techniques, or other materials that belong to others. Furthermore, I hereby certify that I am the sole source of the creative works and/or inventive knowledge described in this thesis.

iv

AUTHORS DECLARATION

I hereby declare that this thesis consists of original work of which I have authorized.

This is a true copy of the thesis, including any required final revisions, as accepted by

my examiners. I authorize the Ontario Tech University to lend this thesis to other

institutions or individuals for the purpose of scholarly research. I further authorize

University of Ontario Institute of Technology to reproduce this thesis by photocopying

or by other means, in total or in part, at the request of other institutions or individuals

for the purpose of scholarly research. I under-stand that my thesis will be made

electronically available to the public.

Student name and signature

v

ACKNOWLEDGMENTS

I would like to express my most profound gratitude and appreciation to my supervisor,

Professor Ibrahim Dincer, for giving me this unique opportunity to work with him. His

immense knowledge, constant guidance, intellect, commitment, and passion as a

scientist encouraged me to be more dedicated throughout my research journey. Dr.

Dincer has unfailingly provided me with his expert guidance, which has significantly

helped me overcome the challenges I faced during my Ph.D. research. Without Dr.

Dincer’s patience, it would not have been easy to complete my research.

I am also grateful to my colleagues and friends in Professor Dincer`s research

group who have consistently given me both support and motivation.

Special thanks and love go to my beautiful wife, friends and family for their

support and encouragement to continue and finish my studies. They always been my

primary motivation during my stay in Canada. I am sincerely grateful for their sacrifice

and blessings.

Last but not the least; I would like to express my most profound love and thanks

to my mother for her endless support, determination, inspiration and reassurance

throughout my life. Her prayers and good wishes have always been with me. I will

always be grateful to her.

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TABLE OF CONTENTS ABSTRACT i

AUTHORS DECLARATION iii

ACKNOWLEDGMENTS v

NOMENCLATURE ix

LIST OF TABLES xi

LIST OF FIGURES xii

INTRODUCTION 1 1.1 Overview and Outlook for Hydrogen 1 1.2 Hydrogen Production and Storage 4 1.3 Hydrogen Liquefaction 5 1.4 Improvements in Hydrogen Liquefaction 7 1.5 Motivation and Novelties of the Thesis 8 1.6 Thesis Objectives 8 1.7 Thesis Outline 9

LITERATURE REVIEW 11 2.1 Ideal Work of Hydrogen Liquefaction 11 2.2 Hydrogen Liquefaction Plants 14 2.3 Previous and Current Plants 15 2.4 Advanced Liquefaction Techniques 17 2.5 Existing Large-Scale Plants 18 2.6 Lab Scale Hydrogen Liquefaction Methods 19 2.7 Future Developments and Presentation of New Conceptual Designs 21 2.8 Orthohydrogen and Parahydrogen 24 2.9 Catalyst Conversion of Ortho- to Parahydrogen 28 2.10 Organic Rankine Cycle (ORC) 29 2.11 Vortex Tube 29 2.12 Closing Remarks 30

SYSTEM DESCRIPTION 31 3.1 Description of System 1 33

3.1.1 System 1A: Reference system without a catalyst 33 3.1.2 System 1B: Reference system with a catalyst 34

3.2 Description of System 2 38

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3.2.1 System 2A: System with ORC and without a catalyst 38 3.2.2 System 2B: System with ORC and with a catalyst 39

3.3 Description of System 3 41 3.3.1 System 3A: System with Vortex Tubes and without a catalyst 42 3.3.2 System 3B: System with Vortex Tubes and with a catalyst 42

SYSTEM ANALYSIS, MODELLING AND SIMULATION 45 4.1 Basic Thermodynamic Concepts 45 4.2 Conservation of Mass Principle 46 4.3 Conservation of Energy Principle 46 4.4 Entropy Balance and Entropy Generation 46 4.5 Exergy Analysis 47 4.6 Components used in the systems 49 4.7 Energy and Exergy Efficiencies 51 4.8 Sustainability Assessment 53 4.9 Exergoeconomic Assessment 54 4.10 Environmental Impact Assessment 56 4.11 Optimization Study 57

RESULTS AND DISCUSSION 60 5.1 Base System Results 60 5.2 Systems 1A and 1B Results 73

5.2.1 Pre-cooling phase at systems S1A and S1B 77 5.2.2 Liquefaction Phase at systems S1A and S1B 83

5.3 Systems 2A and 2B Organic Rankine Cycles 92 5.3.1 Pre-cooling phase at systems S2A and S2B 96 5.3.2 Liquefaction Phase at systems S1A and S1B 101

5.4 Systems 3A and 3B – Vortex tubes 111 5.4.1 Pre-cooling phase at systems S3A and S3B 112 5.4.2 Liquefaction Phase at systems S3A and S3B 118

5.5 Property set 129 5.6 Comparative analysis results 131 5.7 Optimization results 133

5.7.1 Objective function 133 5.7.2 Design conditions 133 5.7.3 Variables 134 5.7.4 Constraints 134

viii

5.8 Optimum case 134 5.9 Simulation comparison 135

CONCLUSIONS AND RECOMMENDATIONS 138 6.1 Conclusions 138 6.2 Recommendations 139

REFERENCES 140

ix

NOMENCLATURE

c Unit cost of exergy, $/kJ E Energy, kJ ex Specific exergy, kJ/kg Ex Exergy rate, kW 𝐸��& Exergy destruction rate, kW h Specific enthalpy, kJ/kg K Equilibrium constant m Mass, kg �� Mass flow rate, kg/s 𝑛 Number of moles �� Molar flow rate, mole/s P Pressure, kPa �� Heat flow rate, kW s Specific entropy, kJ/kg K ��+,- Entropy generation rate, kW/K T Temperature, °C or K t Time, s u Specific internal energy, kJ/kg U Internal energy, kJ V Volume, m3 𝜐 Specific volume, m3/kg Y Yield, % �� Capital cost, $ Greek Letters 𝜌 Density, kg/m3 𝜂 Energy efficiency 𝜓 Exergy efficiency Subscripts 0 Ambient conditions cv Control volume cold Cold ex Exergy hot Hot in Inlet out Exit consumption Consumption loss loss Q Thermal energy s Surface w Work

x

Acronyms CHP Combined heat and power CPV Concentrated photovoltaic COP Coefficient of Performance EX Expansion valve HE Heat Exchanger EX Heat Exchanger HX Heat Exchanger H2FEED Hydrogen Feed LH2 Liquified Hydrogen LN2 Liquid Nitrogen LNG Liquefied Natural gas ORC Organic Rankine Cycle SI Sustainability Index TPD Ton per Day VT Vortex Tube

xi

LIST OF TABLES

Table 1.1 Exergy efficiencies of different renewable energy hydrogen production methods

before and after liquefaction. ..................................................................................................... 5

Table 2.1 Details of Commercial Hydrogen Liquefaction Plants ............................................ 20

Table 2.2 Efficiencies of some conceptual plants .................................................................... 25

Table 4.1. Energy Balance for System Components ............................................................... 50

Table 4.2 Base system components exergy equations ............................................................ 55

Table 4.3. Properties of Hydrogen, Nitrogen and Carbon dioxide .......................................... 59

Table 5.1 Constraints of Selected Variables .......................................................................... 134

Table 5.2 Simulation and initial experimental data of the proposed system ......................... 137

xii

LIST OF FIGURES

Figure 1.1 Total Primary Energy Consumption – World (data from [1]) .................................. 1

Figure 1.2 World primary energy consumption (data from [3]) ................................................ 2

Figure 1.3 Generalized process flow for syngas production and industrial hydrogen (adapted

from [8]). .................................................................................................................................... 4

Figure 1.4 Graphical overview of hydrogen storage technologies (adapted from [9]) .............. 5

Figure 1.5 Schematic of the Vortex Tube, showing the inlet (𝑉𝑖𝑛) and cold (𝑉𝑐𝑜𝑙𝑑) and hot (𝑉ℎ𝑜𝑡)

outlets. Adapted from [13] ......................................................................................................... 7

Figure 2.1 Linde-Hampson liquefaction cycle schematic representation (adapted from [29]).

.................................................................................................................................................. 13

Figure 2.2 Schematic representation of the Claude cycle (adapted from [30]). ...................... 14

Figure 2.3 Flow chart of the hydrogen liquefier system [25]. ................................................. 16

Figure 2.4 Praxair hydrogen liquefaction process flow diagram (left) and improved hydrogen

liquefaction process flow diagram (right), adapted from [38]). ............................................... 17

Figure 2.5 Basic scheme of hydrogen liquefaction process based on renewable energy (adapted

from [52]). ................................................................................................................................ 23

Figure 2.6 Equilibrium composition as a function of temperature (adapted from [54]) .......... 26

Figure 2.7 Spin isomers of molecular hydrogen (adapted from [57]). .................................... 26

Figure 3.1 Advanced hydrogen liquefaction systems considered for analysis. ....................... 31

Figure 3.2 The main systems schematic diagram .................................................................... 35

Figure 3.3 Schematic diagram for the reference system .......................................................... 36

Figure 3.4 Schematic diagram for the main system with reactor ............................................ 37

Figure 3.5 Organic Rankine Cycles ......................................................................................... 38

Figure 3.6 Schematic diagram for the system with ORCs ....................................................... 40

Figure 3.7 Schematic diagram for the system ORCs and reactor ............................................ 41

Figure 3.8 Added VTs .............................................................................................................. 42

Figure 3.9 Schematic diagram for the system with VTs and no reactor .................................. 43

Figure 3.10 Schematic diagram for the system with VTs and reactor ..................................... 44

Figure 5.1 Exergy and energy efficiencies for each component .............................................. 61

xiii

Figure 5.2 Effect of pre-cooling Hydrogen on Compressors work losses and yield. The graphs

illustrate a) work loss of compressors C1, C2, and C3 and, b) daily hydrogen yield against

hydrogen Feed temperature ...................................................................................................... 62

Figure 5.3. Effect of pre-cooling hydrogen on overall energy and exergy efficiencies .......... 63

Figure 5.4. Effect of hydrogen mass flow rate change on work and yield. The graphs show a)

work (kW) of (a)Turbo expander (TE1), Turbo expander (TE2), (b)Adsorber (A0) and; (c)

Liquid hydrogen generation per day against hydrogen Feed Mass Flow rate (kg/h) .............. 64

Figure 5.5 Effect of hydrogen mass flow rate change on the overall energy and exergy

efficiencies ............................................................................................................................... 65

Figure 5.6 Effect of changing turbo expander TE1 pressure on main components work ....... 65

Figure 5.7 Effect of changing turbo expander TE1 pressure on the overall energy and exergy

efficiencies ............................................................................................................................... 66

Figure 5.8 Effect of changing flash drums pressure on the overall yield ................................ 67

Figure 5.9 Effect of flash drum pressure rate change on overall energy and exergy .............. 67

Figure 5.10 Effect of hydrogen feed pressure change on the compressors ............................. 68

Figure 5.11 Effect of hydrogen feed pressure change on the overall energy and exergy

efficiencies ............................................................................................................................... 68

Figure 5.12 Effect of the Nitrogen gas mass flow rate change on the overall energy and exergy

efficiencies ............................................................................................................................... 69

Figure 5.13 Heat Exchanger HX1 Heat composite curves ...................................................... 70



Figure 5.16 Energy efficiencies for main components of System 1A and 1B ......................... 74

Figure 5.17. Effect of hydrogen feed pressure on overall efficiencies for System 1A ............ 75

Figure 5.18 Effect of Compressor C1 pressure on overall efficiencies for System 1A ........... 75

Figure 5.19 Effect of Compressor C2 pressure overall efficiencies for System 1A ................ 75

Figure 5.20 Effect of hydrogen H2 feed pressure on overall efficiencies for System 1B ........ 76

Figure 5.21 Effects of Compressor C2 pressure on overall energy and exergy efficiencies for

System 1B ................................................................................................................................ 76

xiv

Figure 5.22 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger

HX1 at stream N2LIQ inlet for System 1A ............................................................................. 77


HX1 at stream 28 inlet for System 1A ..................................................................................... 78


HX1 at stream 9 inlet for System 1A ....................................................................................... 78






HX1 at stream R inlet for System 1A ...................................................................................... 80

Figure 5.28 Heat Load, Exergy flow vs Temperature for Precooling Phase heat exchanger HX1

at stream N2LIQ inlet for System 1B ...................................................................................... 80


HX1 at stream 28 inlet for System 1B ..................................................................................... 81


HX1 at stream 9 inlet for System 1B ....................................................................................... 81









Figure 5.35 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger

HX2 at stream S10 inlet for System 1A ................................................................................... 84





xv


HX3 at stream 32b inlet for System 1A ................................................................................... 86


HX3 at stream S1 inlet for System 1A ..................................................................................... 87




HX2 at stream S10 inlet for System 1B ................................................................................... 88












HX3 at stream S1 inlet for System 1B ..................................................................................... 91


HX3 at stream 32b inlet for System 1B ................................................................................... 92

Figure 5.49 Systems (a) 2A and (b) 2B Exergy and Energy efficiencies for each component 93

Figure 5.50. Effect of pre-cooling hydrogen feed pressure variations on overall energy and

exergy efficiencies for systems (a) S2A and (b) S2B .............................................................. 94

Figure 5.51 Effect of Compressor 1 (C1) pressure variations on overall energy and exergy

efficiencies for systems (a) S2A and (b) S2B .......................................................................... 95

Figure 5.52 Effect of Compressor 5 (C5) pressure variations on overall energy and exergy

efficiencies for systems (a) S2A and (b) S2B .......................................................................... 96


HX1 at stream N2LIQ inlet for System 2A ............................................................................. 97

xvi




HX1 at stream 9 inlet for System 2A ....................................................................................... 98






at stream N2LIQ inlet for System 2B ...................................................................................... 99


HX1 at stream 28 inlet for System 2B ................................................................................... 100







Figure 5.63 Heat Load, Exergy flow and Temperature for liquefaction Phase heat exchanger

HX2 at stream 49 inlet for System 2A ................................................................................... 102










HX3 at stream 32b inlet for System 2A ................................................................................. 105



xvii


















HX3 at stream 32b inlet for System 2B ................................................................................. 111

Figure 5.79 Systems 3A and 3B Exergy and Energy efficiencies for each component ........ 112


HX1 at stream N2LIQ inlet for System 3A ........................................................................... 113










at stream N2LIQ inlet for System 3B .................................................................................... 116

xviii














HX2 at stream S10 inlet for System 3A ................................................................................. 120






HX3 at stream 32b inlet for System 3A ................................................................................. 122








HX2 at stream S10 inlet for System 3B ................................................................................. 125





xix








HX3 at stream 32b inlet for System 3B ................................................................................. 128

Figure 5.106 T-xy plot for temperature versus liquid composition for isobaric data ........... 129

Figure 5.107 T-x plot for temperature versus liquid composition for isobaric data .............. 130

Figure 5.108 K-values for Vapor-liquid vs fraction of Para-hydrogen and Ortho Hydrogen 130

Figure 5.109 y-x diagram for vapor vs liquid composition for the para-hydrogen ............... 131

Figure 5.110 Activity coefficients vs mole fraction for Para-hydrogen and orthohydrogen . 131

Figure 5.111 Exergy efficiency for the proposed hydrogen liquefaction systems at 0°C, 10°C,

25°C, and 45°C. ..................................................................................................................... 132

Figure 5.112 Energy efficiency for the hydrogen liquefaction systems. ............................... 132

Figure 5.113 Work done for liquefaction per unit mass (kJ/kg). ........................................... 133

Figure 5.114 Overall exergy efficiency ................................................................................. 135

Figure 5.115 Optimized system ............................................................................................. 136

1

INTRODUCTION

1.1 Overview and Outlook for Hydrogen

Over the last few decades, global energy consumption has grown continuously and this

trend is expected to continue as presented in Figure 1.1 [1]. This significant growth has

created considerable interest in sustainable energy, including hydel, wind, solar, and

geothermal resources. Over the past century, factors such as population growth,

increasing water demand, industrial development, and the bulk production of agriculture

products have also motivated the research and development of substitutes for fossil fuels.

Figure 1.1 Total Primary Energy Consumption – World (data from [1])

The share of sustainable energy in the global market is around 24% as of 2017

and advancement is dependent on technological developments, society, world politics,

and the environment. Over 75% of the total energy consumption comes from fossil

fuels, according to the World Bank data. Additionally, carbon-based fuels are very

harmful to the environment, causing 87% of carbon emissions [2]. Hence, a substitute

for fossil fuels is required since they are finite resources that may create and cause

future crises and instability. Alternative clean energy sources are needed to meet rising

energy demands in an environmentally friendly manner. As clean energy sources,

renewable energies can be considered as sustainable alternatives due to their significant

advantages over fossil fuels. Figure 1.2 shows the energy share between traditional and

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renewable sources. Hydrogen, which is considered as a sustainable energy source, can

be considered in direct competition with electricity as an energy carrier, with each using

a separate production and distribution system.

Figure 1.2 World primary energy consumption (data from [3])

As an energy carrier, hydrogen plays a crucial part in the reduction of

greenhouse gas emissions thereby curtailing the effects of climate change. It is therefore

often considered as the energy carrier of the future. Since it can primarily be produced

from water, hydrogen can provide a solution to issues of sustainability, greenhouse gas

and other pollutant emissions, and also offer energy security. A hydrogen economy that

is the same size as the U. S. would require approximately 150 million tons per year of

hydrogen for transportation, which would be equal to the consumption of two to five

billion tons of water, taking into account current hydrogen production efficiencies. This

consumption would be considerably less than the current consumption of water for

thermoelectric power generation in the U. S. in power plants, which is approximately

300 billion tons, while an additional US1.2 billion 1.2 billion is spent in the process of

gasoline production. Therefore, the most likely scenario is that the hydrogen economy

could significantly reduce water consumption in the process of energy generation [4].

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3

Hydrogen energy content per weight is around 125 MJ kg-1, over two times

higher than any other fuel currently in use. Fossil fuels have a weight in the range of

20-50 MJ kg-1, with diesel fuel at the high end and natural gas at the low end, while

batteries have 0.1-0.5 MJ kg-1 [5]. This energy outcome makes hydrogen the most

effective energy carrier. As a result, considerable effort has been expended on

improving its production and storage. However, energy content per volume of hydrogen

is relatively low when not highly compressed or liquefied. Even then, it is significantly

lower than that of fossil fuels: 8 MJ L-1 for liquid hydrogen, the most efficient form of

hydrogen storage, in comparison to 32 MJ L-1 for gasoline. In spite of intense research

efforts that have been devoted to the development of more efficient means of storing

hydrogen, hydrogen liquefaction remains the most economical method of hydrogen

storage to date, regardless of its deficiencies. Therefore, when considering the

efficiency of hydrogen production, the cost of liquefaction should be considered.

The phrase “hydrogen economy”, which was first used by John Bockris in 1970,

refers to a proposed system of energy delivery utilizing hydrogen as an energy carrier.

This concept is meant to alleviate some of the adverse effects of hydrocarbon fuel

consumption as both the primary source of energy and as the main energy carrier. Since

burning fossil fuels leads to the emission of carbon dioxide and other pollutants that

have considerable adverse effects on the environment, a global hydrogen economy is

seen as an environmentally eco-friendlier alternative for delivering energy to end-users,

particularly in transport. A 2004 study found that "most of the hydrogen supply chain

pathways would release significantly less carbon dioxide into the atmosphere than

would gasoline used in hybrid electric vehicles" and that significant reduction in carbon

dioxide emissions could result from utilizing carbon capture or carbon sequestration

technology at the site of energy or hydrogen production [6].

If hydrogen is considered to be a renewable fuel for the future, with the

numerous challenges that come with its production, storage and use, the issue of

efficiency is considered as just one of the many factors that determine the viability and

usability of these systems. The investment of both financial and other resources in

durability, and the stability of operation along with safety are important parameters in

determining the viability of each proposed solution that might become a part of the

hydrogen cycle.

4

1.2 Hydrogen Production and Storage

Most of the hydrogen produced today is obtained from methane reforming, with carbon

dioxide as the side product of the reaction. A generalized process flow is shown in

Figure 1.3. Since this increase greenhouse gas emission, significant efforts have been

invested in the exploration of alternative methods of hydrogen production that rely on

renewable energy. The use of hydrogen produced in this alternative process could

contribute to a reduction in the level of greenhouse gases. There are several of these

methods, which differ significantly in their efficiencies [7].

Figure 1.3 Generalized process flow for syngas production and industrial hydrogen

(adapted from [8]).

Hydrogen storage includes a number of methods that can be broadly divided

into three groups: mechanical storage: storage in a solid material through physisorption:

and solid material storage through chemical bonding or chemisorption. Each of these

has advantages and disadvantages, and a graphical overview of their capabilities is

given in Figure 1.4.

In order to overcome storage issues, hydrogen has been liquefied at high cost

but room for improvement in the liquefaction process is still present. Table 1.1 shows

the efficiencies of different renewable energy sources used in hydrogen production

before and after liquefaction. The difference in efficiencies clearly show that

liquefaction has a great room of improvement through advanced research and

development studies.

Cool/Biomass

Natural gas/ fuel oil

Natural gas/ naphtha

Syngas(H2, CO, CO2, H2O)

O2 and/or Air Steam

(Water- gas shiftH2,O, H2, CO, CO2,

shift to H2 and CO2)

(H2 and CO2 separation)

PSA, physical absorption

H2 and CO2 to ammonia and F-T Processes

CO2:- Vented - To urea production - Enhanced oil recovery

Gasifier

Partial OXI (POX)

Reformer (SMR/ATR)

Shift Reactor Gas Clean up

O2 and/or Air Steam

Feedstock in:

5

Figure 1.4 Graphical overview of hydrogen storage technologies (adapted from [9])

Table 1.1 Exergy efficiencies of different renewable energy hydrogen production methods before and after liquefaction.

Method Exergy efficiency after electrolysis (%)

Exergy efficiency after liquefaction (%)

Photovoltaic solar 7.2 1.0 Photothermal solar 8.6 1.2 Wind power 30.8 4.1 Hydropower 41.6 5.6 Biomass Combustion 27.5 3.9 Biomass Gasification 40.6 5.4

Source: [7].

Mechanical storage methods, in both gas and liquid form, are the most common

hydrogen storage methods used today. Storage of liquid hydrogen creates less risk than

high-pressure gas storage and represents the dominant form of hydrogen storage for

large-scale production and transport.

1.3 Hydrogen Liquefaction

Liquefied hydrogen has thus far played a role mostly as fuel for space exploration and

other related applications, as well as in the semiconductor industry. However, with a

shift in the energy landscape towards clean energy, it is expected that liquid hydrogen

will be used as a future energy carrier, both for automotive transportation and long

6

distance and overseas transport. In order to achieve this, hydrogen liquefaction

technology requires significant improvements to reduce energy consumption from the

current levels of 12-15 kWh per kg of liquid hydrogen to about 6 kWh per kg. This can

only be achieved through a series of improvements in the different stages of hydrogen

liquefaction. Some of them are already available, such as: improved operation of the

recycled gas compression system that applies chillers; closed refrigeration loops for

pre-cooling; upgraded turbine designs; and adjusted concepts for the main refrigeration

loop. Adoption of these technologies could reduce energy consumption to 7.5 to 9 kWh

per kg of liquid hydrogen [34]. Further improvements could be gained by applying

more innovative process schemes, such as that described above, as well as improved

machinery and equipment.

Some of those improvements could include turboexpanders. They are

considered to be one of the most challenging aspects of hydrogen liquefaction [10],

where low molecular weight and size, as well as the high speed of sound of hydrogen,

require very high peripheral speeds. The material properties and the difficulty presented

in forming reliable seals, along with the propensity of hydrogen to embrittle materials,

restrict the speed as well as limit each stage to a low-pressure ratio. In addition,

hydrogen gas must be thoroughly purified to remove oxygen and other possible

contaminants that would freeze in the system, clogging the Joule-Thomson valve or

damaging an expander. The presence of frozen oxygen in the product tank also

represents a potential explosion hazard, therefore purification requirements for

hydrogen liquefiers can be as high as 1 ppm. A number of authors point out that due to

the long development history and long-term qualification procedures, improvements in

modern components can be difficult to achieve and implement, and their

implementation could most likely be achieved only on a long-term basis [11,12], due

to a relatively small number of plants in operation. However, if more hydrogen

liquefaction plants are set up due to increasing demand, it may become economically

feasible to develop new components specifically designed for liquefaction cycle flow

rates and pressure levels, providing an additional increase in the efficiency of the

liquefaction process.

Other techniques to improve the liquefaction process could include Vortex

tubes. The Vortex tube devices cool a fluid by separating the inlet flow into an inner

cold and outer hot stream within the tube, as described in Figure 1.6. To achieve higher

7

efficiency, the components in the streams have to compress. Vortex tubes can be

considered to be effective solutions for heat exchange applications. As shown in Figure

1.6, the fluid is tangentially injected at high pressure and expanded in the tube. The

fluid flows into the tube at a high velocity along the sidewall, and a cold stream is then

created through the expansion of the centre of the tube. Fluid with a higher temperature

than Vin ejects from the Vhot outlet and the flow that has not been ejected flows back

into the center of the tube, exiting through the Vcold outlet[13].

Figure 1.5 Schematic of the Vortex Tube, showing the inlet (𝑉𝑖𝑛) and cold (𝑉𝑐𝑜𝑙𝑑) and hot (𝑉ℎ𝑜𝑡) outlets. Adapted from [13]

1.4 Improvements in Hydrogen Liquefaction

Krasae-in et al. proposed a series of improvements to the overall design of the hydrogen

liquefaction process, primarily through the application of a multi-component refrigerant

refrigeration system [14,15]. The new system uses a mixture of 4% neon, 12% nitrogen,

26% methane, 30% ethane, and 28% butane as a coolant, resulting in lower power

consumption (for the production of 100 tons per day) on the pre-cooling, compared to

conventional refrigeration systems with 1.76 kWh per kg of hydrogen, compared to

4.86 kWh per kg for an actual hydrogen liquefaction plant in Ingolstadt (capacity 4.4

tons per day) [16]. This could reduce the specific energy consumption for liquefaction

of the overall cycle from 13.58 to 5.35 kWh per kg of hydrogen.

In the near future, the application of these improvements, the development of

several new improved technologies and reduced costs due to economies of scale could

make hydrogen liquefaction an economically viable solution for energy storage in a

renewable energy sourced hydrogen economy.

Vin

Vhot

Vcold

8

1.5 Motivation and Novelties of the Thesis

With the current rapid worldwide energy consumption, fossil-fuel reserves prove to be

in constant reduction [17]. In 2012, the total world energy consumption amounted to

580 kJ, which is expected to rise to 711 kJ by 2025 and 860 kJ in 2040 [1]. In 2016,

the U.S. Energy Information Administration reported that the global use of petroleum

and other liquid fossil fuels had risen from 67.2 million barrels per day in1990 to 90.3

million barrels per day in 2012, and would rise to 109.1 million barrels per day in 2030,

and even 120.9 million barrels per day by 2040.

Growing concerns over greenhouse gas emissions from the combustion of fossil

fuels, and an awareness of the need for a clean high-energy fuel, have prompted interest

in the production of hydrogen. Building a hydrogen-based economy for a sustainable

energy system is the long-term view of many [18]. The main goal of building a

hydrogen economy is to replace fossil-based energy sources with hydrogen. The

technology of producing, liquefying, and storing hydrogen is vital for its feasibility.

Therefore, it is extremely important to build a sophisticated and efficient system that is

a feasible replacement for another energy source.

Hydrogen energy systems represent a potential solution for these highly

important problems, where the requirement is to deliver high-efficiency output while

lowering the total emissions per energy used.

From the open literature, it can be seen that researchers and scientists have

attached considerable weight to hydrogen production systems. However, there has not

yet been sufficient research on hydrogen liquefaction systems, especially catalyst-based

models.

The systems presented in the thesis will contribute in the following: (i) new

advanced liquefaction system configurations, (ii) ull and comprehensive analysis for

systems and (iii) simulation of each system.

1.6 Thesis Objectives

The main objective of this work is to outline the novel advanced hydrogen liquefaction

systems. This will include energy, exergy and environmental analyses to compare the

studied systems. In more detail, these objectives can be listed as follows:

9

• To develop and design three advanced hydrogen liquefaction systems with

two configurations each based on a patented commercial system number

US8042357B2 [19] that was never studied and analyzed.

• To perform comparative analysis and simulate multiple advanced hydrogen

liquefaction systems using ASPEN plus.

• To build a novel configuration of an advance hydrogen liquefaction system,

involving the development of a complete thermodynamic model with full

exergy analysis of the proposed systems, including calculating exergy flow,

energy and exergy efficiency, exergy destruction ratios, and other related

thermodynamics measures.

• To have environmental impact assessment of each of the proposed systems

by studying the CO2 emission and sustainability indices.

• To conduct a parametric study to evaluate system performance utilising

parametric study on individual components and the effect of environmental

conditions on each system.

• To determine the optimum design parameters through an optimization

analysis of the proposed liquefaction system using Matlab.

1.7 Thesis Outline

This thesis consists of six main chapters. The first chapter includes an introduction and

background information regarding the hydrogen economy and hydrogen liquefaction

developments over time. Furthermore, the novelties of the proposed integrated systems,

together with the motivation and objectives of this thesis, are included. Chapter 2

provides a comprehensive literature review of different advanced liquefaction

techniques, ortho- parahydrogen and catalysts. Moreover, a literature review of the

different components that will be utilized in the proposed integrated systems, such as

Organic Rankine Cycles (ORC) and Vortex Tubes (VT), is incorporated. Chapter 3

explains in detail the proposed systems and their components. Chapter 4 contains the

general thermodynamic equations that are used to model the introduced integrated

systems along with detailed thermodynamic modelling for the main components in each

integrated system. An exergoeconomic analysis is the main part of the system along

with optimization. Chapter 5 shows the results of the systems, combined with a

comprehensive comparison. The results of the exergoeconomic analysis and an

10

optimization study for each system are also provided. Chapter 6 highlights conclusions

from the thesis together with recommendations for future studies.

11

LITERATURE REVIEW

The significance of hydrogen as a clean energy source in the present world has already

been explained in Chapter 1. This chapter provides a brief literature review of the

hydrogen liquefaction process which is essential for using hydrogen as a fuel. Hydrogen

was first liquefied by Sir James Dewar in 1898 [20]. Following this significant step,

several procedures were developed for hydrogen liquefaction, forming a broad range of

technological solutions from laboratory liquefaction apparatus to large-scale plants

[21]. In this present review, developments from 1898 to 2016 are presented.

Hydrogen has shown potential as an important energy carrier for use in

transportation vehicles of the future, leading to considerable hydrogen research activity.

The greatest challenge today is the relatively low efficiency of the currently used

liquefaction plant cycles. Several recent studies have explained methods and ways to

overcome efficiency issues, where some have proposed conceptual plants with

efficiencies that can be increased up to 40– 60% [20].

In this chapter, the literature available regarding the characteristics of hydrogen

are discussed first. The process of hydrogen liquefaction and its evolution over time is

then covered followed by a discussion on laboratory scale hydrogen liquefaction

processes. The chapter concludes by identifying the literature gaps and the necessity

for researching the various hydrogen liquefaction systems.

2.1 Ideal Work of Hydrogen Liquefaction

The first successful hydrogen liquefaction was achieved in 1898 by a small device made

and invented by Scottish scientist James Dewar [22,23]. Dewar’s process used a

combination of carbolic acid and liquid air to pre-cool compressed hydrogen gas at 180

bars and the Joule-Thompson effect for liquefaction [20]. The amount of work required

by a reversible cycle to bring hydrogen from the initial conditions, e.g. 300K, 100 kPa,

and 25% parahydrogen, to the final liquid state at 100 kPa and equilibrium

parahydrogen content is referred to as the ideal work of hydrogen liquefaction.

Most current hydrogen liquefier systems utilize steady flow processes,

including the pre-cooled Linde-Hampson cycle, the Claude cycle and the helium

hydrogen condensing cycle [24]. The choice of a particular thermodynamic cycle

12

depends on the projected size of the plant, the available level of technology, equipment

cost and, principally, cycle efficiency[25].

The most simplified hydrogen liquefaction cycle is the Linde-Hampson or

Joule-Thomson expansion cycle, as shown in Figure 2.1. The process consists of

compressing gas at ambient pressure, cooling it in a heat exchanger and then passing it

through a throttle valve producing liquid through isenthalpic Joule-Thomson

expansion. The liquid product is then collected and removed, while the cooled gas is

returned to the compressor through the heat exchanger. However, unlike most gasses,

hydrogen warms under expansion at room temperature, and therefore requires pre-

cooling to the temperature below the corresponding inversion temperature (which

depends on pressure), which is typically 78 K. This is usually accomplished using liquid

nitrogen as a coolant, where nitrogen gas can be recovered and then reused in a

continuous refrigeration loop[26,27].

Most large-scale hydrogen liquefaction processes are based on the Claude cycle,

as illustrated in Figure 2.2, where hydrogen is both the product and the working fluid

[16,21]. One or more heat exchangers reduce the temperature of the working fluid and

a Joule-Thomson valve brings the fluid into the two-phase regime when the saturated

liquid is removed from the cycle. The input of gas at the warm end maintains a constant

mass of hydrogen in the system. Modifications of the Claude cycle include the addition

of a second compressor, where the first compresses hydrogen from low to medium

pressures and the second compresses from medium to high pressures. In this case, the

expander operates between medium and low pressures, providing additional cooling to

the high-pressure gas through its exhaust. Variations of this system are often used in

large-scale hydrogen liquefaction plants [23,28], combined with nitrogen pre-cooling,

multiple ortho-para conversion catalysts and, typically, two or three expanders.

One cycle that can be considered as a combination of a helium refrigerator

(Claude cycle) and a hydrogen liquefier (pre-cooled Linde-Hampson) cycle is the

helium hydrogen condensing cycle that utilizes helium as a primary refrigeration

working fluid. The main advantage of this cycle is its safety features, where hydrogen

compression is relatively limited and only increases to a high enough pressure to

overcome the pressure drop in the heat exchangers. Considering this pressure change,

and with the helium gas temperature at below 20 K, complete liquefaction of hydrogen

13

can be achieved after the expansion by using the correct ratio of flow rates of helium

and hydrogen. However, a return hydrogen stream can be constructed by the heat of

conversion of ortho- to parahydrogen in the liquid hydrogen receiver.

Figure 2.1 Linde-Hampson liquefaction cycle schematic representation (adapted from

[29]).

A comparison of these three hydrogen liquefaction cycles [24] shows that while

the Linde-Hampson and Claude cycle have a liquid hydrogen yield of 12-20%, helium

hydrogen condensing achieves a 100% yield for normal hydrogen and a 54% yield for

parahydrogen production. However, energy-wise, the Claude cycle is the most efficient

14

with an energy cost of 100-140 MJ kg-1, followed by helium hydrogen condensing with

120-200 MJ kg-1, and Linde-Hampson with a considerably higher 260-285 MJ kg-1 [26].

The liquefaction processes covered in this section are employed in liquefaction plants

to produce hydrogen. The next sections describe the transition of hydrogen liquefaction

from small scale plants over time starting from plants using Dewar’s process to more

recent plants employing advanced hydrogen liquefaction systems.

Figure 2.2 Schematic representation of the Claude cycle (adapted from [30]).

2.2 Hydrogen Liquefaction Plants

Since Dewar’s first successful hydrogen liquefaction in the late 1800s [23], more

efficient systems, such as Claude, pre-cooled Claude and the helium-refrigerated

system, were developed in the early 1900s [31]. Construction of the first large hydrogen

plants in the United States took place in 1957, to satisfy the growing needs of the

aerospace and petrochemical industries. These plants utilized the modified pre-cooled

Claude cycle, using liquid nitrogen as a pre-coolant, cooling input hydrogen to 80 K,

which was then further cooled down to 20 K using the hydrogen refrigeration system.

These plants had an energy efficiency of less than 20%, with a focus on reliability and

safety rather than maximum efficiency. Since then, there has been little improvement

15

in this regard; most large-scale hydrogen plants in operation today use similar cycles as

these first types with energy efficiencies of up to 40% [32]. A study of the efficiency

of hydrogen liquefaction plants by Linde Kryotechnik and the Nippon Sanso

Corporation [33] showed that about one-third of the exergy is lost in the liquefier

system, with an additional 27% lost before the cold end of the liquefier. Only 39.7%

exergy is left as product flow.

A typical liquefier system is shown in Figure 2.3, which includes a hydrogen

feed stream entering the cold box and featured a continuous conversion of ortho- to

parahydrogen during cooling using a catalyst placed directly in the heat exchangers.

Hydrogen pressure between the seventh and eighth heat exchangers (HE 7 and HE 8)

is reduced to tank pressure with an ejector, which functions similar to a water jet blast,

removing displaced or flash gas from the tank to be re-liquefied in HE 8.

Refrigeration of hydrogen gas down to about 80 K is achieved utilizing a

nitrogen pre-cooler by running liquid nitrogen through a phase separator and flooding

HE 2 with liquid, which cools the hydrogen stream down to approximately 81 K. This

refrigeration process creates evaporated nitrogen, which is further used in HE 1 to cool

the feed stream, while warming up to ambient temperature. Refrigeration from 80 K to

about 30 K is accomplished through expansion of high-pressure hydrogen gas (2 MPa)

in three expanders placed in a series. The Joule-Thomson cycle is applied in HE 7 and

8 to cool hydrogen from 30 K to liquefaction. The system has a throttle valve at the

bottom, where the high-pressure gas is throttled to low pressure, reducing its

temperature in the process. This is the lowest temperature point of the system. After

removal of the liquid, the warmed-up gas is again compressed, to medium pressure, and

inserted into the stream of the return gas from the expanders.

2.3 Previous and Current Plants

As the demand for liquid hydrogen grew, the liquefaction plants also changed to

accommodate the demand. Economies of scale mean that centralized hydrogen

production is more cost effective and energy efficient than distributed production. In

fact, hydrogen liquefaction plants tend to be more efficient with an increase in size [34]

as well as limited by financial rather than technical constraints. The capital costs

accounts for approximately 63% of the total lifetime cost of a hydrogen liquefaction

plant.

16

The main operating cost of hydrogen liquefaction is the cost associated with

input power of 12-15 kWh/kg, accounting for about 32% of the total lifetime cost of a

hydrogen liquefaction plant [35,36]. The capacities of hydrogen liquefaction plants

vary from 5 tons per day for the Air Products plant in Sacramento to 66 tons per day

for the Air Products plant in New Orleans. An economic analysis of three hydrogen

liquefaction systems [37] illustrates that, while power consumption costs remain

relatively constant, fixed charges, as well as operation and maintenance costs, rapidly

decrease with increased production rates. The cost of production decreased to about

0.7$ and 0.8$ per kg H2 for a production rate of 29,700 kg per hour for an optimized

large-scale hydrogen liquefier and a two-stage Claude hydrogen liquefier, respectively.

Figure 2.3 Flow chart of the hydrogen liquefier system [25].

The currently used hydrogen liquefaction process is highly integrated with air

separation and typically uses liquid nitrogen as a coolant. In 2011, Praxair introduced

improvements to its existing hydrogen liquefaction process, as shown two

corresponding flow diagrams in Figure 2.6. The improved process reduces the overall

power consumption by 2.4% and the liquid nitrogen requirement by 11%. The cooling

load is moved from the second to the first heat exchanger, with the result that external

refrigeration is increased by 17% and the recycle flow is reduced [38]. In addition, it

was reported that the novel ortho-para conversion process was able to achieve an overall

17

process improvement of around 8%. However, all of these improvements failed to

deliver the desired 20% improvement in overall process performance. Therefore,

researchers have worked on new advanced methods with the motivation to improve

hydrogen liquefaction efficiency. These are discussed in the next section.

Figure 2.4 Praxair hydrogen liquefaction process flow diagram (left) and improved hydrogen liquefaction process flow diagram (right), adapted from [38]).

2.4 Advanced Liquefaction Techniques

The hydrogen liquefaction process has undergone numerous refinements since the first

successful hydrogen liquefaction. However, the fundamental elements of the early

liquefaction process have not disappeared from currently employed cycles. These

basics include [12]:

• Using Joule-Thomson expansion, where the pressure of the compressed

hydrogen is reduced using a nozzle or valve, which represents an adiabatic

process leading to a reduction in gas temperature. That is, at a hydrogen

inversion temperature of 204 K, this temperature reduction can be achieved only

with gas that has been pre-cooled in the process to a temperature below this

inversion point. Thus, hydrogen liquefaction requires more than just this

refrigeration technique.

18

• Using an external auxiliary refrigerating fluid in the hydrogen liquefaction

process in order to cool hydrogen below inversion point. Typical refrigeration

fluid is liquid nitrogen, where the nitrogen liquefaction may, in turn, use its

auxiliary refrigerating fluid, typically a halogenated hydrocarbon. In addition to

liquid nitrogen, helium is also used for some small-scale liquefiers as the

auxiliary refrigeration liquid in order to achieve hydrogen liquefaction via

Joule-Thomson expansion.

• Using an expansion engine for compressed hydrogen. In addition to expansion

by a Joule-Thomson valve, the compressed hydrogen can also expand in an

expansion engine. The developed work is then excluded from the system using

the engine shaft and can be recovered for additional external use.

The first reported hydrogen liquefaction cycles include the pre-cooled Linde-

Hampson cycle, invented in 1895, and the Claude cycle that was invented in 1903 [20].

Since many of the advancements were based on trying to achieve better efficiency, most

of the related literature focuses on optimization of the liquefaction process,

improvements to the process equipment, and the improvement of the ortho-para

conversion to reduce the amount of power consumed by liquefaction.

A patent by Schwartz et al. [39] identified and quantified some of the ways to

reduce the cost of the liquefaction process, which would, in turn, significantly reduce

the cost of hydrogen distribution. The aim of this research was to achieve a reduction

of 20% in power consumption, followed by a further reduction in capital cost. When

the targeted efficiency improvement was not achieved, Praxair Inc. stopped the project

before the potential savings in capital costs were addressed.

Other projects introduced by NASA were for small and medium-scale hydrogen

liquefaction processes. The technology included domestically produced wet cryogenic

turboexpanders [40]. The few large-scale hydrogen liquefaction plants in operation use

variations of the cycles described above.

2.5 Existing Large-Scale Plants

The most common cycles now used for hydrogen liquefaction are the helium Brayton

cycle and (pre-cooled) Claude cycle [11]. The Brayton cycle achieves refrigeration

capacity solely with expansion turbines, while the Claude cycle uses recycle

compressors with lower power consumption, and optimizing the refrigeration loop with

19

expansion by turbines and finally via a Joule-Thomson valve [11,41]. The energy

consumptions of the Brayton and Claude cycles are 12.3 to 13.4 kWh per kg H2 and

10.8 to 12.7 kWh per kg H2, respectively [11].

The first liquefaction plants were constructed in the late 1950s to support

NASA's programs [20,42]. Today, the greatest use of liquid hydrogen continues to be

in space programs as rocket fuel. In the future, if production efficiency is increased, it

might be used as a fuel for vehicles. Table 2.1 exhibits some hydrogen liquefaction

plants that exist globally. The largest producers are Praxair, Air Products, Air Liquide

and Linde [11].

Praxair Inc. currently has five fully operational hydrogen liquefaction plants in

the United States, with different production capacities ranging from 6 to 35 tons per

day (TPD) of liquid hydrogen. Production is based on a modified pre-cooled Claude

cycle, where the usual rates of power consumption are in the range of 12.5−15

kWh/kgLH2 [11]. Air Products has six hydrogen liquefaction plants: five in the US and

one in the Netherlands. Four out of five of the US plants have a producing capacity of

about 30 TPD, while the others (one in the US and one in the Netherlands) have a 5

TPD capacity [11]. Air Liquide has a plant in each of France and Canada, both with

production capacities of around 10 TPD. The two plants utilize the Claude cycle, where

hydrogen was used as the cycle fluid [11]. Linde is another large-scale producer whose

production is also based on the pre-cooled Claude cycle [11,42]. The plant is located in

Ingolstadt, near Munich. The liquefier has a capacity of 4.4 TPD [32]. In addition to

these large-scale operations, various experimental laboratory scale hydrogen

liquefaction processes are also in existence. It is hoped that over time, some of these

processes will develop sufficiently to be economically viable in an industrial setting.

2.6 Lab Scale Hydrogen Liquefaction Methods

Most of the recent designs on a small and laboratory scale are based on the

magnetocaloric effect. This kind of liquefier, together with an appropriate cyclic

thermodynamic process, can use isentropic demagnetization of a ferromagnetic

material near its Curie point temperature as a refrigeration procedure [42].

Magnetocaloric refrigerators were investigated and constructed for a temperature range

of about 1 K to 20 K.

20

Table 2.1 Details of Commercial Hydrogen Liquefaction Plants

Country Location Operated by Capacity Tons Per Day

Canada

Ontario Air Products 30.0

Quebec

Air Liquide 10.0 Air Liquide 12.0 BOC by Linde 15.0 BOC by Linde 14.0

French Guyana Kourou Air Liquide 5.0

USA

Missouri Air Products 3.0

Florida Air Products 3.2 Air Products 27.0

Mississippi Air Products 32.7 California Union Carbide/Linde Division 54.0 Louisiana Air Products 34.0 New York Praxair 18.0 California Air Products 6.0 New York Praxair 18.0 Mississippi Air Products 30.0 McIntosh Praxair 24.0 Indiana Praxair 30.0

France Lille Air Liquide 10.0

Germany Linde 4.4 Linde 5.0

Netherlands Air Products 5.0 China CALT 0.6

India ISRO 0.3 Asiatic Oxygen 1.2 Andhra Sugars 1.2

Japan

Iwatani 0.6 MHI 0.7 Tashiro 1.4 Pacific Hydrogen 1.4 Japan Liquid Hydrogen 2.2 Japan Liquid Hydrogen 0.3 Air Products 0.3 Iwatani (Hydro Edge) 11.3 Iwatani, built by Linde 10.0

Source: [20]

They can be integrated into thermodynamic processes in a manner similar to

isentropic expansion in expanders with pure gas processes. To accomplish a continuous

cooling process, an appropriate ferromagnetic material, whose Curie point temperature

is in the range of the cooling temperature, has to be cyclically magnetized and

21

demagnetized [42]. Pressure and specific volume of the gas process correspond to the

magnetic field intensity and the magnetization, respectively. Thus, from the

corresponding gas processes, the magnetic Carnot, Brayton, Ericson and Stirling cycles

can be derived [42]. Many of these types of systems and magnetic material have been

investigated, and it is demonstrated that there is a possibility for these systems to be

improved and combined with existing hydrogen liquefaction cycles [43–47].

In research by Kamiya et al. [45], a newly designed magnetic refrigerator is run

by the Carnot cycle and liquefies pre-cooled 20.28 K gas hydrogen, absorbing the latent

heat. This Carnot liquefaction system consists of magnetic materials, a superconducting

magnet and a heat switch. Its capacity is 12 kg LH2/day at 1.25 Hz.

Two-stage active magnetic regenerative refrigerator systems, which have been

investigated and presented, can be improved and used in the future for the liquefaction

of hydrogen [47].

2.7 Future Developments and Presentation of New Conceptual Designs

In research by Ohlig and Decker [11], some future developments with 40-60%

efficiency and about 7.5 kWh/kg/ LH2 energy consumption are recognized as follows:

1. One of the key areas of exergy loss is the recycle compressor of the refrigeration

loop, including its interstage and after-coolers. Improvements could be made by

shifting to the more efficient turbo-compressors with a higher number of stages

and frequent intercooling, hence bringing compression closer to the isothermal

optimum.

2. Precooling to 77 K (usually about 80 K) is achieved by use of liquid nitrogen

which is then released to the atmosphere. Energy for nitrogen generation from

the air is lost and sensible heat only partially used. A closed loop consisting of

a nitrogen re-liquefier increases investment costs by 20 to 30% but results in

significant energy savings of 10%.

3. Shifting towards systems including energy recovery solutions regarding

expansion turbines even though, due to complexity respective, solutions will

need to undergo extensive qualification procedures prior to approval as the

standard for industrial applications.

4. The optimization of the refrigeration loop and use of unproven concepts in this

field such as new refrigeration media.

22

In a study of Kuendig et. al. proposed pre-cooling by LNG instead of liquid

nitrogen. The authors suggested that efficiency would be improved from 10 to 4 kW

h/kg LH2, when compared to currently used liquefaction processes, such as the plant at

Leuna, which uses liquid nitrogen to precool, but with the compression at ambient

temperature. However, this method is only applicable to hydrogen production from

LNG, which would necessitate the location of the plant near a seaport [20].

Another conceptual design is the WE-NET project, which is a 300 TPD large-

scale process delivering LH2 at 1.06 bar from a feed stream of equal pressure [20]. The

plant is based on a pre-cooled Claude cycle and is similar to the plant in Ingolstadt but

with some modifications that lead to an increased efficiency of 46.2% and a specific

liquefaction power calculated to 8.5 kWh/kgLH2 [48].

Quack [49] produced a study in which the design is based on modern helium

liquefiers that are built with up to ten expansion turbines placed strategically in a cycle

to obtain optimal overall efficiency. Efficiencies obtainable by this concept are up to

60% and specific energy consumption is 5-7 kWh/kg LH2.

Another study by Kuz’menko et al. [50] proposed a helium refrigeration cycle

which showed higher efficiency than the Ingolstadt plant. Valenti and Macchi [51]

found an innovative, efficient and large hydrogen liquefier. It is a large-scale plant since

the production rate is 10 kg/s of L H2. The system utilizes four cascaded helium Joule–

Brayton cycles and has a reported efficiency of 47.73% [51].

Research conducted by Ratlamwala et al. [52] reports on hydrogen liquefaction

based on renewable energy (solar photovoltaic/thermal) based on the Linde-Hampson

23

cycle. It is schematically shown in

Figure 2.5. The efficiencies of some of these prototype plants, and some already

in existence, are compared in Table 2.2.

From the open literature, it can be seen that there are three fundamental ways to improve liquefaction efficiency. The first is to improve the ortho-para hydrogen conversion, thereby improving the energy carrying capacity of the final product. The second approach would be to improve the overall energy efficiency of the process through cogeneration or energy recovery. Finally, improving the overall cooling capacity of the system can improve the energy conversion efficiency. In this study, three methods (catalytic ortho-para conversion, integrating Organic Rankine Cycle, and introducing vortex tubes) are investigated to increase the energy conversion efficiency of hydrogen liquefaction. An introduction to there three inventions are provided below.

Figure 2.5 Basic scheme of hydrogen liquefaction process based on renewable energy (adapted from [52]).

24

2.8 Orthohydrogen and Parahydrogen

An understanding of the physical characteristics of hydrogen is essential to develop an

efficient hydrogen liquefaction method. This section explains how the isomeric forms

of hydrogen are relevant in its use as a fuel and how these states are affected by the

liquefaction process. Molecular hydrogen occurs in two isomeric forms depending on

the alignment of its two proton spins namely orthohydrogen which has parallel spin

alignment, and parahydrogen which has antiparallel spin alignment. Since these two

differ in the nuclear spin state, rather than in chemical structure, they are also referred

to as spin isomers. The existence of different hydrogen forms of molecular hydrogen

was first proposed in 1927 by Heisenberg and Hund [53]. The two forms are para- and

orthohydrogen [53]. However, Harteck and Bonhoeffer first synthesized pure

parahydrogen in the following year. Parahydrogen represents a lower energy state than

orthohydrogen although, due to thermal excitation, at room temperature and pressure,

hydrogen consists of around 75% ortho- and 25% parahydrogen [53]. However, at low

temperatures, in the hydrogen liquefaction process, there is a spontaneous increase in

parahydrogen content, accompanied by a release of energy of around 1091 J mol-1,

which is higher than the heat of vaporization of hydrogen (904 J mol-1) [53].

25

Table 2.2 Efficiencies of some conceptual plan

Source: [20]

The two states have different energy levels so that the content of each species

at equilibrium is temperature dependent. Moreover, the energy balance of this reaction

has important implications for hydrogen storage because spontaneous conversion inside

a hydrogen tank can cause significant hydrogen vaporization. At hydrogen's boiling

point of 20K, as shown in Figure 2.6, the equilibrium is shifted almost completely

towards parahydrogen [54].

26

Figure 2.6 Equilibrium composition as a function of temperature (adapted from [54])

In the ortho-state, electron spins are in the opposite direction to the ortho-state,

and in the same direction in the para-state when the temperature falls, as shown in

Figure 2.2. A number of catalysts can be used to accelerate the conversion from ortho-

state hydrogen to parahydrogen [55,56], allowing conversion of all the liquid hydrogen

to parahydrogen prior to storage. One of the challenges of this reaction is that it is

typically performed at low temperature, usually at 77 K, during hydrogen liquefaction,

which additionally decreases the reaction rate.

Figure 2.7 Spin isomers of molecular hydrogen (adapted from [57]).

It was shown as far back as 1949 that the activation of hydrogen gas by

transition metals is connected with the presence of unpaired d-electrons or holes in the

electronic d-band of the metal [58]. Therefore, the key component of any catalyst for

this reaction is a paramagnetic metal ion, typically a transition metal or a lanthanide.

Para Hydrogen Ortho Hydrogen

Proton

Electron

27

Most of the widely used catalysts are made of various metal oxides such as

ferric, chromic, cerium, neodymium, manganese (both supported [12] and

unsupported), activated carbon, different metal compounds (such as uranium and

nickel), rare-earth metals and various organo-metallic compounds. The catalytic

activity of these materials is directly dependent on the specific surface area of the

catalyst [59]. A recent investigation of interconversion kinetics using paramagnetic

complexions as catalysts revealed a direct correlation between the rate constants and

the concentration of the catalyst, and that second-order rate constants are related to the

magnetic moment of solvated metals and, in most cases, to the size of the ligand in the

complex. While the dependence on magnetic moment can be explained using Wigner's

theory [60], the size of the ligand has a greater effect on the second-order rate constants

than previously expected [56]. Furthermore, recent research has demonstrated the

ability of C60 fullerene to act as a catalyst for ortho-para conversion in liquid oxygen at

77 K [61]. After the removal of oxygen, enriched parahydrogen adsorbed at C60

fullerene is stable for many days (half-life of around 15 days in the absence of

paramagnetic catalyst), while the conversion rate back to orthohydrogen of this material

dissolved in organic solvent at room temperature was determined to be three orders of

magnitude lower than the conversion rate of parahydrogen dissolved in organic solvents

and not protected by a C60 fullerene shell [62–64].

Hydrogen with parahydrogen in excess of its natural 3:1 ratio is used to study

hydrogenation reactions because the resulting products exhibit hyperpolarized signals

in proton NMR spectra [65,66]. Parahydrogen and orthohydrogen also exhibit some

significant differences in properties, such as ideal gas specific heat. Parahydrogen

exhibits as much as 50% higher ideal gas isobaric specific heat than orthohydrogen in

the 65-320K region, leading to a significant difference in behaviour between it and

ordinary hydrogen, which is 75% orthohydrogen [67]. This leads to different equations

of state proposed for parahydrogen and normal hydrogen. Recently, Leachman et al.

formulated new fundamental equations of state for parahydrogen, normal hydrogen,

and orthohydrogen, with upper limitations of pressure and temperature of 2000 MPa

and 1000 K, respectively [68]. Based on this work, Lemmon et al. developed a truncated

virial equation for use in fuel consumption applications, using the normal hydrogen

equation of state, providing a correlation for density as a function of temperature and

pressure [69]. After understanding the behaviour of hydrogen at various states, the next

28

step is to study the literature available on hydrogen liquefaction. The various hydrogen

liquefaction process is described in the next section.

2.9 Catalyst Conversion of Ortho- to Parahydrogen

When hydrogen is condensed, the transition from ortho- to parahydrogen form is slow

at 1.14%/ hour [70]. For this reason, catalysts are used to speed up this transition. In

1933, a mechanism for this catalysis was proposed by Wigner [71], and it is still the

most widely used version even though other mechanisms have since been proposed

[70,71]. Wigner’s mechanism was refined by Kalckar and Teller [23]. In research

published in 1965 by Leffler [23] and Kasai et al. [72], two general mechanisms were

distinguished:

1. Magnetic: Wigner’s mechanism belongs to this group, which assumes that a

strongly inhomogeneous magnetic field decouples the proton spins and allows

the ortho-para transition to occur.

2. Dissociative: This mechanism assumes disassociation of hydrogen molecule

and rearrangement such that parahydrogen is obtained.

The most commonly used catalysts are Cr2O3, Fe-oxides and hydroxides [70]. In

[73], CrO3/SiO2-gel, chromic anhydride, nickel silica gel, FeNi alloys, activated

charcoal, Fe(OH)3, and Fe2O3 are reported as commonly used catalysts. In [73], the use

of LaFeO3 as a catalyst is proposed.

A study by Boeva et al. [74] reported the use of gold nanoparticles as a catalyst

at low temperatures. It seems that gold nanoparticles exhibit magnetic properties at low

temperatures.

Iron, Platinum and Nickel Cata1ysts were investigated for use in ortho-para

conversion of hydrogen in a study conducted by Emmett and Harknes [75], which also

examined the influence of temperature, pressure, time of contact and poisons.

The theory was also formulated as a case study rather than for general purposes.

One of them is published in an article by Ishii and Sugano [76], where the theoretical

framework for the conversion of ortho- to parahydrogen on magnetic surfaces was

presented. This involved two energy flow paths to the translational motion of the

hydrogen molecule and spin excitation of the substrate, which was then compared with

the experimental results of two catalysts: antiferromagnetic α-Cr2O3, and ferromagnetic

29

EuO. Three factors identified by the authors as important in ortho-para hydrogen

conversions in [62], namely Fermi contact interaction; the Steric effect; and Dynamical

quantum filtering have been investigated.

2.10 Organic Rankine Cycle (ORC)

The Organic Rankine Cycle (ORC) works with organic, high molecular mass working

fluid, which has the normal for having the fluid vapour stage change, or breaking point,

at a lower temperature than the water-steam stage change. The ORC is a potential

contender for integrated systems. A significant amount of research has been conducted

related to Rankine cycle performance in energy systems. Researchers have investigated

the ideal working liquid, and the ideal reinjection state of the liquid, the capacity of

cogeneration and the economic analysis of the systems.

Organic Rankine Cycles (ORCs) offer several advantages. A standout feature

amongst the most essential qualities of natural working liquids is their generally low

enthalpy drop through the turbine that causes a higher mass flow rate and reduces the

entire waste, and which therefore builds turbine adiabatic proficiency. Additionally,

superheated vapour at the turbine exit of an ORC cycle, stimulates avoiding dissolution,

permitting dependable activity and quick start-up [77,78].

2.11 Vortex Tube

Through broad research endeavours, it is recognized that there have been no recent

critical investigations into the conduct of a Vortex Tube (VT) using supercritical,

cryogenic hydrogen as its working liquid. Accordingly, there exists no approved

computational liquid element (CFD) model of a VT under these conditions. The earliest

study to report the phenomenon of energy separation in a VT was conducted by Ranque

[79]. Imperative exploratory examinations on VT parameters were led by researchers

such as Takahama [80], while hypothetical and expository depictions of the vitality

detachment, as well as temperature and speed profiles in a VT were provided by

Deissler and Perlmutter [81] and Ahlborn et al. [82]. Saidi and Yazdi [83] investigated

an exergy examination on a VT while Valipour [79] confirmed a trial model of a VT

icebox.

30

2.12 Closing Remarks

The review of existing literature shows that there is a need for moving away from fossil

fuels and hydrogen as a fuel has significant advantages. However, the efficiency of

hydrogen liquefaction plants is at present too low to make the process environmentally

and economically viable. Multiple attempts have been made by researchers to improve

the efficiency of hydrogen liquefaction systems. Some of these have resulted in large

scale hydrogen liquefaction plants while new methods and modifications are largely

available in laboratory settings. While there has been research on individual methods

for efficiency improvement, little data is available on comparing possible hydrogen

liquefaction systems. Such an approach would enable researchers to identify the system

that is likely to generate liquified hydrogen at the best possible efficiency. In this light,

the thesis investigates various hydrogen liquefaction systems with a focus on energy

and exergy efficiency. Such an investigation will help identify the system

configurations that are likely to have maximum energy and exergy efficiency. The

systems employing catalyst, ORC, and vortex tubes being simulated are described in

detail in the next chapter.

31

SYSTEM DESCRIPTION

Chapter 3 highlights the importance of understanding various hydrogen liquefaction

systems so that the process can be carried out at a greater efficiency. In that light, this

chapter examines various possible combinations and system configurations that may

generate better efficiencies. For this purpose, seven hydrogen liquefaction systems are

introduced, the first of which is based on a patent introduced by Schwartz et al. [84].

The remaining six are advanced hydrogen liquefaction systems that simulate large plant

size processes. The proposed integrated systems are described in detail in order to

demonstrate how they function. Assuming initial pressure to be atmospheric for the H2

feed and liquid H2 product as a saturated liquid, these systems were simulated in Aspen

Plus in two versions: the process is simulated once without the catalyst and again with

the catalyst. The energy systems for the production of liquid hydrogen, which are

discussed in this chapter, are modified from the basic patent in order to reach optimum

process. It is expected that these systems will meet the desired exergy objectives. Figure

3.1 illustrates the analyzed systems in this proposal.

Figure 3.1 Advanced hydrogen liquefaction systems considered for analysis.

These efficiency improvements provide a benefit and can lead to future

improvements in the design of a hydrogen liquefaction system. These systems increase

understanding of the intricacies of the process of hydrogen liquefaction using modelling

and a thorough examination of the ortho-para conversion process and their separation.

The process modelling techniques provide additional benefit to the public. The planning

and projections incorporate the effect of parahydrogen into process modelling software,

allowing more accurate modelling of the liquefaction processes compared to only using

Main Liquefaction

System

Advanced Liquefaction System (S1)

Without Catalyst (S1-

A)

With Catalyst (S1-B)

Advanced Liquefaction System with CO2 Orgainc Rankine Cycle (S2)


A)


Advanced Liquefaction System with Vortex tubes (S3)


A)


32

normal hydrogen. An additional simple method of reduction in liquefaction power

consumption has also been identified.

Figure 3.2 illustrates the main system simulated and analyzed to provide an

understanding of the actual patent and validate results against it. A brief description of

the system and hydrogen flow are provided below.

H2FEED (stream 1) enters mixer before travelling to compressor C1 where it is

compressed from atmospheric pressure to 3 bar. Stream 3 exits from compressor C1

and enters heat exchanger EX1 as a hot product. A stream of liquid CO2, serves as the

cold stream. The exchanger lowers the temperature of the stream and it exits at a

temperature of -20 oC as stream 4. Stream 4 further enters mixer M2. Stream 5 exits

M2 and enters compressor C2, where the gas is compressed to 20 bar. Stream 6 exits

C2 and enters heat exchanger EX2 as a hot product, where it is cooled until -23.15 oC.

Stream 7 exits EX2 and enters the first multi-heat exchanger HX1, as a hot feed.

The output stream is stream 8 with the temperature of -173.15oC, which enters splitter

D1. In D1 it is split on streams 9 and 15. Stream 9 enters mixer M3, from which exits

stream 10 and enter the second multi-heat exchanger HX2 as a hot feed. The output

stream is stream 11, which exists at -230 oC. Stream 11 further enters the third heat

exchanger HX3 as a hot feed. Stream 12 is an outlet stream on -253 oC that goes directly

to valve V1, where the pressure is decreased again to atmospheric. Exit of V1 is stream

13, which again enters HX3 and then exits it as a final product – stream H2LIQProduct,

a liquid H2 product at the atmospheric pressure and temperature of -253 oC.

Stream 15, after splitter D2, enters HX2 as a hot feed and exits at -230 oC to

enter D3 as stream 16. Stream 16 goes to splitter D3, where it is split on streams 17 and

18. Stream 17 enters turbo expansion compressor TE2, from which exits stream 29 as

a vapor on 20 bar. This stream goes directly to the HX3 as a hot feed, from where it

exits as stream 30 as a liquid on -253 oC. Stream 30 further enters mixer M5 and exits

as stream 31. Stream 30 enters now HX2 as a hot feed, from where it exits as stream 32

at -235 oC. Stream 32 is split on streams 33 and 39 in splitter D4. Stream 33 goes to the

absorption tank AD, where ortho H2 species are absorbed. They are routed as stream

34B and mixed in mixer M4. From M4 stream 26 enters HX1 as a hot feed. Output

33

stream 27 exits on -120 C and goes to mixer M1, where it is mixed together with

H2FEED.

Para species from AD tank are represented as stream 34 and it enters HX1 as a

cold stream. The output stream 35 at -23.15 oC enters compressors C3, where it is

compressed to 20 bar and exists as stream 36. Stream 36 is a hot feed for heat exchanger

EX3, where it is cooled to -20 oC and rerouted back to HX1 as stream 37 as a hot feed.

Stream 45 is a cooling medium in EX3, which is liquid CO2, which exits as stream 46.

Stream 37 is cooled up to -173.15 oC and exits as stream 38, which is mixed in mixer

M3 together with stream 9.

Stream 39 from splitter D4 enters HX1 as a cold stream and exits as stream 40

at 215.5 oC. It is then mixed with stream 4 in mixer M2. Stream 18 from splitter D3

enters HX3 as a hot feed. It is cooled until -253 oC and it exits as stream 23, which goes

to valve V2 to decrease the pressure to atmospheric and becomes a saturated vapor as

stream 20. It is then flashed in F1 and the vapor fraction, stream 21, is rerouted back to

HX1 as a cold feed. The output stream 24 exits as vapor at-230 oC and goes directly to

HX2 as a cold stream. It exits as stream 25 at 26 oC, and enters mixer M5 together with

stream 34B.

Stream 17 from splitter D2 goes to turbo expansion valve TE1, from where it

exits as stream 28 at 20 bar. Stream 28 then enters mixer M5, where it is mixed together

with stream 30. Liquid product of flash separation F1 is stream 22 and it enters HX3 as

a cold feed. The output stream enters again F1. The inlet to flash separator F2 is a stream

of N2 Liq, from where the vapor fraction is rerouted to HX2 as a cold feed. The output

stream enters F2 again at 26 oC. The liquid part after F1 enters HX1 as a cold feed, and

exits HX1 as a stream of N2Gas at 215.5 oC. The additional cold feed to HX1 is stream

N2LIQ, which exits as N2GAS.

3.1 Description of System 1

Two hydrogen liquefaction systems are proposed in this section. The base system, which

is considered as a reference (base), is shown in Figure 3.3 while Figure 3.4 displays the

first developed integrated system with added catalysts.

3.1.1 System 1A: Reference system without a catalyst

34

This system is based on the layout of the original system described in the previous

section. Figure 3.3 shows the modified base system. The base system that is analyzed,

this system has a hydrogen feed stream and a liquefier. Ortho-species of hydrogen the

hydrogen feed stream are converted to the para-species in higher and lower temperature

converters. An adsorption unit, between the higher and lower temperature catalytic

converters, adsorbs a portion of the ortho content of the feed stream. The adsorbed

portion is desorbed during regeneration of an adsorbent in the bed of the adsorption

unit. It is then re-circulated in the higher temperature catalytic converter to reduce the

degree to which the ortho-species are converted to the para-species in the lower

temperature catalytic converter and at lower temperatures.

The process starts with feeding hydrogen, containing both the ortho- and para-

species of hydrogen, to the system from outside and recycling it from inside the system.

The proportion is about 75% ortho-species and 25% para-species, which are the

approximate values at atmospheric condition.

3.1.2 System 1B: Reference system with a catalyst

In this simulation, three catalysts were added to the heat exchangers to speed up the

conversion and reach 90% para-conversion. The system layout otherwise remains the

same as that of the reference system except for the addition of catalysts. With the

addition, ortho-para conversion consumes a significant amount of refrigeration energy

because it requires cooling at low temperatures. Further improvements in ortho-para

conversion can lead to a significant reduction in power requirement. Figure 3.4 shows

the configuration with the added catalysts. The colored exchangers are two exchanger

with catalysts. The catalysts are expected to help in speeding up the liquefaction

process but on the other hand it is also expected that efficiency will decrease due to the

high-energy requirement to run and operate the catalyst. Changes made on this

configuration are significant for utilizing the developed system imitating the

commercial system for further improvements are needed to increase the overall

efficiency of the system and could lead to less CO¬2 emissions. This is configuration

is aimed to lead to improvements to: reach simplicity, system integration, high thermal

efficiency and quick thermal response, wide turn-down window, low emissions, and

better fuel flexibility in the same design.

35

Figure 3.2 The main systems schematic diagram

Legend:HX : Heat ExchangerAO : Adsorption Unit B : BoilerC : CompressorD : SplitM : Mixer/DividerTE : Turbo ExpanderF : FlashV : Valve

36

Figure 3.3 Schematic diagram for the reference system

37

Figure 3.4 Schematic diagram for the main system with reactor

38


In a further improvement to the efficiency of the system, three Organic Rankine

Cycles (ORCs) are added to the compressors. The ORC represents the most

commonly used low heat source temperature-based system.

3.2.1 System 2A: System with ORC and without a catalyst

System 2A explores the use of ORC in hydrogen liquefaction. While the basic system

remains the same, an ORC is integrated to improve compressor work gain. An ORC

system consists of the same components as a conventional steam power plant. However,

the working fluid is an organic component, which exhibits a lower boiling temperature

than water, allowing a reduction in the evaporating temperature. The selection of this

CO2 fluid in an ORC is mainly due to the nature of use in the cycle and, most

importantly, the maximum temperature of the cycle. Figure 3.6 shows a schematic

representation of the proposed integrated system, where an ORC provides the necessary

electricity.

The additions of the ORCs to the compressor gained work by a total of about

140 kJ in which the total system efficiency can be improved. Figure 3.5 illustrate the

ORCs that are added to the system.

(a)

(b)

(c)

Figure 3.5 Organic Rankine Cycles

For each ORC, the work outcome could make a significant positive impact on

the system’s overall efficiency and can be optimized for future improvements. In a

further improvement to the efficiency of the system, three ORCs are added to the

compressors. The ORC, which represents the most commonly used low heat source

temperature-based system, consists of the same components as a conventional steam

power plant. However, the working fluid is an organic component that exhibits a lower

M1

H2FEED

C1

2

EX1

3

M2

C2

4

9

HX1

46

M3

S4

14

17

V-1

HX2

N2LIQ

1

D2

D3

D519

20

C4

C5

30

32

21

31HX3

22

49

F1

V-2

23

24

S1

25

33

D1

M412 13

B627

51

28

26

AO

D4

34

44

35

PROD

N2GAS

29

43

EX2

S9

33

S3

36

E-51

E-52

S3

S6

E-53

S11

S5

46B

8

E-57

E-58

E-59

11

10

E-56

S8

S7

S12

Ex3

B13

E-61

41

40

B11

S13

S15

P-91

39

P-93

M1

H2FEED

C1

2

EX1

3

M2

C2

4

9

HX1

46

M3

S4

14

17

V-1

HX2

N2LIQ

1

D2

D3

D519

20

C4

C5

30

32

21

31HX3

22

49

F1

V-2

23

24

S1

25

33

D1

M412 13

B627

51

28

26

AO

D4

34

44

35

PROD

N2GAS

29

43

EX2

S9

33

S3

36

E-51

E-52

S3

S6

E-53

S11

S5

46B

8

E-57

E-58

E-59

11

10

E-56

S8

S7

S12

Ex3

B13

E-61

41

40

B11

S13

S15

P-91

39

P-93

M1

H2FEED

C1

2

EX1

3

M2

C2

4

9

HX1

46

M3

S4

14

17

V-1

HX2

N2LIQ

1

D2

D3

D519

20

C4

C5

30

32

21

31HX3

22

49

F1

V-2

23

24

S1

25

33

D1

M412 13

B627

51

28

26

AO

D4

34

44

35

PROD

N2GAS

29

43

EX2

S9

33

S3

36

E-51

E-52

S3

S6

E-53

S11

S5

46B

8

E-57

E-58

E-59

11

10

E-56

S8

S7

S12

Ex3

B13

E-61

41

40

B11

S13

S15

P-91

39

P-93

39

boiling temperature than water, allowing a reduction in the evaporating temperature.

The selection of this CO2 fluid in an ORC is mainly due to the nature of use in the cycle

and, most importantly, the maximum temperature of the cycle.

For each ORC, the work outcome could make a significant positive impact on

the system’s overall efficiency and can be optimized for future improvements. The use

of ORCs allows or a reduction in CO2 usage in exchangers EX1-3 from 350kg/sec to

5kg/sec.

3.2.2 System 2B: System with ORC and with a catalyst

For a catalyst-infused advanced hydrogen system, three ORCs have also been added.

The importance of system relay on the catalyst beds in the heat exchangers that have

been added to speed up the process of liquefaction. The objective of the ORCs is to

successfully utilize the outlet temperature. This also limits the emissions that will be

reflected in the environmental analysis by the reduction of CO2 emissions per MW of

energy. The colored exchangers are two exchangers with catalysts. The catalysts are

expected to help in speeding up the liquefaction process but on the other hand it is also

expected that efficiency will decrease due to the high-energy requirement to run and

operate the catalyst but with ORC its can be said that ORCs will compensate on the

high emissions. Changes made on this configuration are significant for utilizing the

developed system imitating the commercial system for further improvements are

needed to increase the overall efficiency of the system and could lead to less CO2

emissions. This is configuration is aimed to lead to improvements to: reach simplicity,

system integration, high thermal efficiency and quick thermal response, wide turn-

down window, low emissions, and better fuel flexibility in the same design.

This has shown interesting results and was analyzed thoroughly in the next

chapter to maximize the efficiency with keeping in mind other factors within the limits

and the boundaries of the optimization models.

40

Figure 3.6 Schematic diagram for the system with ORCs

41

Figure 3.7 Schematic diagram for the system ORCs and reactor


In a further improvement to the efficiency of the system, Vortex tubes before the

exchanger inlet. Two configurations are explained further.

42

3.3.1 System 3A: System with Vortex Tubes and without a catalyst

In this system, Vortex Tubes (VTs) were added using a splitter and a turbo expander to

simulate the splitting and cooling effects of the VTs. Apart from the addition of VTs,

the system layout corresponds to that of the base system. The split was simulated to be

at 50% and the turbo expander outlet pressure was simulated to be 15 bars. Adding a

VT in this manner resulted in additional cooling capacity in the system. There is an

optimization opportunity to split more of the stream through to the VT, and to further

reduce the outlet pressure on the expander to increase cooling capacity. Figure 3.9

shows the VTs added to the main system.

Figure 3.8 Added VTs

3.3.2 System 3B: System with Vortex Tubes and with a catalyst

In this system, Vortex Tubes (VTs) were added using a splitter and turbo expander to

simulate the splitting and cooling effects of the VT. The split was simulated to be at

50% and the turbo expander outlet pressure was simulated to be 15 bars. Adding a VT

in this manner resulted in additional cooling capacity in the system. There is an

optimization opportunity to split more of the stream through to the VT, and to further

reduce the outlet pressure on the expander to increase cooling capacity. The use of the

reactors puts additional energy requirement on the system. This energy requirement is

reduced by using the VTs.

The next step of this study is to investigate the energy and exergy efficiencies

of the six proposed systems. Systems 1A and 1B are the base configurations with and

without catalyst addition. Systems 2A and 2B are modifications of the same systems

with the addition on ORC while systems 3A and 3B include vortex tubes.

It is hoped that this analysis of varying configurations will yield insights

regarding methods to improve efficiency of hydrogen liquefaction. The next chapter

therefore presents energy and exergy analysis of each configuration for comparison.

M1

H2FEED

C1

2

EX1

3

M2

C2

4

9HX1

46

M3

S4

14

17

V-1

HX2

N2LIQ

1

D2

D3

D519

20

C4

C5

30

32

21

31HX3

22

49

F1

V-2

23

24

S1

25

33

D1

M412

13

27

51

28

26

AO

D4

34

44

35

PROD

N2GAS

29

43

EX2

S9

33

S3

36

VT1

VT2

P-78

text 51 44

78 32

43

Figure 3.9 Schematic diagram for the system with VTs and no reactor

44

Figure 3.10 Schematic diagram for the system with VTs and reactor

45

SYSTEM ANALYSIS, MODELLING AND

SIMULATION

The thermodynamics analyses of the proposed systems will be based on energetic and

exergetic methods. Exergoeconomic concepts will be utilized to examine the created

frameworks economically. The performances of the proposed systems will be assessed

by deciding the energy and exergy efficiencies for the presented frameworks. In this

chapter, basic equations of energy and exergy will be presented. The investigation of

the principle controlling choices will be portrayed.

4.1 Basic Thermodynamic Concepts

In thermodynamic analyses, overall mass, energy, entropy and exergy balance

equations are written for the fuel that will be blended with ammonia. In order to

understand the combustion process, how the generator operates and its performance, a

comprehensive thermodynamic analysis is carried out regarding energy and exergy

studies to evaluate efficiencies of the system. The general assumptions taken in to

account for the thermodynamic analysis and calculations of the system can listed as:

• The reference temperature is selected as T0 = 25 oC (outside temperature during the

experimental studies) and reference pressure P0 = 101.325 kPa.

• The variations in the kinetic and the potential energies and exergies are ignored.

• The ideal gas laws apply for the gases operating in the system.

• Air used in the system is an ideal gas with constant specific heat.

• The relative humidity of the inlet air and hydrogen is taken as 90%.

• No pressure drops in the system

• No heat losses in pipes nor other equipment (perfectly insulated)

• The kinetic and potential energy changes are negligible.

• The liquefaction capacity is 36000 kg/day.

• Compressors are assumed to be isentropic.

Thermodynamic systems are, in principle, open or closed in nature [85]. Open

systems interact with the environment, exchanging heat, mass and work, while closed

systems include only exchange heat and work. When the system is defined as positive,

the Mass flow and the transfer of heat goes into, and the transfer of work goes out of

46

the system. During the hydrogen liquefaction process, there is a constant flow of

hydrogen mass in and out of the system; a hydrogen liquefaction system would

represent an open system.

4.2 Conservation of Mass Principle

For a non-steady flow process, occurring during a time interval from t1 to t2, the balance

of mass for this process can be written as follows:

∑ 𝑚; −∑ 𝑚> = 𝑚@ −𝑚A>; (4.1)

where mi and mo denote the mass entering the system through the input and exiting the

system through the output in time (t2-t1), respectively. In a more generalized form, for

an infinitesimally period dt, the conservation of mass in the system can be described as:

&BC&D

= ∑ ��; −; ∑ ��>> (4.2)

where ṁ represents mass flow (dm/dt), with i and o denoting the input and output of

the system, respectively; mv denotes the mass inside the system volume V.

4.3 Conservation of Energy Principle

The balance of energy for a non-steady flow process can be written as:

∑ (𝑒 + 𝑃𝑣);𝑚;; −∑ (𝑒 + 𝑃𝑣)>𝑚> +∑ (𝑄K)A,@ −(𝑊)A,@K> = 𝐸@ −𝐸A (4.3)

where e, P, v and m represent specific energy, pressure, specific volume and mass, i

and o denote input and output of the system, (Qr)1,2 denotes the heat transferred into the

system volume across the region r, (W)1,2 denotes the work transferred out of the system

and E1 and E2 represent the energy of the system at time t1 and t2, respectively. For a

hydrogen liquefier, which is a steady state system, where the energy of the system is a

constant and infinitesimally small period of time, this equation can be rewritten as:

��N + ��N + ∑ ��;ℎ; = ��P + ��P + ∑ ��> ℎ> (4.4)

where �� and �� represent the heat transfer and work rate exchanged between the system

and its environment, and ṁ and h represent mass flow rate and the specific enthalpy of

the streams.

4.4 Entropy Balance and Entropy Generation

Entropy generation occurring during the processes these systems undergo can be

described using the following equation:

47

&QRC&D

= ∑ ��;-𝑠;- − ∑ ��>TD𝑠>TD + ∑URCV+ ��+,- (4.5)

where s denotes specific entropy and ��+,- represents the rate of entropy generation.

4.5 Exergy Analysis

Exergy analysis is a technique of thermodynamic analysis formed based on the Second

Law of Thermodynamics. It gives alternative means of assessing and comparing

different systems and processes in a meaningful way, yielding efficiencies that

represent a factual representation of how close the performance of a given system

comes to an ideal form, and allows us to identify the causes and locations of

thermodynamic loses more effectively than with the energy analysis [85]. In this regard,

exergy analysis can help to improve and optimize system designs. Exergy

corresponding to a particular quantity of energy represents a quantitative assessment of

its usefulness, recognizing that, while energy cannot be created or destroyed, its quality

and value can be degraded. For an energy storage system, which liquid hydrogen

essentially is, exergy analysis determines the maximum potential of the incoming

energy. Liquid hydrogen is preserved and recovered only if the process of storing

energy is fully reversible. Since this can never occur under realistic working conditions,

where processes are always irreversible, the process of energy storage represents a

source of loss in the system’s potential for exergy recovery. This means that exergy

analysis quantitatively specifies more practical boundaries by providing the

information on losses as indicated by lost exergy.

For a steady state system, the exergy balance equation considering the system

components can be constructed in the following general form [86]:

&WXY&D

= ∑ ��𝑥U − ∑ ��𝑥Z + ∑ ��𝑥[\>Z; − ∑ ��𝑥[\>Z> − 𝐸��& (4.6)

where ��𝑥U denotes the rate of exergy transfer with the heat energy exchange across the

system volume. ��𝑥Z denotes the rate of exergy transfer by the boundary or work

applied on or done by the system. The term ��𝑥[\>Z represents the exergy transfer rate

with flow transfer through the system. The exergy destruction, which describes the

system irreversibility, is shown in the equation as 𝐸��&.

The exergy transfer due to the exchange of heat with the environment can be

formulated as follows:

48

Ex]^ = QN `1 −bcbd^e (4.7)

where To is the environmental temperature and TS is the temperature of the source (for

heat penetration process) or sink (for a heat loss process).

The exergy transfer rate associated with work, taking into consideration the

potential change of the volume of the system as a result of this work, can be written as

follows:

��𝑥Z = ��f + 𝑃g&hC&D

&BC&D

= ∑ ��; − ∑ ��> (4.8)

where m and �� denote mass and mass flow rate, respectively. The subscripts v, i and o

indicate the control volume and the inlet and exit of the control volume, respectively.

where P0 is the pressure of the system in a dead state.

Exergy of a flowing stream of gas or liquid can be represented as a sum of the

different exergies (chemical, physical, kinetic and/or potential) of the flow:

𝐸𝑥i>D = 𝐸i>D (4.9)

𝐸𝑥j;- = 𝐸j;- (4.10)

𝐸𝑥kl,B = 𝐸x = ∑ m𝜇; − 𝜇,op𝑁;; (4.11)

𝐸𝑥[\>Z = (𝐻; − 𝐻g) −𝑇g(𝑆; − 𝑆g) (4.12)

where Epot and Ekin are the potentials and kinetic energy, respectively, μ is the chemical

potential (i denotes environmental state, eq denotes equilibrium state), T0 is the

temperature, Ni is the component mole fraction and H and S represent the enthalpy and

the entropy of the system, respectively. The kinetic and potential terms of the flow

exergy can be disregarded as negligible since the changes in velocities and elevation

across the system components are too small compared to the values of other two terms.

Therefore, the flow exergy term can be written like this:

𝑒𝑥; = ℎ; − ℎg − 𝑇g(𝑠; − 𝑠g) + 𝑒𝑥 (4.13)

where h and s denote the enthalpy and the entropy, respectively, while ex-denotes

chemical exergy term.

The exergy destruction rate can be calculated according to the following:

49

𝐸��&t = 𝑇g��+,-,; (4.14)

whereSvwx,Ndenotes the rate of entropy generation for the system component i, which

can be correlated to the entropy balance equation for a steady state operation of each

system component:

��+,-,; = ∑ ��> 𝑠> − ∑ ��;𝑠; −∑(UV) (4.15)

For each individual system component, the corresponding exergy balance

equations and exergy efficiency, including those of individual subsystems. The exergy

efficiency for a particular process can be expressed through the ratio of exergy output

produced by the system to the total exergy input.

From the above, the exergy balance equation can be reformulated as follows:

&WXC&D

= ��𝑥f = ∑ y1 − VcVz �� − ��f + 𝑃g

&hC&D+ ∑ ��;𝑒𝑥; − ∑ ��>𝑒𝑥> − 𝑇g��+,-(4.16)

Exergy destruction in each component can be determined to utilize the exergy

balance on the system components at a steady state, as follows:

𝐸��&; = 𝐸𝑥U{ − 𝐸��|t + ∑ ��;𝑒𝑥; − ∑ ��> 𝑒𝑥> (4.17)

where Ex}N denotes the rate of exergy destruction occurring at the system component i,

Ex~^ and Ex]^ represent the exergy rates corresponding to work and heat transfer,

respectively, across the system limitations, while the exergy rates carried in and out of

the system with the flow are represented by exN, exP.

4.6 Components used in the systems

System components that are utilized in the different systems can be analyzed

energetically and exergetically in Table 4.1. Some components included in the table are

pressure regulator, expander, and compressor, for each component mass, energy,

entropy and exergy balance equations are listed.

50

Table 4.1. Energy Balance for System Components

Component Component Name

Balance Equation

Mixer Energy Balance ��AℎA + ��@�ℎ@� = ��@ℎ@ Entropy Balance ��A𝑠A + ��@�𝑠@� + ��+,-,BA = ��@𝑠@ Exergy Balance ��A𝑒𝑥A + ��@�𝑒𝑥@� = ��@𝑒𝑥@ + 𝐸��&,BA

Compressor Energy Balance ��@ℎ@ + ��kA = ��ℎ� Entropy Balance ��@ℎ@ + ��+,-,kA = ��ℎ� Exergy Balance ��@𝑒𝑥@ + ��kA = ��𝑒𝑥� + 𝐸��&,kA

Cooler, Heat Exchanger (Without heat exchanger efficiency)

Energy Balance ��ℎ� + ��ℎ� = ��ℎ� + ��ℎ� Entropy Balance ��𝑠� + ��𝑠� + ��+,-,,XA = ��𝑠� + ��𝑠� Exergy Balance ��𝑒𝑥� + ��𝑒𝑥� = ��𝑒𝑥� + ��𝑒𝑥� + 𝐸��&,,XA

Cooler, Heat Exchanger (With heat exchanger efficiency)

Energy Balance ��ℎ� + ��ℎ� = ��ℎ� + ��ℎ� + ��\>��,,XA Entropy Balance ��𝑠� + ��𝑠� + ��+,-,,XA = ��𝑠� + ��𝑠� +U��,��𝑻𝒃𝒐𝒖𝒏𝒅𝒓𝒚

𝑻𝒃𝒐𝒖𝒏𝒅𝒓𝒚𝒐𝒓𝑻𝒔𝒖𝒓𝒇𝒂𝒄𝒆𝒐𝒓𝑻𝟎𝒐𝒓𝑻𝒂𝒎𝒃𝒊𝒆𝒏𝒕 Exergy Balance ��𝑒𝑥� + ��𝑒𝑥� = ��𝑒𝑥� + ��𝑒𝑥� +𝐸��&,,XA + 𝐸��U��,�� 𝐸��U��,�� = y1 − Vc

V��z ��\>��,,XA𝑎𝑛𝑑𝐸��&,,XA = 𝑇g��+,-,,XA

𝐹𝑜𝑟𝑇,XA𝑚𝑎𝑘𝑒𝑎𝑛𝑎𝑠𝑠𝑢𝑚𝑝𝑡𝑖𝑜𝑛𝑑𝑒𝑝𝑒𝑛𝑑𝑖𝑛𝑔 𝑜𝑛𝑠𝑡𝑎𝑡𝑒𝑝𝑜𝑖𝑛𝑡𝑡𝑒𝑚𝑝𝑒𝑟𝑎𝑡𝑢𝑟𝑒𝑠.

Valve Energy Balance ��@�ℎ@� = ��@�ℎ@� Entropy Balance ��@�𝑠@� + ��+,-,f@ = ��@�ℎ@� Exergy Balance ��@�𝑒𝑥@� = ��@�𝑒𝑥@� + 𝐸��&,f@

M

2

1

3

C1

2

HX4

5

6

1

2

3

HX4

5

6

1

2

3

V

12

51

Flash Energy Balance ��@�ℎ@� + ��@ℎ�@ = ��@�ℎ@� + ��@ℎ�@ Entropy Balance ��@�𝑠@� + ��@𝑠�@ + ��+,-,[A = ��@�𝑠@� + ��@𝑠�@ Exergy Balance ��@�𝑒𝑥@� + ��@𝑒𝑥�@ = ��@�𝑒𝑥@� + ��@𝑒𝑥�@ +𝐸��&,[A

Vortex Tube Energy Balance ��@ℎ@ + ��kA = ��ℎ� Entropy Balance ��@ℎ@ + ��+,-,kA = ��ℎ� Exergy Balance ��@𝑒𝑥@ + ��kA = ��𝑒𝑥� + 𝐸��&,kA

4.7 Energy and Exergy Efficiencies

The exergy efficiency for the liquefaction process of the system is expressed as follows:

Ψ = WX¦t§|¨�©ª«��«

(4.18)

For each component exergy and exergy efficiency, in Table 4.2, Exergy

destruction and exergy efficiency are presented.

Efficiency can be defined as "the ability to produce the desired effect without

waste of, or with minimum use of, energy, time, resources, etc.,”. It is typically used in

the context of effectiveness which something is produced from something else, or how

close to the ideal situation the system is in performing a given task [85]. In engineering

systems, efficiency is usually expressed through nondimensional ratios of quantities,

like the energy in systems for transformation of energy. Energy efficiency, formulated

in this manner, is based on the First Law of Thermodynamics, where it can be stated

that the maximum efficiency is achieved if the input of energy equals the recoverable

output of energy. However, efficiency determined in this form does not mean its a true

measure of mimicking the ideal situation. A more meaningful efficiency can be

determined using exergy: through a formulation that the maximum efficiency can only

be achieved through a fully reversible process. This can be quantified through entropy,

where maximum efficiency can be achieved in through conservation of entropy, and,

correspondingly, the magnitude of the creation of entropy can be utilized as the measure

F

2

1

3

VT

12

52

of the degree to which the process is irreversible. However, using the entropy ratios

would not provide the measure of how close the process is to the ideal.

If maximum efficiency is defined to have been attained when, at the end of a

process, the sum of all energy in the process has the ability to perform work equal to

the sum before the process occurred, then exergy, as the measure of the ability to

perform work, has to be conserved in a process with maximum efficiency [67]. This

approach provides a measure of an approach of the system to the ideal. Another

advantage of exergy efficiency is that the values between 0 and 100% are always

obtained, unlike for energy consideration, where factors like coefficient of performance

can sometimes have values greater 100%. The energy (η) and the exergy (ψ)

efficiencies can be written as:

𝜂 = W�Wt

(4.19)

𝜓 = WX�WXt

= 1 − ∑WX��WXt

(4.20)

where i and o denote input and output of the system, and E and Ex denote energy and

exergy of the system, respectively. Exergy destruction ratio, or depletion factor Dp, can

be defined as the ratio of exergy destruction rate to the input exergy rate:

𝐷i =WX

WXt (4.21)

There are two other, commonly used exergy-based equations for the efficiency

of a steady-state device:

𝑅𝑎𝑡𝑖𝑜𝑛𝑎𝑙𝑒𝑓𝑓𝑖𝑐𝑖𝑒𝑛𝑐𝑦 = 1 −WX,K+±k>-�TBiD;>-V>D²\,X,K+±;-iTD

(4.22)

𝑇𝑎𝑠𝑘𝑒𝑓𝑓𝑖𝑐𝑖𝑒𝑛𝑐𝑦 = Vl,>K,D;k²\B;-;BTB,X,K+±;-iTD³,oT;K,&´kDT²\,X,K+±;-iTD

(4.23)

Exergy efficiencies weigh individual energy flows in the system, taking into

consideration their respective exergy contents, and separate inefficiencies into two

groups according to the cause: those originating from effluent losses and those

originating from irreversibilities in the system. This makes the information about

inefficiencies more useful, because they provide a quantitative measure of the system’s

potential for improved efficiency.

The yield calculation is based on the fraction of liquefaction:

53

𝑌 = B¶

B= l�·l¸

l�·l¶ (4.24)

where 𝑚[ is the mass flow rate of the fraction of liquefaction. The total system

efficiency is determined by:

𝜂@-&(%) = ml¶·l�p·Vc.(�¶·��¸)

Zº�º.B¶. 100 (4.25)

The energy per unit mass liquefied in the system is calculated by:

𝑤; = |B¶

= |B½

(4.26)

For the cooling cycle, the Coefficient of performance (COP) is determined by:

𝐶𝑂𝑃 = ml¶·l�pZº�ºÀ�

(4.27)

4.8 Sustainability Assessment

Exergy analysis, as a whole, evaluates the quality of the underlying thermodynamic

processes and serves as a potential tool for achieving maximum sustainability.

However, when considering a particular process in the real world, there are two main

aspects to be considered: its cost or efficiency and its environmental impact. Increase

in exergy efficiency of a particular system serves both to minimize the destructed

exergy (and reduce waste) and decrease the environmental impact of the process,

making exergoeconomic and exergoenvironmental analysis invaluable tools for

achieving sustainable development.

Another parameter is the sustainability index SI [87]:

𝑆𝐼 = A[ (4.28)

which represents a measure of the effectiveness of the process: higher sustainability

index means that less exergy is destroyed during the process as a portion of the input

exergy. This also means that the process with a higher sustainability index would have

a lower environmental impact.

54

4.9 Exergoeconomic Assessment

The exergoeconomic analysis represents a combination of exergy analysis and

economic analysis of a process. The goal is to help achieve maximum exergy system

optimization by determining the trade-off between the input cost (like the cost of fuel),

and capital and production cost. It uses a combination of thermodynamic and economic

principles to describe the system at the individual component level, providing

information for design improvement and cost-effective operation. The exergy part of

the model considers the exergy efficiencies at each point of the system, while the

economic part of the model takes into account the costs associated with the capital,

operation and maintenance of the system to determine the flow of costs in the system

and provide optimization of both specific variables of a particular component and the

system as a whole.

Exergy balance equation can be rewritten to include the cost associated with

each term of exergy flow:

𝐸��|t𝐶| = 𝐸𝑥U{ 𝐶V − 𝐸��&;𝐶Q + m∑ ��;𝑒𝑥; − ∑ ��> 𝑒𝑥>p𝐶g +𝑍j (4.29)

where CW is the unit cost of the rate of work production, while CT, CS, and C0 represent

unit costs corresponding to the thermal exergy flow, exergy rate of destruction and

chemical exergy flow, respectively. The term 𝑍j includes all financial costs correlated

with ownership, operation and maintenance of any particular system component, and

can be defined as [62]:

𝑍j = ÂÃÄÃ

��gg×ÆÇ (4.30)

where ϕk is maintenance factor (typically equal to 1.06), ��j is the unit cost rate and Nh

is the annual number of operational hours of the system component.

The cost balance equation can also be written in terms of unit cost rates for each

individual exergy flow component:

𝐶| = 𝐶V −𝐶Q +𝐶g + ��j (4.31)

where different �� parameters represent the corresponding unit cost rates: �� = 𝐶𝐸𝑥.

Application of exergy cost balance equation for every system component yields a set

of non-linear equations which can be solved for �� or C.

55

Table 4.2 Base system components exergy equations

Component Exergy Destruction Rate Exergy Efficiency Cooler 1 Exd,EX1 = m3ex3 + m5ex5

− m4ex4− m6ex6

ηex,EX1 =(m4ex4 − m6ex6)(m3ex3 − m5ex5)

Cooler 2 Exd,EX2 = m8ex8 + m10ex10− m9ex9− m11ex11


Cooler 3 Exd,EX3 = m38ex38 + m40ex40− m39ex39− m41ex41


Expansion Valve 1 Exd,V1 = m18ex18 − m16ex16 ηex,V1 =

(m16ex16)(m18ex18)

Expansion Valve 2 Exd,V2 = m23ex23 − m24ex24 ηex,V2 =

(m24ex24)(m23ex23)

Compressor 1 Exd,C1 = m3ex3 + Win,C1− m2ex2

ηex,C1 =(m3ex3 − m2ex2)

Win,C1


ηex,C2 =(m8ex8 − m7ex7)

Win,C2


ηex,C3 =(m38ex38 − m37ex37)

Win,C3

Turbo Expander 1 Exd,T1 = m31ex31 − m30ex30− Wout,T1 ηex,T1 =

Wout,GT1

(m31ex31 − m30ex30)

Turbo Expander 2 Exd,T2 = m32ex32 − m21ex21− Wout,T2 ηex,T2 =

Wout,GT2

(m32ex32 − m21ex21)=

Flash Drum 1 Exd,F1 = ms1exs1 + m25ex25−ms2exs2− m24ex24− Wout,F1

ηex,F1 =Wout,ST

(ms2exs2 − m24ex24)

Flash Drum 2 Exd,F2 = ms4exs4 + ms5exs5−ms3exs3− ms6exs6− Wout,F2

ηex,F2 =Wout,ST

(m56ex56 − m54ex54)

Heat Exchanger 1 Exd,HX1 = m1ex1 + m29ex29 +m37ex37 + m42ex42 + m46ex46+mN2gasexN2gas + ms5exs5 −m9ex9− m28ex28 −m36ex36 + m39ex39 −m44ex44− mN2liqexN2liq − ms7exs7

Heat Exchanger 2 𝐸��𝑑,𝐻𝑋2 = ��15𝑒𝑥15 + ��19𝑒𝑥19 +��27𝑒𝑥27 + ��34𝑒𝑥34

+��𝑠4𝑒𝑥𝑠4 −��6𝑒𝑥6 − ��13𝑒𝑥13 −��14𝑒𝑥14+ ��26𝑒𝑥26 −��49𝑒𝑥49

Heat Exchanger 3 Exd,HX3 = m18ex18 + m23ex23 +m26ex26 + m33ex33+mH2liqproexH2liqpro + ms2exs2 −m15ex15− m16ex16 −m22ex22 − m25ex25 −m32ex32−ms1exs1

56

The ratio of exergy loss to the capital cost of the system provides information

about the relative exergy loss (waste plus destruction) in the system compared to its

capital cost:

𝑅 = �� ⁄ (4.32)

where �� represents the overall financial cost of the system and �� is the total exergy loss

in the system defined as:

�� = 𝐸��& + 𝐸��\>�� (4.33)

where 𝐸��\>�� represents the sum of exergy losses due to heat exergy transfer and flow

exergy leaving the system.

Finally, exergoeconomic factor for the component k can be defined as the ratio

of the capital cost of the system to the sum of the capital cost and exergy loss of the

system [88]:

𝑓,,j = ëÃ

ëÃìÄí,ÃîÃ= ëÃ

ëÃìÄí,Ã(WXìWX��) (4.34)

where CF,k is the unit cost of the exergy of the fuel expended by the component k. A

higher value of the,k means that the process is more efficient because the cost of lost

exergy is smaller compared to the capital cost of the system.

4.10 Environmental Impact Assessment

Environmental problems and issues have become a major factor in the adoption of new

technologies and construction of new production facilities due to increased public

awareness over the past few decades. These include an increasing number of pollutants

and ecosystem deterioration factors affecting the environment at the local, regional and

global level. Since these problems are often complex and in a state of constant flux,

where industrial and technological development often creates new environmental

problems, requiring analysis of environmental impact before their introduction.

Since hydrogen liquefaction is a thermal process, involving no exhaust gas or

another pollutant, the most important environmental factor stems from its energy

consumption. Since a large amount of energy is obtained by burning fossil fuels, it has

to be taken into consideration that any waste in the system would produce additional

greenhouse gas emissions. The amount of greenhouse gas emission from any thermal

process can be calculated using thermodynamic analysis and then compared to the

57

similar systems to evaluate its environmental impact relative to the technology currently

in use. The environmental impact of a process can be correlated directly with the rate

of exergy consumption or destruction through environmental impact factor f [87]:

𝑓 = ïªð

(4.35)

which represents the efficiency of exergy depletion, i.e. what portion of exergy flow in

the system is destroyed.

4.11 Optimization Study

Cost accounting is concerned with calculating the real cost of production, providing a

rational economic basis for pricing, providing means for allocation and control of

expenditure and providing information for making an evaluation of operating decisions.

Since cost can be defined, in the broadest of terms, as the number of resources necessary

to obtain a functional product, that amount of resources can be expressed through

exergy, allowing us to exercise thermoeconomic optimization of the system, with the

goal of minimizing the cost associated with exergy flow. In order to do this, it is

necessary to choose the proper equipment, in type and size, and the best configuration

of the system, along with optimal operating pressure and temperature ranges for the

process. In calculating the approximate average cost of thermal processes, it is

recommended to use the lowest possible aggregation level, since the sophistication of

the formulation of the cost balances has a considerable effect on the results of any

thermoeconomic analysis. This level is typically represented by the individual system

components, even in cases where all of the data is not available, because it is generally

preferable to make appropriate assumptions about the exergy cost of each individual

component, rather than to consider only a group of components as a whole [85].

The first step in the definition of the optimization problem is to clearly define

the system boundaries in such a way to include all the important components and

system parameters. Secondly, the selection of criteria that will form the basis of the

system's evaluation and optimization represents the key element of any optimization

[89]. The optimization criteria can be divided into three groups:

- Economic (capital investment, cost of operation and maintenance, net profit)

- Technological (efficiency, production time and rate, fuel consumption)

- Environmental (rate of pollutant emission)

58

Optimized system design is characterized by a minimum and maximum value

for each of the selected criteria [87]. A third essential element in system optimization

is the selection of design variables which characterize the available design options,

taking care to include all the possible factors which affect the efficiency and cost

effectiveness of the system. These can be independent or design variables, dependent

variables and parameters.

A mathematical model for solving the optimization problem consists of an

objective function that is to be minimized, a set of equality constraints and a set of

inequality constraints. The objective function, in this case, can be the maximization of

exergy efficiency or minimization of exergy loss or destruction or minimization of the

product cost. Equality constraints stem from appropriate thermodynamic and economic

models with properly defined system boundaries. Inequality constraints usually specify

the allowed operating ranges, minimum and maximum performance required and limits

on the available resources. Thermoeconomic optimization methods typically use a

primary optimization performance measure: minimizing the total cost of the system

product, although multicriteria optimization and environmental factors may also be

considered. The differences between the systems include is the amount of para species

in the liquid product. Without a catalyst, there were no improvements in the transition

of ortho to para species. The catalyst system provided up to 90% of parahydrogen

species in total, which is good as parahydrogen provides stability of liquid hydrogen.

On the other hand, the total liquefaction yield (total liquid product per total hydrogen

input) is the same for both systems, due to the fact the amount of liquid product was the

same in both simulations. It is assumed that is because less energy was needed in the

main heat exchangers. The catalyst system provided up to 90% of parahydrogen species

in total, which is good as parahydrogen provides stability of liquid hydrogen.

Simulation outcome and parameter changes have been utilized for the parametric study.

For Ortho-Para reactor conversion of the Hydrogen in the catalyst-based

systems The o-p conversion reaction is set as the following for the reactors:

𝐻𝑦𝑑𝑟𝑜𝑔𝑒𝑛 ⟶ 𝑝 − 𝐻𝑦𝑑𝑟𝑜𝑔𝑒𝑛 + ℎ𝑒𝑎𝑡 (4.36)

Conversion percentage for the equilibrium concentration of ortho and para hydrogen

is temperature dependent the and follows:

𝐶𝑜𝑛𝑣𝑒𝑟𝑠𝑖𝑜𝑛(%) = 𝐶g + 𝐶A ∗ 𝑇 + 𝐶@ ∗ 𝑇@ (4.37)

59

where C0, C1, and C2 denote conversion confidents and T denotes the temperature in

kelvin. Yet, the coefficients are adjusted based on the composition of the outgoing

streams of the conversion reactors meets the experimental data. Properties of Hydrogen,

Nitrogen and Carbon dioxide used in the simulation shown in the Table 4.3.

Table 4.3. Properties of Hydrogen, Nitrogen and Carbon dioxide

Parameters Units Component

H2 LH2 N2 CO2

API Gravity 340 1860.95 340 340

Freeze Point C -259.2 -259.347 -210.001 -56.57

Std. Enthalpy cal/mol -57757.7 -57772 0 0

Heat of Fusion cal/mol 27.9689 28.0644 171.969 2154.15

Molecular Weight 2.01588 2.01588 28.0135 44.0098

Pitzer Acentric Factor -0.215993

-0.220755 0.0377215 0.223621

Critical Pressure bar 13.13 12.928 34 73.83

Some of Entropies cal/mol-K

31.2124 45.765 50.3696

Specific Gravity 0.3 0.0710179 0.3 0.3

Boiling Point C -252.76 -252.882 -195.806 -78.45

Critical Temp. C -239.96 -240.174 -146.95 31.06

Triple Point Temp C -259.2 -259.347 -210.001 -56.57

Liquid Molar Vol. cc/mol 28.5681 28.4572 34.6723 35.0189

Critical Vol. cc/mol 64.147 64.144 89.21 94

Standard Liquid Molar Vol. cc/mol 53.5578 28.4572 53.5578 53.5578

Critical Compressory Factor 0.305 0.302 0.289 0.274

60

RESULTS AND DISCUSSION

This chapter discusses the performance details of the proposed systems. This includes the

base system, and results obtained from systems 1A and 1B, 2A and 2B, and 3A and 3B.

The results of the exergoeconomic analysis and optimization study are also included.

Finally, comparative analysis between the introduced integrated system is carried out. The

analyzed and assessed systems are:

• Base Liquefaction System

• Advanced Liquefaction System without Catalyst (S1A)

• Advanced Liquefaction System with Catalyst (S1B)

• Advanced Liquefaction System without Catalyst with ORGs (S2A)

• Advanced Liquefaction System with Catalyst with ORGs (S2B)

• Advanced Liquefaction System without Catalyst with VTs (S3A)

• Advanced Liquefaction System with Catalyst with VTs (S3B)

5.1 Base System Results

The reference system presented was simulated based on a patent assigned to a

commercial entity in [19]. A comprehensive energy and exergy analyses was carried

out to examine the system performance.

Figure 5.1 illustrates the energy and exergy efficiencies of individual system

components. Equipment variables changed to test system outcomes as they are changed.

The Cooler (EX1), with 23% efficiency is the least efficient equipment.

Expansion Valve (V2) has the highest exergy efficiency and Expansion Valve

(V1) and Compressor (C2) are the second and third highest efficient equipment among

other units. Noticeably, Compressors and Valves are working with efficiency higher

than 80% and heat exchangers and expanders are working with lower exergy efficiency

than other equipment. The lower exergy efficiency of individual equipment affects the

overall system exergy efficiency. Different catalyst types and performance

enhancement of the heat exchangers could create a dramatic improvement to the

process efficiency.

61

Figure 5.1 Exergy and energy efficiencies for each component

Figure 5.2 demonstrates the effect of feeding the system with precooled

hydrogen starting from negative 50 °C on Compressors work losses. Compressor (C1)

is affected to most with loss of 20% as temperature increases while the rest of the

compressors are not affected by the feed temperature change

Figure 5.3 illustrates the effect of the pre-cooled hydrogen on the overall energy

and exergy efficiencies. The energy efficiency increases as feed temperature rises but

the overall exergy efficiency does not change on a degree of significance. This is due

to the change on only mainly one compressor C1. Kanoglu et al. in [90] researched

precooling using geothermal energy and showed improvement by decreasing the work

consumption by

Figure 5.4 displays the effect of hydrogen mass flow rate change on the work

of the turbo expanders, the adsorption unit and the overall yield. It can be noticed that

the work of turbo expanders TE1, TE2 and the overall yield are increased in a linear

manner while the work of the adsorption unit declined. In case of both turbo expanders,

the work increases from zero to roughly 3500 kw as the hydrogen mass flow rate

increases from 0 to 7000kg/h

24

49

34

9299

69

82 81

53

65

46

6253

47

26

80 80 80

100100 99

88 90

68 6876

66

99 99 99

0

20

40

60

80

100

Cooler (EX1)

Cooler (EX2)

Cooler (EX3)

Expansio

n Valve (V

1)

Expansio

n Valve (V

2)

Compressor (

C1)

Compressor (

C2)

Compressor (

C3)

Turbo Expander (T

E1)

Turbo Expander (T

E2)

Flash

Drum (F1)

Adsorp

tion Unit

(AO)

Heat Exch

anger (HX1)

Heat Exch

anger (HX2)

Heat Exch

anger (HX3)

Effic

ienc

y (%

)

System Components

Exergy Efficiency

62

(a)

(b)

Figure 5.2 Effect of pre-cooling Hydrogen on Compressors work losses and yield. The graphs illustrate a) work loss of compressors C1, C2, and C3 and, b) daily

hydrogen yield against hydrogen Feed temperature

0

100

200

300

400

500

600

-60 -40 -20 0 20 40 60

Wor

k Lo

ss (k

W)

H2 Feed temperature (oC)

Compressor C1

Compressor C2

Compressor C3

3.2

3.3

3.4

3.5

3.6

3.7

3.8

3.9

-50 -40 -30 -20 -10 0 10 20 30 40 50

Yiel

d (T

ons/

Day)

H2 Feed Temperature (oC)

63

Figure 5.3. Effect of pre-cooling hydrogen on overall energy and exergy efficiencies

(a)

10

11

12

13

14

15

16

17

18

19

20

-60 -40 -20 0 20 40 60

Effic

ienc

y (%

)

Hydrogen Feed Tempreture (oC)

Exergy efficiency

Energy Effeciency

0

500

1000

1500

2000

2500

3000

3500

4000

0 1000 2000 3000 4000 5000 6000 7000

Wor

k (k

W)

H2 Mass flow rate (kg/h)

Turbo Expander TE1

Turbo Expander TE2

64

(b)

(c)

Figure 5.4. Effect of hydrogen mass flow rate change on work and yield. The graphs show a) work (kW) of (a)Turbo expander (TE1), Turbo expander (TE2), (b)Adsorber (A0) and; (c) Liquid hydrogen generation per day against hydrogen Feed Mass Flow

rate (kg/h)

The effect of hydrogen mass flow rate change on the overall energy and exergy

efficiencies is shown in Figure 5.5. The exergy efficiency plummets after 10% increase

and energy efficiency declines gradually by a small percentage. An optimum mass flow

rate can therefore give a higher yield without compromising on the exergy efficiency.

-20

-18

-16

-14

-12

-10

-8

-6

-4

-2

0

0 1000 2000 3000 4000 5000 6000 7000

Work of Adsorber A0

0

20

40

60

80

100

120

140

160

180

200

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000 5500 6000 6500 7000 7500

Liqu

id H

2yi

eld

(ton

es/d

ay)

H2 Feed Mass Flow Rate (kg/h)

65

Figure 5.5 Effect of hydrogen mass flow rate change on the overall energy and exergy efficiencies

Figure 5.6 illustrates how changing turbo expander TE1 pressure is affecting

main components work. It can clearly be seen from figure 5.6 that the performance of

key components are not affected by variations in pressure of TE1. The graphs clearly

show that the pressure of turbo expander TE1 does not affect the work rate of the other

main components in the system.

Figure 5.6 Effect of changing turbo expander TE1 pressure on main components work

15.46

0.32

0

2

4

6

8

10

12

14

16

18

0 2000 4000 6000 8000

Effic

ienc

y (%

)

Hydrogen Feed Mass Flow Rate, kg/h

Exergy Effeciency

Energy Effeciency

-10

190

390

590

790

990

0 50 100 150 200

Wor

k Ra

te (k

W)

Turbo Expander TE1 Pressure, bar

Turbo Expander TE1 Compressor C1 Compressor C2

Compressor C3 Adsorption Unit A0 Turbo Expander TE2

66

Figure 5.7 illustrates the effect of changing turbo expander TE1 pressure on the

overall energy and exergy efficiencies. It is seen that TE1 has a notable impact on the

energy efficiency and a slight impact on the exergy efficiency. Both values decrease

with increase in TE1 pressure. In this system design, the turbo expander TE1 works

independently where is failure does not affect the overall system.

Figure 5.7 Effect of changing turbo expander TE1 pressure on the overall energy and exergy efficiencies

Figure 5.8 demonstrates the effect of changing pressure of flash drums on the

overall yield. Though it shows volatility in the production per day, it’s not significant

for a system of its scale and has minimal effect on the outcome. However, the yield

values show no discernible trend or consistency.

The effect of pressure of flash drums on energy and exergy are also negligible

as shown in Figure 5.9. In fact, the effective change in overall energy and exergy is

close to zero. This shows that the system startup period has a noticeable effect on the

10

11

12

13

14

15

16

0 50 100 150 200

Effic

ienc

y (%

)

TE1 Pressure (bar)

Exergy Effeciency Energy Effeciency

67

flash drum. It requires about 400 iterations in the simulation model to reach the steady

state for having the system to give required results.

Figure 5.8 Effect of changing flash drums pressure on the overall yield

Figure 5.9 Effect of flash drum pressure rate change on overall energy and exergy

10

11

12

13

14

15

16

0 1 2 3 4 5 6 7

Effic

ienc

y (%

)

Flash Drum Pressure (bar)

Exergy Effeciency

Energy Effeciency

68

Figure 5.10 Effect of hydrogen feed pressure change on the compressors

Changing the feed pressure does not help the overall system form compressors

perspective. In addition, both the exergy and energy efficiencies are unaffected by

variations in feed pressure. Therefore, the only effect of hydrogen feed pressure

increase is an increase in the work of compressor C1 as shown in Figure 5.10.

Figure 5.11 Effect of hydrogen feed pressure change on the overall energy and exergy efficiencies

236.28

236.3

236.32

236.34

236.36

236.38

236.4

0

100

200

300

400

500

600

700

800

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15

Wor

k (k

W)

Wor

k Ra

te (k

W)

H2 Feed Pressure (bar)

Work of Compressor C2 Work of Compressor C3 Work of compressor C1

10

11

12

13

14

15

16

17

18

19

20

0 5 10 15 20 25

Effic

ienc

y (%

)

Hydrogen Feed Pressure (bar)

Exergy Effeciency

Energy Effeciency

69

Figure 5.11 depicts the effect of the hydrogen feed pressure change on energy

and exergy are also negligible. In Figure 5.12, the Nitrogen gas mass flow rate change

is changing on the overall energy and exergy efficiencies. The change in energy

efficiency is noticeable but there is no significant change in exergy efficiency. The feed

can be utilized to cool the hydrogen feed. Other gases may also be utilized for the

cooling process that can replace nitrogen.

Figure 5.12 Effect of the Nitrogen gas mass flow rate change on the overall energy and exergy efficiencies

Temperature approach of the heat exchangers has a considerable effect on the

liquefaction plants. The more the pinch temperatures, the more the operating costs or

energy costs of the process. Figure 5.13, 5.14, and 5.15 show that hot composite curves

and heat flow through the main heat exchangers HX1 to HX2, and HX3 respectively.

10

11

12

13

14

15

16

0 10000 20000 30000 40000

Effic

ienc

y (%

)

Nitrogen Gas Mass Flow Rate (kg/h)

Exergy efficiency

Energy efficiency

70

Figure 5.13 Heat Exchanger HX1 Heat composite curves

-100

-50

0

50

100

150

-1.2E+08 -80000000 -40000000 0

Tem

pera

ture

(o C)

Heat Duty (kJ/h)(a)

-200

-150

-100

-50

0

-5E+08 -3E+08 -1E+08

Tem

pera

ture

(o C)

Heat Duty (kJ/h)(b)

-200

-150

-100

-50

0

-12000000 -8000000 -4000000 0

Tem

pera

ture

(o C)

Heat Duty (kJ/h)(c)

-250

-200

-150

-100

-50

0

1000000 6000000 11000000 16000000

Tem

pera

ture

(o C)

Heat Duty (kJ/h)(d)

-235

-185

-135

0 2000000 4000000 6000000 8000000 10000000 12000000 14000000

Tem

pera

ture

(o C)

Heat Duty (kJ/h)

(e)

71


Heat exchangers HX1, HX2, and HX3 Heat duties are the very sensitive to the temperatures and

each inlet is duty increase with the temperature increase. It can be said that in extreme cooling process there

is need to optimize the process with consideration of the inlet stream intake in the efficiency.

-230

-220

-210

-200

-190

-180

-1.35E+08 -8.50E+07 -3.50E+07 1.50E+07

Tem

pera

ture

(o C)

Heat Duty (kJ/h)

(a)

-230

-220

-210

-200

-190

-180

-170

-3.60E+07 -2.60E+07 -1.60E+07 -6.00E+06

Tem

pera

ture

(o C)

Heat Duty (kJ/h)

(b)

200

250

300

350

400

450

0.00E+00 1.00E+08 2.00E+08

Tem

pera

ture

(o C)

Heat Duty (kJ/h)

(c)

-235

-225

-215

-205

-195

-2.00E+07 -1.50E+07 -1.00E+07 -5.00E+06 0.00E+00

Tem

pera

ture

(o C)

Heat Duty (kJ/h)

(d)

72


-254

-249

-244

-239

-234

-1.00E+08 -6.00E+07 -2.00E+07

Tem

pera

ture

(o C)

Heat Duty (kJ/h)(a)

-253

-252.9

-252.8

-252.7

-252.6

-1.80E+06 -1.30E+06 -8.00E+05 -3.00E+05 2.00E+05

Tem

pera

ture

(o C)

Heat Duty (kJ/h)

(b)

-254

-249

-244

-239

-234

-2.00E+07 -1.50E+07 -1.00E+07 -5.00E+06 0.00E+00

Tem

pera

ture

(o C)

Heat Duty (kJ/h)

(c)

-248

-244

-240

-236

-232

-228

0.00E+00 2.00E+06 4.00E+06 6.00E+06 8.00E+06

Tem

pera

ture

(o C)

Heat Duty (kJ/h)

(e)

-255

-235

-215

-195

-175

-1.20E+07 -7.00E+06 -2.00E+06 3.00E+06

Tem

pera

ture

(o C)

Heat Duty (kJ/h)

(d)

-250

-150

-50

50

150

250

350

-3.00E+07 2.00E+07 7.00E+07 1.20E+08

Tem

pera

ture

(o C)

Heat Duty (kJ/h)

(f)

73

5.2 Systems 1A and 1B Results

With results in the base system, it was clear that some of the system components need

to be changed. The performance aspects of the base system are investigated by a

comprehensive study covering energy and exergy analyses. Various operating conditions,

reference state parameters, and system parameters are altered to examine each parameter’s

effects on the performance of system 1A. From this analysis, it is evident that some aspects

had to be modified and changed to produce a working functional system illustrates the

energy and exergy efficiencies of the equipment in the system individually. Equipment

variables changed to test system outcomes as they are changed.

The results show that the cooler (EX1) has the lowest efficiency (23%).

Expansion Valve (V2) has the highest exergy efficiency and Expansion Valve (V1) and

Compressor (C2) are the second and third highest efficient equipment among other

units. In fact, Compressors and Valves are working with efficiency higher than 80%,

while, heat exchangers and expanders are working with lower exergy efficiency than

other equipment. The lower exergy efficiency of almost all equipment affects the



process efficiency. Figure 5.16 shows the exergy and energy efficiency of systems 1A

and 1B.

Figure 5.17 shows the effect of hydrogen feed pressure on energy and exergy

efficiency of system 1A. It can be seen that the energy efficiency does vary with

variation in hydrogen feed pressure. In fact, the energy efficiency increased 12.4% as

the hydrogen feed pressure increased from 1 bar to 5 bars. The exergy efficiency also

increases from 17.42% to 18.04% for the same change.

74

(a)

(b) Figure 5.16 Energy efficiencies for main components of System 1A and 1B

Figure 5.18 shows that changes in compressor (C1) pressure does have an effect

on both energy and exergy efficiency of system 1A. The energy efficiency increases

from 41.83% to 67.59% as the compressor pressure increased from 2 bar to 4 bar. For

the same change in compressor pressure, the exergy efficiency increases moderately

from 17.59% to 16.78%.

24

49

34

92 99

69

82 81

53

65

23

62

31

63

41

99 99

8389

98

64

85

28

74

0102030405060708090

100

Heat Ex

chan

ger (H

X1)

Heat Ex

chan

ger (H

X2)

Heat Ex

chan

ger (H

X3)

Expan

sion Valv

e (V1)

Expan

sion Valv

e (V2)

Compressor (C

1)

Compressor (C

2)

Compressor (C

3)

Compressor (C

3)

Compressor (C

3)

Flash

Drum (F1)

Adsorptio

n Unit (AO)

Effic

ienc

y (%

)

Exergy Efficiency

Energy Efficiency

77

18

42

99

60

23

77 78 7867

23

62

23

55

78

30

99 99 9987

30

74

0102030405060708090

100

Heat Ex

chan

ger (H

X1)

Heat Ex

chan

ger (H

X2)

Heat Ex

chan

ger (H

X3)

Expan

sion Valv

e (V1)

Expan

sion Valv

e (V2)

Compressor (C

1)

Compressor (C

2)

Compressor (C

3)

Compressor (C

3)

Compressor (C

3)

Flash

Drum (F1)

Adsorptio

n Unit (AO)

Effic

ienc

y (%

)

Exergy Efficiency

Energy Efficiency

75

Figure 5.17. Effect of hydrogen feed pressure on overall efficiencies for System 1A

Figure 5.18 Effect of Compressor C1 pressure on overall efficiencies for System 1A

Figure 5.19 Effect of Compressor C2 pressure overall efficiencies for System 1A

10

15

20

25

30

35

0 5 10 15 20 25

Effic

ienc

y (%

)


Exergy Efficiency

Energy Efficiency

0

10

20

30

40

50

60

70

80

1.5 2 2.5 3 3.5 4 4.5

Effic

ienc

y (%

)

Compressor C1 Pressure (bar)

Energy Efficiency

Exergy Efficiency

7

9

11

13

15

17

19

1 2 3 4 5 6

Effic

ienc

y (%

)

Compressor C2 Pressure (bar)

Energy Efficiency

Exergy Efficiency

76

Figure 5.19 shows how the effect of variations in compressor C2 on the energy

and exergy efficiency of system 1A. The variation in exergy efficiency is negligible.

However, the energy efficiency initially drops and then rises. As the compressor (C2)

rises from 2 bar to 4 bar, the energy efficiency decreases from 15.85% to 11.14%.

However, as the compressor (C2) pressure rises from 4 bar to 5 bar, the energy

efficiency rises from 11.14% to 13%.

Figure 5.20 shows that changes in H2 feed pressure on overall energy and exergy

efficiency of system 1B. It can be seen that the variation in efficiency caused by feed

pressure variation is minimal in both cases. Both efficiencies do rise, but by less than

0.5% in both cases.

Figure 5.20 Effect of hydrogen H2 feed pressure on overall efficiencies for System 1B

Figure 5.21 Effects of Compressor C2 pressure on overall energy and exergy efficiencies for System 1B

17

18

19

20

21

22

23

0 5 10 15 20 25

Effic

ienc

y (%

)

Pressure (bar)

Exergy Efficiency

Energy Efficiency

1517192123252729313335

2 2.5 3 3.5 4

Effic

ienc

y (%

)

Pressure (bar)

Exergy Efficiency

Energy Efficiency

77

Figure 5.21 clearly shows that variations in compressor (C2) pressure has no

effect on either energy or exergy efficiency of system 1B.

5.2.1 Pre-cooling phase at systems S1A and S1B

The precooling phase in the liquefaction cycle helps cool the hydrogen gas for faster

liquefaction. An analysis has been conducted to understand the precooling phase of the

heat exchangers. Figure 5.22 shows the Heat Load, Exergy flow and Temperature

against specific exergy flow for Precooling Phase heat exchanger HX1 at the liquid

nitrogen inlet. The graphs indicate that as specific exergy decreases, total exergy

decreases while heat load increases beyond a certain value. While temperature initially

rises with the decrease in specific exergy, it soon stabilises.

Figure 5.23 shows the heat load, exergy flow and temperature for Precooling

Phase heat exchanger HX1 at stream 9 inlet and depicts a different view from the

Nitrogen inlet but in decreasing exergy flow. At inlet 9 and inlet 28, the exergy flow

increases slightly in the heat exchanger as a sign of need of optimization as shown in

Figure 5.23 and Figure 5.24.

Figure 5.22 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream N2LIQ inlet for System 1A

-195.805

-195.8

-195.795

-195.79

-195.785

-195.78

-195.775

0

200

400

600

800

1000

1200

760.2 760.1 709.6 658.9 608.3 557.7 507.1 456.5 405.9 355.3 304.7 254.1 203.5

Tem

pera

ture

(˚C

)

Hea

t Loa

d an

d Ex

ergy

Flo

w (k

W)

Specific Exergy Flow (kJ/kg)

Total ExergyHeat LoadTemperature

78

The precooling phase shows good indicator that it can effect the changes on the

overall system and possibly on the efficiency as can bee seen at the end of the chapter

with the analysis.

Figure 5.23 Heat Load, Exergy flow and Temperature for Precooling Phase heat

exchanger HX1 at stream 28 inlet for System 1A



At inlet 36 and 44 the exergy flow drops at an almost constant rate as shown in

Figure 5.25 and Figure 5.26. It can be seen that in both cases, as the specific exergy

-140

-120

-100

-80

-60

-40

-20

0

-20

-15

-10

-5

0

5

10

15

20

25

30

401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



-150

-100

-50

0

50

100

150

-400

-300

-200

-100

0

100

200

300

400

3903.8 3822.1 3756.5 3709.2 3682.8 3680.5 3706.6 3766.2 3866.3 4016.2 4228.7 4522.1

Tem

pera

ture

(o C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)

Specific Exergy (kj/kg)


79

flow rises, heat load increases and exergy decreases. Additionally, higher specific

energy flows correspond to higher inlet temperatures in both cases.

Figure 5.25 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 36 inlet for System 1A


Figure 5.27 shows the Heat Load, exergy flow, and temperature for heat

exchanger HX1 at stream (R) inlet for system 1A. It can be seen that both the heat load

and exergy plots diverge with increase in specific energy flow. The total exergy flow

increases while heat load increases with increase in specific exergy flow. Lower values

of specific exergy flow also correspond to lower temperatures. In addition, Figure 5.28

-200

-180

-160

-140

-120

-100

-80

-60

-40

-20

0

0

10

20

30

40

50

60

70

401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



-200

-180

-160

-140

-120

-100

-80

-60

-40

-20

0

0

5

10

15

20

25

30

35

40

45

401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



80

shows the heat load, exergy flow, and temperature for heat exchanger (HX1) at stream

N2LIQ inlet for system 1B and signals that as specific exergy goes down also the total

exergy goes down with slight change in temperature value.

Figure 5.27 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream R inlet for System 1A

Figure 5.28 Heat Load, Exergy flow vs Temperature for Precooling Phase heat

exchanger HX1 at stream N2LIQ inlet for System 1B

Figure 5.29 shows the changes in heat load, Exergy flow and Temperature for

Precooling Phase heat exchanger HX1 at stream 28 inlet. It is seen that with increase in

specific exergy flow rate, the total exergy flow increases while heat load decreases.

Additionally, lower temperatures correspond to higher values of specific exergy flow.

-200-180-160-140-120-100-80-60-40-200

-20

0

20

40

60

80

100

401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



-195.805

-195.8

-195.795

-195.79

-195.785

-195.78

-195.775

0

200

400

600

800

1000

1200

760.2760.1709.2658.3607.3556.4505.5454.5403.6352.6301.7250.8199.8

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



81

Figure 5.29 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 28 inlet for System 1B

Figure 5.30 illustrates the variation of heat load, exergy flow, and temperature

for heat exchanger HX1 in the precooling phase at stream inlet 9. The total exergy flow

has a nonlinear relationship with specific exergy flow in this case. The trend is that as

specific exergy flow increases, the total exergy flow also increases, but the two have a

non-linear relationship.


Figure 5.31 shows the variation of temperature, exergy flow and heat load for

precooling heat exchanger HX1 at stream 36 inlet. With increase in specific exergy

flow, it is seen that total exergy flow rate decreases and heat load increases. Similarly,

lower inlet temperatures correspond to lower values of specific exergy flow. In this

-140

-120

-100

-80

-60

-40

-20

0

-20-15-10

-505

1015202530

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



-150

-100

-50

0

50

100

150

-400

-300

-200

-100

0

100

200

300

400

3042.8 2961.9 2897.0 2850.3 2824.5 2822.6 2849.0 2908.7 3008.6 3158.0 3369.6 3661.8

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



82

case, it can be noted that the relationship between specific exergy flow and all three

other variables is linear.


exchanger HX1 at stream 36 inlet for System 1B




precooling heat exchanger HX1 at stream 44 inlet. With the increase in specific exergy

-180

-175

-170

-165

-160

-155

-150

-145

-140

-135

0

10

20

30

40

50

60

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



-180

-175

-170

-165

-160

-155

-150

-145

-140

-135

0

5

10

15

20

25

30

35

40

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



83


lower inlet temperatures correspond to lower values of specific exergy flow.

5.2.2 Liquefaction Phase at systems S1A and S1B

This section deals with the variation of Heat Load, exergy, and temperature for the

liquefaction phase of systems S1A and S1B. Figure 5.33 shows heat load, Exergy flow

vs Temperature for Precooling Phase heat exchanger HX2 at stream 49 inlet for System

1A. It can be seen that total exergy flow increases with increase in specific exergy flow

while heat load decreases. The variation of heat load is linear, while total exergy flow

has a nonlinear relationship. Higher specific exergy flows can be achieved at lower

temperatures at the actual liquidation phase.



Figure 5.34 illustrates the variation of heat load, exergy flow and temperature

for Precooling Phase heat exchanger HX2 at stream 13 inlet for System 1A. Much like

in Figure 5.33, it can be seen that total exergy flow increases with increase in specific

exergy flow while heat load decreases. The variation of heat load is linear, while total

exergy flow has a nonlinear relationship. Higher specific exergy flows are achieved at

lower temperatures. The variation of these parameters for Precooling Phase heat

exchanger HX2 at stream inlet S10 is shown in Figure 5.35. The trends are similar to

those at stream inlet 13 and stream inlet 9. However, the nonlinearity in total exergy

flow is more noticeable at stream inlet S10.

-200

-180

-160

-140

-120

-100

-80

-60

-40

-20

0

-100

-50

0

50

100

150

200

401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



84


Figure 5.35 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream S10 inlet for System 1A


for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 1A. It can be

seen that the heat exchanger HX1, the specific exergy flow is lower at lower inlet

temperatures. Additionally, the heat load increases linearly with increase in specific

exergy flow, while total exergy flow has a nonlinear, but decreasing relationship with

specific exergy flow.

-200

-180

-160

-140

-120

-100

-80

-60

-40

-20

0

-100

0

100

200

300

400

500

401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



-250

-200

-150

-100

-50

0

0

20

40

60

80

100

120

140

401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



85

The variation brought about by pre cooling in heat exchanges HX3 can be seen

in Figure 5.37, where the heat load, exergy flow and temperature for Liquefaction Phase

heat exchanger HX3 at stream 22 inlet for System 1A is illustrated. The key difference

is that total exergy flow now increases with increase in specific exergy flow while

overall heat load decreases. Specific exergy flow is higher at lower temperatures,

though the variation of temperature vs. specific exergy is minor.

Figure 5.38 and 5.40 illustrate variations of temperature, total exergy flow,

and heat load for liquefaction phase heat exchanger HX3 at stream inlets 32b and S1

respectively. At stream inlet 32b, the variation of total exergy flow offers a unique

trend. It is seen that as specific exergy flow increases, the total exergy flow increases

initially before dropping sharply at a point. The specific exergy values are higher at

lower temperatures. The heat load decreases slightly with increase in specific exergy.

With variations at S1, it is seen in Figure 5.39 that the heat load and total exergy flow

remains constant with variations in specific exergy. However, the temperature and

specific exergy flow have a slightly nonlinear relationship with specific exergy being

lower at lower temperatures.

Figure 5.36 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 1A

Figure 5.40 shows the variation of heat load, exergy and temperature for HX3

at stream inlet 17 for system 1A. It can be seen that heat load increases with increase in

specific exergy flow while total exergy flow decreases. At lower temperatures, the

-180

-160

-140

-120

-100

-80

-60

-40

-20

0

0

20

40

60

80

100

120

140

401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



86

specific exergy flow is higher and vice versa. The next set of graphs illustrate the

variation of exergy, heat load and temperature for system 1B.

Figure 5.37 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat



exchanger HX3 at stream 32b inlet for System 1A

Figures 5.41, 5.42, and 5.43 illustrate the variations in temperature, heat load,

and exergy for system 1B for stream inlets S10, 14, and 13. Figure 5.41 and Figure 5.42

show similar trends. The total exergy decreases with increase in specific exergy flow

while the heat load increases. Specific exergy flow is lower at lower temperatures.

-202

-200

-198

-196

-194

-192

-190

-188

-186

-184

-182

-20

0

20

40

60

80

100

120

140

401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



-235

-230

-225

-220

-215

-210

-205

-10

0

10

20

30

40

50

60

70

401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



87



-245.04

-245.03

-245.02

-245.01

-245

-244.99

-244.98

-1400

-1200

-1000

-800

-600

-400

-200

0

200

401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



-70

-60

-50

-40

-30

-20

-10

0

0

2

4

6

8

10

12

401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1

Tem

pera

ture

(˚C

)

Hea

t Loa

d an

d Ex

ergy

Flo

w (k

W)



88

Figure 5.41 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream S10 inlet for System 1B

Figure 5.42 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 1B

However, the variation of heat load is more noticeable in inlet 14, and the

variation of total exergy flow exhibits a nonlinear pattern in this case (Figure 5.42). In

figure 5.43 (corresponding to inlet 13) the trends are different. There is a slight increase

in total exergy flow and a slight increase in heat load as the specific exergy flow

increases. The value of specific exergy flow is higher at lower temperatures.

-203

-202.5

-202

-201.5

-201

-200.5

-200

-199.5

-199

-198.5

0

20

40

60

80

100

120

140

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



-180

-160

-140

-120

-100

-80

-60

-40

-20

0

0

20

40

60

80

100

120

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



89


In the next set of figures, the variations in exergy, temperature, and heat load

for heat exchanger HX3 in system 1B is illustrated. Figure 5.44 and Figure 5.45 show

the variations related to stream inlets 49 and 22 respectively. Both exhibit similar

trends, with lower temperatures corresponding to higher values of specific exergy flow.

The heat load in both cases decrease with increase in specific exergy flow while total

exergy flow increases.

Figure 5.46 corresponds to variations in heat load, exergy and temperature for

HX3 at stream inlet 17. The value of specific exergy flow is lower at lower values of

temperature. The heat load increases with increase in specific exergy flow. However,

the variation in total exergy flow shows an interesting trend. Initially, with increase in

specific exergy flow the total exergy flow increases. However, at a certain point, the

trend changes, and the total; exergy flow decreases in a nonlinear manner with increase

in specific exergy flow.

-131

-130

-129

-128

-127

-126

-125

-124

-123

-122

-50

0

50

100

150

200

250

300

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



90



Figure 5.47 ilustrates the variations corresponding to stream inlet S1. It can be

seen that the total exergy flow and heat load are unvaried in this case with increase in

specific exergy flow. The temperature and specific exergy flow have a slightly

nonlinear relationship, with lower values of specific exergy flow corresponding to

lower temperatures.

-200

-180

-160

-140

-120

-100

-80

-60

-40

-20

0

-40

-20

0

20

40

60

80

100

120

140

160

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



-250

-200

-150

-100

-50

0

-100

-50

0

50

100

150

200

250

300

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2Te

mpe

ratu

re (˚

C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



91


Figure 5.47 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream S1 inlet for System 1B

-250

-200

-150

-100

-50

0

-100

-50

0

50

100

150

200

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



-245.012

-245.01

-245.008

-245.006

-245.004

-245.002

-245

-244.998

-244.996

-244.994

-1600

-1400

-1200

-1000

-800

-600

-400

-200

0

200

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



92

Figure 5.48 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 32b inlet for System 1B

The final figure in this series, Figure 5.48 shows the variation in heat load, Exergy

flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 32b inlet

for System 1B. Specific exergy flow is lower at higher temperatures in this case. It is

seen that the heat load decreases with increase in specific exergy flow. The variation of

total exergy flow initially increases with increase in specific exergy flow up to a point,

but then decreases linearly.

5.3 Systems 2A and 2B Organic Rankine Cycles

Figure 5.49 illustrates the energy and exergy efficiencies of the equipment in the system

individually. Equipment variables changed to test system outcomes as they are

changed. The least efficiency is at 23 % for the Cooler (EX1). Expansion Valve (V2)

has the highest exergy efficiency and Expansion Valve (V1) and Compressor (C2) are

the second and third highest efficient equipment among other units. Noticeably,

Compressors and Valves are working with efficiency higher than 80%. While, heat

exchangers and expanders, are working with lower exergy efficiency than other

equipment. The lower exergy efficiency of almost all of some equipment affects the



process efficiency.

-235

-230

-225

-220

-215

-210

-205

-10

0

10

20

30

40

50

60

70

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



93

Figure 5.49 Systems (a) 2A and (b) 2B Exergy and Energy efficiencies for each component

Figure 5.50 illustrates the effect of the hydrogen gas feed pressure on the overall

energy and exergy efficiencies. The energy efficiency increases as feed temperature

rises but the overall exergy efficiency does not change on a degree of significance. This

is due to the change on only mainly one compressor C. Kanoglu et al. in [90] researched

precooling using geothermal energy and showed improvement by decreasing the work

consumption by 25%.

44

8

4732

11

33 35

9

77

9

77 76 77 75

37

57

10

61

38

13

40 46

12 12

92 99 92 90

44

0102030405060708090

100

Heat Ex

chan

ger (H

X1)

Heat Ex

chan

ger (H

X2)

Heat Ex

chan

ger (H

X3)

Cooler EX1

Cooler EX2

Cooler EX3

Cooler EX4

Cooler EX5

Cooler EX6

Compressor (C

1)

Compressor (C

2)

Compressor (C

3)

Compressor (C

4)

Compressor (C

5)

Flash

Drum (F1)

Effic

ienc

y (%

)

Exergy Efficiency

Energy Efficiency

(a)

25

7 5

32 30

86

1832

9

77 76 77 78 78

4030

8 6

42 39

23

42

11

92 91 92 94

48

0102030405060708090

100

Heat Ex

chan

ger (H

X1)

Heat Ex

chan

ger (H

X2)

Heat Ex

chan

ger (H

X3)

Cooler EX1

Cooler EX2

Cooler EX3

Cooler EX4

Cooler EX5

Cooler EX6

Compressor (C

1)

Compressor (C

2)

Compressor (C

3)

Compressor (C

4)

Compressor (C

5)

Flash

Drum (F1)

Effic

ienc

y (%

)

Exergy Efficiency

Energy Efficiency

(b)

94

(a)

(b)

Figure 5.50. Effect of pre-cooling hydrogen feed pressure variations on overall energy and exergy efficiencies for systems (a) S2A and (b) S2B

The effect of Compressor (C1) pressure change on the overall energy and

exergy efficiencies is shown in Figure 5.51 and it shows the exergy efficiency plummets

at after the 10% increase and the energy efficiency declines by a small percentage. An

optimum mass flow rate can make give and a higher yield without compromising on

the exergy efficiency.

0.0%

10.0%

20.0%

30.0%

40.0%

50.0%

60.0%

1.0 3.9 6.7 9.6 12.4 15.3 18.1

Perc

enta

ge


Exergy Efficiency

Energy Efficiency

13.00%

13.50%

14.00%

14.50%

15.00%

15.50%

1.0 3.9 6.7 9.6 12.4 15.3 18.1

Perc

enta

ge


Exergy Efficiency

Energy Efficiency

95

(a)

(b)

Figure 5.51 Effect of Compressor 1 (C1) pressure variations on overall energy and exergy efficiencies for systems (a) S2A and (b) S2B

Figure 5.52 illustrates the effect of changing Compressor 5 (C5) pressure on the

overall energy and exergy efficiencies that C5 has an impact on the energy efficiency

and very slight impact on the exergy efficiency.

0.00%

10.00%

20.00%

30.00%

40.00%

50.00%

60.00%

70.00%

80.00%

1.50 2.00 2.50 3.00 3.50 4.00 4.50

Effic

ienc

y

Compressor 1 (C1) pressure (bar)

Exergy Efficiency

Energy Efficiency

0.00%

10.00%

20.00%

30.00%

40.00%

50.00%

60.00%

70.00%

1.50 2.00 2.50 3.00 3.50 4.00 4.50

Effic

ienc

y


Exergy Efficiency

Energy Efficiency

96

(a)

(b) Figure 5.52 Effect of Compressor 5 (C5) pressure variations on overall energy and

exergy efficiencies for systems (a) S2A and (b) S2B 5.3.1 Pre-cooling phase at systems S2A and S2B



heat exchangers. Figure 5.53 shows the Heat Load, Exergy flow and Temperature

against specific exergy flow for Precooling Phase heat exchanger HX1 at the liquid

nitrogen inlet. The graphs indicate that as specific exergy decreases, total exergy

decreases while heat load increases beyond a certain value. While temperature initially

rises with the decrease in specific exergy, it soon stabilises.

0.0%

10.0%

20.0%

30.0%

40.0%

50.0%

60.0%

70.0%

2.00 2.22 2.44 2.67 2.89 3.11 3.33 3.56 3.78 4.00 20.00

Effic

ienc

y


Exergy Efficiency

Energy Efficiency

0.00%

5.00%

10.00%

15.00%

20.00%

25.00%

2.00 2.22 2.44 2.67 2.89 3.11 3.33 3.56 3.78 4.00 20.00

Effic

ienc

y


Exergy Efficiency

Energy Efficiency

97




increases as can be seen in Figure 5.54 and Figure 5.55.


exchanger HX1 at stream N2LIQ inlet for System 2A






energy flows correspond to higher inlet temperatures in both cases.

-195.805

-195.8

-195.795

-195.79

-195.785

-195.78

-195.775

-10000

0

10000

20000

30000

40000

50000

60000

70000

80000

181.6 181.6 181.6 181.6 181.6 181.6 181.6 181.5 181.5 181.5 181.5 181.5 45.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)

Spesific Exergy Flow (kJ/kg)

TOTAL EXERGYFL

Heat duty

Temperature

-160

-140

-120

-100

-80

-60

-40

-20

0

-60

-40

-20

0

20

40

60

80

100

120

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)


TOTAL EXERGYFLHeat duty

98



-200

-180

-160

-140

-120

-100

-80

-60

-40

-20

0

-600

-400

-200

0

200

400

600

800

1000

891.4 901.9 915.9 934.0 956.8 984.9 1019.3 1061.1 1111.6 1172.8 1247.4 1339.3

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)


TOTAL EXERGYFLHeat dutyTemperature

-250

-200

-150

-100

-50

0

0

20

40

60

80

100

120

140

160

180

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



99






Precooling Phase heat exchanger HX1 at stream 28 inlet. It is seen that with increase in

specific exergy flow rate, the total exergy flow increases while heat load decreases.

Additionally, lower temperatures correspond to higher values of specific exergy flow.

-250

-200

-150

-100

-50

0

0

1000

2000

3000

4000

5000

6000

7000

8000

9000

10000

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-1.96E+02

-1.96E+02

-1.96E+02

-1.96E+02

-1.96E+02

-1.96E+02

-1.96E+02

-10000

0

10000

20000

30000

40000

50000

60000

70000

80000

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



100










-1.60E+02

-1.40E+02

-1.20E+02

-1.00E+02

-8.00E+01

-6.00E+01

-4.00E+01

-2.00E+01

0.00E+00

-30

-20

-10

0

10

20

30

40

50

60

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-200

-180

-160

-140

-120

-100

-80

-60

-40

-20

0

-300

-200

-100

0

100

200

300

400

500

600

700

3732.7 3775.8 3833.7 3908.4 4002.1 4117.8 4259.0 4430.3 4637.5 4888.7 5195.0 5572.4

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)


TOTAL EXERGYFLHeat dutyVapor fraction

101














This section deals with the variation of Heat Load, exergy, and temperature for the

liquefaction phase of systems S2A and S2B. Figure 5.63 shows heat load, Exergy flow

vs Temperature for liquefaction Phase heat exchanger HX2 at stream 49 inlet for

System 2A. It can be seen that total exergy flow increases with increase in specific



lower temperatures.

-2.50E+02

-2.00E+02

-1.50E+02

-1.00E+02

-5.00E+01

0.00E+00

0

50

100

150

200

250

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



102



Figure 5.63 Heat Load, Exergy flow and Temperature for liquefaction Phase heat



for Liquefaction Phase heat exchanger HX2 at stream 13 inlet for System 2A. Much

like in Figure 5.63 it can be seen that total exergy flow increases with increase in

specific exergy flow while heat load decreases. The variation of heat load is linear,

while total exergy flow has a nonlinear relationship. Higher specific exergy flows are

-2.50E+02

-2.00E+02

-1.50E+02

-1.00E+02

-5.00E+01

0.00E+00

0

20

40

60

80

100

120

140

160

180

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-250

-200

-150

-100

-50

0

50

-600

-400

-200

0

200

400

600

800

1000

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



103

achieved at lower temperatures. The variation of these parameters for Liquefaction

Phase heat exchanger HX2 at stream inlet 26 is shown in Figure 5.65. The trends are

similar to those at stream inlet 13 and stream inlet 9. However, the nonlinearity in total

exergy flow is more noticeable at stream inlet 26.



-250

-200

-150

-100

-50

0

-200

0

200

400

600

800

1000

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-300

-250

-200

-150

-100

-50

0

-800

-600

-400

-200

0

200

400

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



104



seen that as with HX1, the specific exergy flow is lower at lower inlet temperatures.

Additionally, the heat load increases linearly with increase in specific exergy flow,

while total exergy flow has a nonlinear, but decreasing relationship with specific exergy

flow.


in Figure 5.67, where the heat load, exergy flow and temperature for Liquefaction Phase






Figure 5.68 Figure 5.69 illustrate variations of temperature, total exergy flow,

and heat load for liquefaction phase heat exchanger HX3 at stream inlets 32b and S1

respectively. At stream inlet 32b, the variation of total exergy flow offers a unique




-300

-200

-100

0

100

200

300

400

500

600

700

800

0

100

200

300

400

500

600

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



105







Figure 5.68 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 32b inlet for System 2A

-2.42E+02

-2.40E+02

-2.38E+02

-2.36E+02

-2.34E+02

-2.32E+02

-2.30E+02

-2.28E+02

-2.26E+02

-2.24E+02

-800

-700

-600

-500

-400

-300

-200

-100

0

100

200

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-235

-230

-225

-220

-215

-210

-205

-50

0

50

100

150

200

250

300

350

400

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



106


exchanger HX3 at stream S1 inlet for System 2A





variation of exergy, heat load and temperature for system 2A.



-3.00E+02

-2.50E+02

-2.00E+02

-1.50E+02

-1.00E+02

-5.00E+01

0.00E+00

-20000

0

20000

40000

60000

80000

100000

120000

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



0.00E+00

1.00E+02

2.00E+02

3.00E+02

4.00E+02

5.00E+02

6.00E+02

7.00E+02

8.00E+02

-500

-400

-300

-200

-100

0

100

200

300

400

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



107

Figure 5.71 shows heat load, Exergy flow vs Temperature for Liquefaction

Phase heat exchanger HX2 at stream 49 inlet for System 2B. It can be seen that total

exergy flow increases with increase in specific exergy flow while heat load decreases.

The variation of heat load is linear, while total exergy flow has a nonlinear relationship.

Higher specific exergy flows are achieved at lower temperatures.


Figure 5.72 and Figure 5.73 and Figure 5.74 illustrate the variations in

temperature, heat load, and exergy for system 2B for stream inlets 26, 14, and 13. Figure

5.73 Figure 5.74 show similar trends. The total exergy decreases with increase in

specific exergy flow while the heat load increases. Specific exergy flow is lower at

lower temperatures. However, the variation of heat load is more noticeable in inlet 14,

and the variation of total exergy flow exhibits a nonlinear pattern in this case. In figure

Figure 5.74 (corresponding to inlet 13) the trends are different. There is a slight increase

in total exergy flow and a slight increase in heat load as the specific exergy flow

increases. The value of specific exergy flow is higher at lower temperatures.

-2.50E+02

-2.00E+02

-1.50E+02

-1.00E+02

-5.00E+01

0.00E+00

-300

-200

-100

0

100

200

300

400

500

600

700

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



108





The next set of figures illustrate the variations in exergy, temperature, and heat

load for heat exchanger HX3 in system 2B. Figure 5.75 and Figure 5.76 show the

variations related to stream inlets 22 and 17 respectively. Both exhibit similar trends,

with lower temperatures corresponding to higher values of specific exergy flow. The

heat load in both cases decrease with increase in specific exergy flow while total exergy

flow increases.

-203

-202.5

-202

-201.5

-201

-200.5

-200

-199.5

-199

-198.5

0

20

40

60

80

100

120

140

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



-180

-160

-140

-120

-100

-80

-60

-40

-20

0

0

20

40

60

80

100

120

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



109





Figure 5.77 corresponds to variations in heat load, exergy and temperature for

HX3 at stream inlet 17. The value of specific exergy flow is lower at lower values of

temperature. The heat load increases with increase in specific exergy flow. However,

the variation in total exergy flow shows an interesting trend. Initially, with increase in

specific exergy flow the total exergy flow increases. However, at a certain point, the

-131

-130

-129

-128

-127

-126

-125

-124

-123

-122

-50

0

50

100

150

200

250

300

396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2

Tem

pera

ture

(˚C)

Heat

Loa

d an

d Ex

ergy

Flo

w (k

W)



-2.42E+02

-2.40E+02

-2.38E+02

-2.36E+02

-2.34E+02

-2.32E+02

-2.30E+02

-2.28E+02

-2.26E+02

-2.24E+02

-350

-300

-250

-200

-150

-100

-50

0

50

100

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



110

trend changes, and the total; exergy flow decreases in a nonlinear manner with increase

in specific exergy flow.




exchanger HX3 at stream S1 inlet for System 2B




-2.50E+02

-2.00E+02

-1.50E+02

-1.00E+02

-5.00E+01

0.00E+00

0

50

100

150

200

250

300

350

400

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-3.00E+02

-2.50E+02

-2.00E+02

-1.50E+02

-1.00E+02

-5.00E+01

0.00E+00

-1500

-1000

-500

0

500

1000

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



111


lower temperatures.

The final figure in this series, Figure 5.78 shows the variation in heat load, Exergy

flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 32b inlet

for System 1B. Specific exergy flow is lower at higher temperatures in this case. It is

seen that the heat load decreases with increase in specific exergy flow. The variation of

total exergy flow initially increases with increase in specific exergy flow up to a point,

but then decreases linearly.


exchanger HX3 at stream 32b inlet for System 2B

5.4 Systems 3A and 3B – Vortex tubes

Figure 5.79 illustrates the energy and exergy efficiencies of the equipment in the system

individually. Equipment variables changed to test system outcomes as they are

changed. The least efficiency is at 23 % for the Cooler (EX1). Expansion Valve (V2)

has the highest exergy efficiency and Expansion Valve (V1) and Compressor (C2) are

the second and third highest efficient equipment among other units.

Noticeably, Compressors and Valves are working with efficiency higher than

80% while, heat exchangers and expanders, are working with lower exergy efficiency

than other equipment. The lower exergy efficiency of almost all of some equipment

affects the overall system exergy efficiency. Different catalyst types and performance

-2.35E+02

-2.30E+02

-2.25E+02

-2.20E+02

-2.15E+02

-2.10E+02

-2.05E+02

-50

0

50

100

150

200

250

300

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



112


process efficiency.

Figure 5.79 Systems 3A and 3B Exergy and Energy efficiencies for each component

Noticeably, Compressors and Valves are working with efficiency higher than

80% while, heat exchangers and expanders, are working with lower exergy efficiency

than other equipment. The lower exergy efficiency of almost all of some equipment

affects the overall system exergy efficiency. Different catalyst types and performance


process efficiency.

5.4.1 Pre-cooling phase at systems S3A and S3B



heat exchangers. shows the Heat Load, Exergy flow and Temperature against specific

7

84 81

23

2.2 4

35

9

77 7… 78 78

78

97 95

23

3 5

36

9

85 8095 93

36

0102030405060708090

100

Heat Ex

chan

ger (H

X1)

Heat Ex

chan

ger (H

X2)

Heat Ex

chan

ger (H

X3)

Expan

sion Valv

e (V1)

Expan

sion Valv

e (V2)

Expan

sion Valv

e (V3)

Vortex T

ube (VT1

)

Vortex T

ube (VT2

)

Compressor (C

1)

Compressor (C

2)

Compressor (C

3)

Flash

Drum (F1)

Adsorptio

n Unit (AO)

Effic

ienc

y (%

)

Exergy Efficiency

Energy Efficiency

93

16

81

2331

3

35

9

7786 90 96

15

96

19

94

2434

3

39

9

85 100 99 99

44

0102030405060708090

100

Heat Ex

chan

ger (H

X1)

Heat Ex

chan

ger (H

X2)

Heat Ex

chan

ger (H

X3)

Expan

sion Valv

e (V1)

Expan

sion Valv

e (V2)

Expan

sion Valv

e (V3)

Vortex T

ube (VT1

)

Vortex T

ube (VT2

)

Compressor (C

1)

Compressor (C

2)

Compressor (C

3)

Flash

Drum (F1)

Adsorptio

n Unit (AO)

Effic

ienc

y (%

)

Exergy Efficiency

Energy Efficiency

113

exergy flow for Precooling Phase heat exchanger HX1 at the liquid nitrogen inlet. The

graphs indicate that as specific exergy decreases, total exergy decreases while heat load

increases beyond a certain value. While temperature initially rises with the decrease in

specific exergy, it soon stabilises.




increases as can be seen in Figure 5.81 and Figure 5.82

Figure 5.80 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream N2LIQ inlet for System 3A




energy flows correspond to higher inlet temperatures in both cases

-250

-200

-150

-100

-50

0

0

200

400

600

800

1000

1200

760.2 760.1 687.9 615.6 543.4 471.1 398.8 326.6 254.3 190.4 182.2 125.0 84.8 56.1

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



114



.

-130

-125

-120

-115

-110

-105

-100

-20

-10

0

10

20

30

40

50

60

643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-130

-125

-120

-115

-110

-105

-100

-20

-10

0

10

20

30

40

50

60

643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



115



-152

-150

-148

-146

-144

-142

-140

-10

0

10

20

30

40

50

643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-152

-150

-148

-146

-144

-142

-140

-5

0

5

10

15

20

25

30

35

643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



116





Precooling Phase heat exchanger HX1 at stream N2LIQ inlet. It is seen that with

increase in specific exergy flow rate, the total exergy flow increases while heat load

decreases. Additionally, lower temperatures correspond to higher values of specific

exergy flow.

-195.805

-195.8

-195.795

-195.79

-195.785

-195.78

-195.775

0

200

400

600

800

1000

1200

760.2 760.1 711.5 662.8 614.1 565.5 516.8 468.1 419.5 370.8 322.1 273.4 224.8

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-195.805

-195.8

-195.795

-195.79

-195.785

-195.78

-195.775

0

200

400

600

800

1000

1200

760.2 760.1 711.5 662.8 614.1 565.5 516.8 468.1 419.5 370.8 322.1 273.4

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



117








-195.805

-195.8

-195.795

-195.79

-195.785

-195.78

-195.775

0

200

400

600

800

1000

1200

760.2 760.1 711.5 662.8 614.1 565.5 516.8 468.1 419.5 370.8 322.1 273.4

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-195.805

-195.8

-195.795

-195.79

-195.785

-195.78

-195.775

0

200

400

600

800

1000

1200

760.2 760.1 711.5 662.8 614.1 565.5 516.8 468.1 419.5 370.8 322.1 273.4

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



118
















This section deals with the variation of Heat Load, exergy, and temperature for

the liquefaction phase of systems S3A and S3B. Figure 5.90 shows heat load, Exergy

flow vs Temperature for Liquefaction Phase heat exchanger HX2 at stream 49 inlet for

System 3A. It can be seen that total exergy flow increases with increase in specific


-180

-175

-170

-165

-160

-155

-150

-145

-140

-135

0

5

10

15

20

25

30

35

40

760.2 760.1 711.5 662.8 614.1 565.5 516.8 468.1 419.5 370.8 322.1 273.4

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



119


lower temperatures.






for Liquefaction Phase heat exchanger HX2 at stream 13 inlet for System 3A. Much

like in Figure 5.92 it can be seen that total exergy flow increases with increase in

-160

-140

-120

-100

-80

-60

-40

-20

0

-40

-20

0

20

40

60

80

100

120

140

643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-132

-130

-128

-126

-124

-122

-120

-118

-50

0

50

100

150

200

250

300

350

643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



120

specific exergy flow while heat load decreases. The variation of heat load is linear,

while total exergy flow has a nonlinear relationship. Higher specific exergy flows are

achieved at lower temperatures. The variation of these parameters for Liquefaction

Phase heat exchanger HX2 at stream inlet S10 is shown in Figure 5.93 the trends are

similar to those at stream inlet 13 and stream inlet 9. However, the nonlinearity in total

exergy flow is more noticeable at stream inlet 26.


exchanger HX2 at stream S10 inlet for System 3A



seen that as with HX1, the specific exergy flow is lower at lower inlet temperatures.

Additionally, the heat load increases linearly with increase in specific exergy flow,

while total exergy flow has a nonlinear, but decreasing relationship with specific exergy

flow.


in Figure 5.94 where the heat load, exergy flow and temperature for Liquefaction Phase





-203

-202.5

-202

-201.5

-201

-200.5

-200

-199.5

-199

-198.5

0

20

40

60

80

100

120

140

160

643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



121

Figure 5.95 and Figure 5.96 illustrate variations of temperature, total exergy

flow, and heat load for liquefaction phase heat exchanger HX3 at stream inlets 32b and

S1 respectively. At stream inlet 32b, the variation of total exergy flow offers a unique













variation of exergy, heat load and temperature for system 3A.

-180

-160

-140

-120

-100

-80

-60

-40

-20

0

0

20

40

60

80

100

120

643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



122


Figure 5.95 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 32b inlet for System 3A

-160

-140

-120

-100

-80

-60

-40

-20

0

-40

-20

0

20

40

60

80

100

120

140

643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-235

-230

-225

-220

-215

-210

-205

-10

0

10

20

30

40

50

60

70

80

643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



123



-245.005

-245

-244.995

-244.99

-244.985

-244.98

-244.975

-244.97

-244.965

-244.96

-244.955

-20

0

20

40

60

80

100

120

140

643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-160

-140

-120

-100

-80

-60

-40

-20

0

-40

-20

0

20

40

60

80

100

120

140

643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



124

Figure 5.98 shows heat load, Exergy flow vs Temperature for Liquefaction

Phase heat exchanger HX2 at stream 49 inlet for System 3B. It can be seen that total

exergy flow increases with increase in specific exergy flow while heat load decreases.

The variation of heat load is linear, while total exergy flow has a nonlinear relationship.

Higher specific exergy flows are achieved at lower temperatures.


Figure 5.99, Figure 5.100 and Figure 5.101 illustrate the variations in

temperature, heat load, and exergy for system 3B for stream inlets S10, 14, and 13.

Figure 5.100 show similar trends. The total exergy decreases with increase in specific

exergy flow while the heat load increases. Specific exergy flow is lower at lower

temperatures. However, the variation of heat load is more noticeable in inlet 14, and

the variation of total exergy flow exhibits a nonlinear pattern in this case. In figure

Figure 5.101 (corresponding to inlet 13) the trends are different. There is a slight

increase in total exergy flow and a slight increase in heat load as the specific exergy

flow increases. The value of specific exergy flow is higher at lower temperatures.

-200

-180

-160

-140

-120

-100

-80

-60

-40

-20

0

-40

-20

0

20

40

60

80

100

120

140

160

3381.5 3457.7 3541.2 3632.7 3732.8 3842.5 3962.6 4094.3 4238.8 4397.7 4572.6 4765.7

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



125



The next set of figures illustrate the variations in exergy, temperature, and heat

load for heat exchanger HX3 in system 3B. Figure 5.102 and Figure 5.103 show the

variations related to stream inlets 22 and 17 respectively. Both exhibit similar trends,

with lower temperatures corresponding to higher values of specific exergy flow. The

heat load in both cases decrease with increase in specific exergy flow while total exergy

flow increases.

The value of specific exergy flow is lower at lower values of temperature in

inlet 17 in Figure 5.103. The heat duty increases with increase very slightly in specific

exergy flow. However, the variation in total exergy flow shows a moderate trend.

However, the total exergy flow is remaining steady with decrease in specific exergy

flow.

-203

-202.5

-202

-201.5

-201

-200.5

-200

-199.5

-199

-198.5

0

20

40

60

80

100

120

140

2858.3 2848.0 2837.7 2827.4 2817.2 2807.1 2797.0 2786.9 2776.9 2766.9 2756.9 2747.0

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)


TOTAL EXERGYFL

Heat duty

Temperature

126









lower temperatures.

-203

-202.5

-202

-201.5

-201

-200.5

-200

-199.5

-199

-198.5

0

20

40

60

80

100

120

140

2858.3 2848.0 2837.7 2827.4 2817.2 2807.1 2797.0 2786.9 2776.9 2766.9 2756.9 2747.0

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-131

-130

-129

-128

-127

-126

-125

-124

-123

-122

-50

0

50

100

150

200

250

300

3661.8 3668.5 3675.2 3681.9 3688.6 3695.4 3702.3 3709.2 3716.1 3723.1 3730.1 3737.2

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



127





The final figure in this series, Figure 5.105 shows the variation in heat load,

Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream

32b inlet for System 3B. Specific exergy flow is lower at higher temperatures in this

case meaning that more power is needed to cool and liquefy. It is seen that the heat load

decreases with increase in specific exergy flow. The variation of total exergy flow

increases with increase in specific exergy flow.

-250

-200

-150

-100

-50

0

-100

-50

0

50

100

150

200

250

300

3737.2 3842.2 3956.8 4081.9 4218.6 4368.2 4532.1 4712.3 4910.7 5129.8 5372.8 5643.2

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-250

-200

-150

-100

-50

0

-1600

-1400

-1200

-1000

-800

-600

-400

-200

0

200

-1388

0.1

-1387

6.6

-1387

3.1

-1386

9.5

-1386

6.0

-1386

2.5

-1385

9.0

-1385

5.5

-1385

2.0

-1384

8.5

-1384

5.0

-1384

1.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



128




exchanger HX3 at stream 32b inlet for System 3B

-245.012

-245.01

-245.008

-245.006

-245.004

-245.002

-245

-244.998

-244.996

-244.994

-1600

-1400

-1200

-1000

-800

-600

-400

-200

0

200

-1388

0.1

-1387

6.6

-1387

3.1

-1386

9.5

-1386

6.0

-1386

2.5

-1385

9.0

-1385

5.5

-1385

2.0

-1384

8.5

-1384

5.0

-1384

1.5

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)



-250

-200

-150

-100

-50

0

-100

-50

0

50

100

150

200

250

300

3737.2 3842.2 3956.8 4081.9 4218.6 4368.2 4532.1 4712.3 4910.7 5129.8 5372.8 5643.2

Tem

pera

ture

(˚C

)

Hea

t Dut

y an

d Ex

ergy

Flo

w (k

W)


TOTAL EXERGYFL

Heat duty

Temperature

129

5.5 Property set

In the property set in Figure 5.106 the T-xy diagram (vapor mole fraction versus liquid

mole fraction) for vapor-liquid equilibrium (VLE) is shown for a Hydrogen mixture

being fed to the cycle and it shows that hr composition changes at temperatures between

-253 °C and -245 °C at the liquid phase. Figure 5.107 shows the T-x diagram. K-values

for Vapor-liquid and fraction of Para-hydrogen and Ortho Hydrogen is shown in Figure

5.108 and Figure 5.109 illustrates the y-x diagram for vapor vs liquid composition for

the para-hydrogen

Figure 5.106 T-xy plot for temperature versus liquid composition for isobaric data

Figure 5.110 shows the activity coefficients vs mole fraction for Para-hydrogen

and orthohydrogen and illustrates the point of where ortho and para hydrogen have

similar molar fraction with Activity Coefficient ~=1.2.

Liquid/vapor mole fraction, HYDRO-01

Tem

pera

ture

, C

0.000 0.025 0.050 0.075 0.100 0.125 0.150 0.175 0.200 0.225 0.250 0.275 0.300 0.325 0.350 0.375 0.400 0.425 0.450 0.475 0.500 0.525 0.550 0.575 0.600 0.625 0.650 0.675 0.700 0.725 0.750 0.775 0.800 0.825 0.850 0.875 0.900 0.925 0.950 0.975 1.000-253.0

-252.5

-252.0

-251.5

-251.0

-250.5

-250.0

-249.5

-249.0

-248.5

-248.0

-247.5

-247.0

-246.5

-246.0

-245.5

-245.0x 1.0133 bary 1.0133 bar

130

Figure 5.107 T-x plot for temperature versus liquid composition for isobaric data

Figure 5.108 K-values for Vapor-liquid vs fraction of Para-hydrogen and Ortho

Hydrogen

T-x diagram for HYDRO-01/HYDRO-02

Liquid mole fraction, HYDRO-01

Tem

pera

ture

, C

0.000 0.025 0.050 0.075 0.100 0.125 0.150 0.175 0.200 0.225 0.250 0.275 0.300 0.325 0.350 0.375 0.400 0.425 0.450 0.475 0.500 0.525 0.550 0.575 0.600 0.625 0.650 0.675 0.700 0.725 0.750 0.775 0.800 0.825 0.850 0.875 0.900 0.925 0.950 0.975 1.000-253.0

-252.5

-252.0

-251.5

-251.0

-250.5

-250.0

-249.5

-249.0

-248.5

-248.0

-247.5

-247.0

-246.5

-246.0

-245.5

-245.0

1.0133 bar

K-values for HYDRO-01/HYDRO-02


K-va

lues

0.000 0.025 0.050 0.075 0.100 0.125 0.150 0.175 0.200 0.225 0.250 0.275 0.300 0.325 0.350 0.375 0.400 0.425 0.450 0.475 0.500 0.525 0.550 0.575 0.600 0.625 0.650 0.675 0.700 0.725 0.750 0.775 0.800 0.825 0.850 0.875 0.900 0.925 0.950 0.975 1.0000.0

0.5

1.0

1.5

2.0

2.5

3.0

3.5

4.0

4.5

5.0

5.5

6.0

6.5

7.0

7.5

8.0

Liquid1 HYDRO-01 1.0133 barLiquid1 HYDRO-02 1.0133 bar

131

Figure 5.109 y-x diagram for vapor vs liquid composition for the para-hydrogen

Figure 5.110 Activity coefficients vs mole fraction for Para-hydrogen and orthohydrogen

5.6 Comparative analysis results

The exergy efficiencies of the systems are looking quite low when T0 is changed. The

systems were simulated at 0°C, 10°C, 25°C, and 45°C. System S2A shows the highest

exergy efficiency at 40% as shown in Figure 5.111.

y-x diagram for HYDRO-01/HYDRO-02


Vapo

r mol

e fra

ctio

n, H

YDRO

-01

0.00 0.05 0.10 0.15 0.20 0.25 0.30 0.35 0.40 0.45 0.50 0.55 0.60 0.65 0.70 0.75 0.80 0.85 0.90 0.95 1.000.00

0.05

0.10

0.15

0.20

0.25

0.30

0.35

0.40

0.45

0.50

0.55

0.60

0.65

0.70

0.75

0.80

0.85

0.90

0.95

1.00

1.0133 bar

Activity coefficients for HYDRO-01/HYDRO-02


Activ

ity c

oeffi

cient

s

0.000 0.025 0.050 0.075 0.100 0.125 0.150 0.175 0.200 0.225 0.250 0.275 0.300 0.325 0.350 0.375 0.400 0.425 0.450 0.475 0.500 0.525 0.550 0.575 0.600 0.625 0.650 0.675 0.700 0.725 0.750 0.775 0.800 0.825 0.850 0.875 0.900 0.925 0.950 0.975 1.0001.00

1.05

1.10

1.15

1.20

1.25

1.30

1.35

1.40

1.45

1.50

1.55

1.60

1.65

1.70

1.75

1.80

1.85

1.90

1.95

2.00

2.05

2.10

2.15

2.20

2.25

2.30

2.35

2.40

Liquid1 HYDRO-01 1.0133 barLiquid1 HYDRO-02 1.0133 bar

132

Figure 5.111 Exergy efficiency for the proposed hydrogen liquefaction systems at 0°C, 10°C, 25°C, and 45°C.

Figure 5.112 shows system S2A shows the highest energy efficiency at 76%

followed by S3B then S3A. Energy efficiency with Vortex tubes is creating better

efficiencies overall. Similarly, work done per unit liquefaction is highest for S3B and

lowest for S3A when compared to peer systems as shown in Figure 5.113.

Figure 5.112 Energy efficiency for the hydrogen liquefaction systems.

12.0

11.9 16

.2

42.0

6.0

21.0

15.4

13.4 17

.7 19.3 24

.0

10.6

17.5

16.8

11.9 17

.4

18.9 22

.4

10.4 17

.2

16.5

9.4

8.0

5.0

12.0

4.0 6.

7

2.3

0

10

20

30

40

50

Main S1A S1B S2A S2B S3A S3B

Exer

gy E

ffici

ency

(%)

Exergy effeciency (0°C) Exergy effeciency (10°C)

Exergy effeciency (25°C) Exergy effeciency (45°C)

10 12

32

76

7

54

74

0

10

20

30

40

50

60

70

80

Main S1-A S1-B S2-A S2-B S3-A S3-B

Ener

gy E

ffici

ency

(%)

133

Figure 5.113 Work done for liquefaction per unit mass (kJ/kg).

5.7 Optimization results

The optimization model is solved using genetic algorithm utilizing MATLAB

software through the internal software’s calculator and an optimizer.

5.7.1 Objective function

The objective function for optimal operation is simpler than for optimal design,

discussed by Jensen [91], because the investment costs, the capital costs, and others are

not considered. The simplified cost function to reduce total compressor consumption to

be minimized then becomes:

min �� + 𝐶

max 𝐸𝑥

subject to ��[,,& = 𝑔𝑖𝑣𝑒𝑛 = 3.628𝑘𝑔/ℎ𝑟

𝑐 ≤ 0

Here, �� is the sum of all compressor powers (kW). 𝑐 ≤ 0represents the

mathematical formulation of the operational constraints and the model equations. And

feed ��[,,& is maintained at the nominal feed rate.

5.7.2 Design conditions

Feed hydrogen gas stream: normal hydrogen gas enters with P = 1 bar and T = 25 oC

after gas.

15941.215017.9

12688.1

3403.9

1657.8

17216.1

0.0

2000.0

4000.0

6000.0

8000.0

10000.0

12000.0

14000.0

16000.0

18000.0

20000.0

S1A S1B S2A S2B S3A S3B

Wor

k (k

w)

134

Nominal flow rate is 3.628 TPD = 3628 kg/24 hours = 151.2 kg/hour =

0.042 kg/s.

Product: 95% para-liquid hydrogen is at P = 1.3 bar and T = −253 oC equivalent

to the product at Ingolstadt plant by Bracha et al [32].

Pressure: Pressure drops inside HX1, HX2, and HX3 are assumed to be zero

because the information about design criteria of all heat exchangers is assumed to be

insignificant.

5.7.3 Variables

The number of manipulated variables are based on the number of main components

seen showing changes in the system overall through the sensitivity analyses for each

system. Hydrogen compressor powers, Expansions valves, Flash power are mainly

manipulated. Additional equipment based on the proposed systems were manipulated

to calculation ran to get results.

5.7.4 Constraints

The exergy efficiencies for each system are optimized to be the maximum possible value.

The total cost rates obtained from each system is optimized to have the lowest value to

reduce the cost of the system. Eventually, for each system, the exergy efficiency equation

and the total cost rate equation are combined in a function, at which the exergy efficiency

is divided by the total cost rate and the function is set to be maximized. The constraints of

some selected variables are shown in Table. 5.1, at which the upper and lower bounds are

set based on the available date from previous studies.

Table 5.1 Constraints of Selected Variables Variable Lower Upper Unit

Ambient temperature, 𝑇0 -10 50 °C

Compressor Pressure -10 100 °C

LN Pressure 400 1500 Bar

Heat exchanger minimum approach temperature

5 10 °C

5.8 Optimum case

After running the optimization for the all the systems, results show that slight changes

to model do make significant improvement but integrating the systems make a great

improvement. In the case of system 2B along with 3B, it can be found that adding

135

Vortex tubes with ORCs increased the exergy efficiency to 27% with the configuration

shown in Figure 5.115 with at work done per liquefaction mass at 3476.3(kJ/kg).

Figure 5.114 Overall exergy efficiency

5.9 Simulation comparison

For validating the liquefaction system Krasae-In et. Al. [15] experiment data is

compared with the simulation model assumption and setup. Table 5.2 indicates that the

simulation and experimentally measured power consumptions of the two compressors

were equal because the simulation data was calculated using the experimental data. The

compressor power was calculated from the flow rate, inlet and outlet pressures, and

temperatures. According to the law of conservation of energy, the calculated brake

horsepower was the same as the measured one. The hydrogen gas flow rate from the

measurement was only 0.6 kg/h instead of the initially designed 2.0 kg/h.

The main conclusion is that the compressor power and liquefier efficiency were

the same as the simulation data. Although the test rig was capable of cooling hydrogen

gas using the refrigeration system, it was only able to reach a temperature of −158 °C

instead of the designed value of −230 °C.

As the proposed plans are larger that the one in the experiment by [15], the

proposed simulation data are at 1:3 ratio with ±10%change.

17 19

22

10

17 16

27

0%

5%

10%

15%

20%

25%

30%

S1A S1B S2A S2B S3A S3B OPTIMUM

Effic

ienc

y

136

Figure 5.115 Optimized system

137

Table 5.2 Simulation and initial experimental data of the proposed system Parameter Simulation data Experimental data

H2 mass flow 3.628 kg/h 0.6 kg/h

H2 compressor power 0.4 kW 0.067 kW

Isentropic efficiency of compressor 80% 80%

Actual work 6.8 kW 1.127 kW

138

CONCLUSIONS AND RECOMMENDATIONS

6.1 Conclusions

This report gives a brief view of the prime movers of the hydrogen economy, hydrogen

liquefaction and the advanced liquefaction systems. In the literature review, a fair

number of research papers discussing liquid hydrogen production and basic liquefaction

processes and cycles have been discussed. The objective of this work is to propose

different hydrogen liquefaction systems based on an existing one created by Praxair

Inc. A comprehensive thermodynamic is then performed. Energy and exergy

efficiencies will be analyzed to assess and make further improvements. System

feasibility will be assessed with Economic and environmental evaluation. Operating

conditions will be optimized to generate the best scenarios with and will be validated.

The results for the analyzed cases show that there is room for improvement of

on advanced liquefaction system and novel work can be produced that can overcome

efficiency challenges.

• The overall energy and exergy efficiencies of systems 1A and 1B are found to

be 16.69 % and 18.87 % respectively.

• The work done per liquefaction mass of systems 1A and 1B are found to be

15941.20 kJ/kg and 15017.91 kJ/kg respectively.









• The overall energy and exergy efficiencies of the optimized system is found to

be 26.89 % and the work done per liquefaction mass is 3476.34 kJ/kg.

139

• The power obtained and supplied from and to the optimized system was

sufficient to provide nearly half of the power required for compressor work.

6.2 Recommendations

Regarding the future development of study liquefaction cycles of increasing

complexity. An attempt was made create VTs, TEs and ORCs. The expander added

extra variables that made the cycle much more complex. The recommendations for the

future studies are as follows; the use of chemical reactions in the simulations may create

more realistic conclusions.

• Base system simulation cycle indicated that the TEs may improve the cycles

efficiency. However, modified cycle designs could yield more positive results.

Testing TEs in different parts may achieve interesting results.

• In view of this study, it can be inferred that more cryogenic experiments should

be conducted using a VT. Experimentation has the possibility to determine the

expected range of effectiveness values for a cryogenic VT, under any condition.

• Optimization of the proposed large-scale plant explained is simplified and it is

preliminary. More information is still required for more complicated work. It is

a must that there is a study about computer simulation work deep inside about

optimization of the new more efficient cycle.

• For the proposed systems, performing dynamic modeling should be done to

evaluate their actual performances and the released emissions during the

different phases of the driving cycles.

• Considering the hydrogen compression method to compress liquid hydrogen is

required to allow for storing hydrogen onboard at very high pressures.

• Conducting life cycle assessments for the proposed systems to confirm that the

operational emissions produced by the proposed systems are feasible for

hydrogen and sustainable fuel source.

140

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