Analysis and Assessment of Advanced Hydrogen Liquefaction Systems

ANALYSIS AND ASSESSMENT OF ADVANCED HYDROGEN LIQUEFACTION SYSTEMS

Anwar Hammad

A Thesis Submitted in Partial Fulfillment

of the Requirement for Doctor of Philosophy Degree

Mechanical Engineering

Faculty of Engineering and Applied Science

University of Ontario Institute of Technology

Oshawa, Ontario, Canada

September 2019

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 Dr. Haoxiang Lang

University Examiner 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.

i 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

ii 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.

iii 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

iv 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.

v 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 vi 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 vii 5.8 Optimum case 134 5.9 Simulation comparison 135

CONCLUSIONS AND RECOMMENDATIONS 138 6.1 Conclusions 138 6.2 Recommendations 139

REFERENCES 140

viii NOMENCLATURE c Unit cost of exergy, $/kJ E Energy, kJ ex Specific exergy, kJ/kg Eẋ Exergy rate, kW ��̇ Exergy destruction rate, kW h Specific enthalpy, kJ/kg K Equilibrium constant m Mass, kg �̇ Mass flow rate, kg/s � 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 � Specific volume, m3/kg Y Yield, % �̇ Capital cost, $

Greek Letters � Density, kg/m3 � Energy efficiency � 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

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

x 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

xi 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 (��) and cold (��) and hot (�ℎ��) 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

xii 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 5.5 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 5.14 Heat Exchanger HX2 Heat composite curves ...... 71

Figure 5.15 Heat Exchanger HX3 Heat composite curves ...... 72

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

xiii Figure 5.22 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream N2LIQ inlet for System 1A ...... 77

Figure 5.23 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 28 inlet for System 1A ...... 78

Figure 5.24 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 9 inlet for System 1A ...... 78

Figure 5.25 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 36 inlet for System 1A ...... 79

Figure 5.26 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 44 inlet for System 1A ...... 79

Figure 5.27 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger 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

Figure 5.29 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 28 inlet for System 1B ...... 81

Figure 5.30 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 9 inlet for System 1B ...... 81

Figure 5.31 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 36 inlet for System 1B ...... 82

Figure 5.32 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 44 inlet for System 1B ...... 82

Figure 5.33 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX2 at stream 49 inlet for System 1A ...... 83

Figure 5.34 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX2 at stream 13 inlet for System 1A ...... 84

Figure 5.35 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream S10 inlet for System 1A ...... 84

Figure 5.36 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 1A ...... 85

Figure 5.37 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 22 inlet for System 1A ...... 86

xiv Figure 5.38 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 32b inlet for System 1A ...... 86

Figure 5.39 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream S1 inlet for System 1A ...... 87

Figure 5.40 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 17 inlet for System 1A ...... 87

Figure 5.41 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream S10 inlet for System 1B ...... 88

Figure 5.42 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 1B ...... 88

Figure 5.43 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 13 inlet for System 1B ...... 89

Figure 5.44 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 49 inlet for System 1B ...... 90

Figure 5.45 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 22 inlet for System 1B ...... 90

Figure 5.46 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 17 inlet for System 1B ...... 91

Figure 5.47 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream S1 inlet for System 1B ...... 91

Figure 5.48 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger 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

Figure 5.53 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream N2LIQ inlet for System 2A ...... 97

xv Figure 5.54 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 28 inlet for System 2A ...... 97

Figure 5.55 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 9 inlet for System 2A ...... 98

Figure 5.56 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 36 inlet for System 2A ...... 98

Figure 5.57 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 44 inlet for System 2A ...... 99

Figure 5.58 Heat Load, Exergy flow vs Temperature for Precooling Phase heat exchanger HX1 at stream N2LIQ inlet for System 2B ...... 99

Figure 5.59 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 28 inlet for System 2B ...... 100

Figure 5.60 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 9 inlet for System 2B ...... 100

Figure 5.61 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 36 inlet for System 2B ...... 101

Figure 5.62 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 44 inlet for System 2B ...... 102

Figure 5.63 Heat Load, Exergy flow and Temperature for liquefaction Phase heat exchanger HX2 at stream 49 inlet for System 2A ...... 102

Figure 5.64 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 13 inlet for System 2A ...... 103

Figure 5.65 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 26 inlet for System 2A ...... 103

Figure 5.66 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 2A ...... 104

Figure 5.67 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 22 inlet for System 2A ...... 105

Figure 5.68 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 32b inlet for System 2A ...... 105

Figure 5.69 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream S1 inlet for System 2A ...... 106

xvi Figure 5.70 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 17 inlet for System 2A ...... 106

Figure 5.71 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 49 inlet for System 2B ...... 107

Figure 5.72 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 26 inlet for System 2B ...... 108

Figure 5.73 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 2B ...... 108

Figure 5.74 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 13 inlet for System 2B ...... 109

Figure 5.75 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 22 inlet for System 2B ...... 109

Figure 5.76 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 17 inlet for System 2B ...... 110

Figure 5.77 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream S1 inlet for System 2B ...... 110

Figure 5.78 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 32b inlet for System 2B ...... 111

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

Figure 5.80 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream N2LIQ inlet for System 3A ...... 113

Figure 5.81 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 28 inlet for System 3A ...... 114

Figure 5.82 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 9 inlet for System 3A ...... 114

Figure 5.83 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 36 inlet for System 3A ...... 115

Figure 5.84 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 44 inlet for System 3A ...... 115

Figure 5.85 Heat Load, Exergy flow vs Temperature for Precooling Phase heat exchanger HX1 at stream N2LIQ inlet for System 3B ...... 116

xvii Figure 5.86 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 28 inlet for System 3B ...... 116

Figure 5.87 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 9 inlet for System 3B ...... 117

Figure 5.88 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 36 inlet for System 3B ...... 117

Figure 5.89 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 44 inlet for System 3B ...... 118

Figure 5.90 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 49 inlet for System 3A ...... 119

Figure 5.91 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 13 inlet for System 3A ...... 119

Figure 5.92 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream S10 inlet for System 3A ...... 120

Figure 5.93 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 3A ...... 121

Figure 5.94 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 22 inlet for System 3A ...... 122

Figure 5.95 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 32b inlet for System 3A ...... 122

Figure 5.96 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream S1 inlet for System 3A ...... 123

Figure 5.97 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 17 inlet for System 3A ...... 123

Figure 5.98 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 49 inlet for System 3B ...... 124

Figure 5.99 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream S10 inlet for System 3B ...... 125

Figure 5.100 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 3B ...... 126

Figure 5.101 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 13 inlet for System 3B ...... 126

xviii Figure 5.102 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 22 inlet for System 3B ...... 127

Figure 5.103 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 17 inlet for System 3B ...... 127

Figure 5.104 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream S1 inlet for System 3B ...... 128

Figure 5.105 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger 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

xix 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.

900 Current Forecast 800 700 600 500 400 300

Energy (QuadrillionEnergy kJ) 200 100 0 1990 2000 2012 2020 2030 2040 Year

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 1 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.

180000 160000 140000 120000 100000 80000

Energy (TWh) Energy 60000 40000 20000 0 1990 1991 1992 1993 1994 1995 1996 1997 1998 1999 2000 2001 2002 2003 2004 2005 2006 2007 2008 2009 2010 2011 2012 2013 2014 2015 2016 2017 Year

Coal Solar Crude oil Natural gas Traditional biofuels Other renewables Hydropower Nuclear

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].

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.

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].

O2 and/or Air Steam

Feedstock in: Cool/ Gasifier Gas Clean up Biomass Shift Reactor Syngas Partial OXI Natural gas/ (H2, CO, CO2, H2O) (Water- gas shift (H and CO H ,O, H , CO, CO , 2 2 fuel oil (POX) 2 2 2 separation) shift to PSA, physical absorption H and CO ) Natural gas/ Reformer 2 2 naphtha (SMR/ATR)

CO2: O2 and/or Air Steam - Vented - To urea production H2 and CO2 - Enhanced oil recovery to ammonia and F-T Processes 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.

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 Exergy efficiency after electrolysis (%) 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 5 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 6 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].

Vin

Vhot

Vcold

Figure 1.5 Schematic of the Vortex Tube, showing the inlet (��) and cold (��) and hot (�ℎ��) 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.

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:

• 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

9 optimization study for each system are also provided. Chapter 6 highlights conclusions from the thesis together with recommendations for future studies.

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

11 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 precooling 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

12 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 13 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 14 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.

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

16 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.

• 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 18 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.

Table 2.1 Details of Commercial Hydrogen Liquefaction Plants Capacity Location Operated by Country Tons Per Day Ontario Air Products 30.0 Air Liquide 10.0 Canada Quebec Air Liquide 12.0 BOC by Linde 15.0 BOC by Linde 14.0 French Guyana Kourou Air Liquide 5.0 Missouri Air Products 3.0 Air Products 3.2 Florida Air Products 27.0 Mississippi Air Products 32.7 California Union Carbide/Linde Division 54.0 Louisiana Air Products 34.0 USA 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 Linde 4.4 Germany Linde 5.0 Netherlands Air Products 5.0 China CALT 0.6 ISRO 0.3 India Asiatic Oxygen 1.2 Andhra Sugars 1.2 Iwatani 0.6 MHI 0.7 Tashiro 1.4 Pacific Hydrogen 1.4 Japan 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

20 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. 21

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

22 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]).

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].

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].

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.

Electron

Proton

Ortho Hydrogen Para Hydrogen

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.

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

27 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

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.

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.

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.

Main Liquefaction System

Advanced Liquefaction Advanced Liquefaction System with Advanced Liquefaction System (S1) CO2 Orgainc Rankine Cycle (S2) System with Vortex tubes (S3)

Without Without Without With Catalyst With Catalyst With Catalyst Catalyst (S1- Catalyst (S2- Catalyst (S2- (S1-B) (S2-B) (S2-B) A) A) A)

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 31 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, o a liquid H2 product at the atmospheric pressure and temperature of -253 C.

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

32 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

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.

Legend: HX : Heat Exchanger AO : Adsorption Unit B : Boiler C : Compressor D : Split M : Mixer/Divider TE : Turbo Expander F : Flash V : Valve

Figure 3.2 The main systems schematic diagram

Figure 3.3 Schematic diagram for the reference system

Figure 3.4 Schematic diagram for the main system with reactor

3.2 Description of System 2

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

S12 140 kJ in which the total system efficiency can be improved. Figure 3.5 illustrate the

S6 4 S5 ORCs that are added to the system. 2 C2 46 S12 M1 E-51 10 H2FEED S3 S7 4 S12 N2GAS S6 8 E-58 S5 EX1 2 C2 46 S6 4 E-53 M2 M1 46B E-51 10 2 S5 C2 H2FEED 40 S3 S7 46 E-57 M1 E-51 S11 E-56 N2GAS S15 10 S3 8 E-58 S3 9 39 B13 H2FEED S7 EX1 46B N2GAS E-53 S8 M2 8 E-58 11 P-93 E-52 EX1 40 46B Ex3 E-57 E-53 S15 M2 HX1 S3 S11 B11 E-56 C1 N2LIQ 9 40 E-59 39 B13 36 E-57 S15 S13 S3 S11 E-56 41 11 S8 9 3 PD-191 P-93 39 B13 28 E-52 1 Ex3 11 S8 14 HX 1 B11 P-93 C1 M3 N2LIQ E-61 E-52 29 E-59 S4 36 Ex3 HX1 B11 32 C1 41 N2LIQS13 (a) (b) P-91 (c) E-59 3 12 13 D1 M4 28 21 1 36 43 41 S13 D5 P-91 S9 14 M3 3 D1 19 29 E-61 28 1 HX2 S4 EX2 30 Figure 3.5 Organic Rankine Cycles 51 20 14 M3 E-61 D2 32 29 17 C5 12 13 S4 44 33 M4 21 B6 4943 D3 27 For each ORC, the work outcome could make a significant positive impact on 32 AO D5 19 S9 12 13 M4 21 D4 HX2 43 the system’s overall efficiency EXand2 can be 30optimized for future improvements. In a 34 51 20 S9 D5 D2 19 35 17 C5 HX2 further improvement to the efficiency of the system, three ORCs are added to the EX2 30 44 B6 31 49 D3 33 51 20 HX3 AO 27 D2 33 26 22 compressors. The ORC, which represents the most commonly used low heat source 17 C5 44 B6 49 D3 33 C4 D4 25 27 34 AO 24 temperature-based system, consists of the same components as a conventional steam V-1 D4 35 S1 F1 23 power plant. However, the working fluid is an organic component that exhibits a lower31 34 V-2 PROD HX3 33 26 22 35 31 C4 38 25 HX3 33 26 22 24 V-1 S1 F1 C4 23 25

V-2 PROD 24 V-1 S1 F1 23

V-2 PROD 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.

Figure 3.6 Schematic diagram for the system with ORCs

Figure 3.7 Schematic diagram for the system ORCs and reactor

3.3 Description of System 3

In a further improvement to the efficiency of the system, Vortex tubes before the exchanger inlet. Two configurations are explained further.

2 M1 H2FEED 4 C2 EX1 46

S3 N2GAS

9 3.3.1 System 3AC1 : System with Vortex Tubes and without a catalystN2LIQ HX1

36 In this system, Vortex Tubes3 (VTs) were added using a splitter and a turbo expander to D1 28 1 14 simulate the splitting and cooling effects of the 29VTs. Apart from the addition of VTs, M3 S4 13 the system layout corresponds to that of the base system. The split was simulated to be 12 43 M4 21 at 50% and the turbo expander outlet pressure was simulated to be 15 bars. Adding a D5 19 S9 VT in this manner resulted in additional cooling capacity in the system. There is an HX2 30 EX2 20 D2 optimization opportunity to split more of the stream through to the VT, and to further 17 C5 44 49 D3 33 27 reduce the outlet pressure on the expander to increaseAO cooling capacity. Figure 3.9 D4 shows the VTs added to the main system. 34

35 51 32 31 HX3 text 44 51 33 26 22

VT1 C4 25 24

V-1 S1 F1 23 P-78 78 V-2 32 PROD VT2

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. 42

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

Figure 3.10 Schematic diagram for the system with VTs and reactor

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:

o • The reference temperature is selected as T0 = 25 C (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 45 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:

∑ � − ∑ � = � − � (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:

= ∑ �̇ − ∑ �̇ (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:

∑(� + ��)� − ∑(� + ��)� + ∑(�), − (�), = � − � (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:

̇ ̇ ̇ ̇ � + � + ∑ �̇ ℎ = � + � + ∑ �̇ 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: 46

̇ = ∑ �̇ � − ∑ �̇ � + ∑ + �̇ (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]:

= ∑ �̇ � − ∑ �̇� + ∑ �̇ � − ∑ �̇ � − ��̇ (4.6) where �̇ � denotes the rate of exergy transfer with the heat energy exchange across the system volume. �̇ � denotes the rate of exergy transfer by the boundary or work applied on or done by the system. The term �̇ � 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 ��̇ .

The exergy transfer due to the exchange of heat with the environment can be formulated as follows:

Eẋ = Q̇ 1 − (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:

�̇ � = �̇ + � = ∑ �̇ − ∑ �̇ (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:

�� = � (4.9)

�� = � (4.10)

�� = �x = ∑� − �� (4.11)

�� = (� − �) − �(� − �) (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:

�� = ℎ − ℎ − �(� − �) + �� (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:

̇ ̇ �� = ��, (4.14) where Ṡ ,denotes 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:

̇ �̇ = ∑ �̇ � − ∑ �̇ � − ∑() (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:

= �̇ � = ∑ 1 − �̇ − �̇ + � + ∑ �̇ �� − ∑ �̇ �� − � �̇ (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:

̇ ̇ ̇ ∑ ∑ �� = �� − �� + �̇ �� − �̇ �� (4.17)

̇ where Ex 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 ex, ex.

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.

Table 4.1. Energy Balance for System Components

Component Balance Equation Component Name Mixer Energy Balance 1 �̇ ℎ + �̇ ℎ = �̇ ℎ M Entropy Balance �̇ � + �̇ � + �̇ = �̇ � 2 , Exergy Balance 3 �̇ �� + �̇ �� = �̇ �� + ��̇ , Compressor Energy Balance 1 ̇ C �̇ ℎ + � = �̇ ℎ Entropy Balance �̇ ℎ + �̇ , = �̇ ℎ Exergy Balance �̇ �� + �̇ = �̇ �� + ��̇ 2 ,

Cooler, Heat Energy Balance HX Exchanger �̇ ℎ + �̇ ℎ = �̇ ℎ + �̇ ℎ 1 4 (Without heat Entropy Balance exchanger �̇ � + �̇ � + �̇ = �̇ � + �̇ � 2 , 5 efficiency) Exergy Balance ̇ 3 6 �̇ �� + �̇ �� = �̇ �� + �̇ �� + ��,

Cooler, Heat Energy Balance Exchanger �̇ ℎ + �̇ ℎ = �̇ ℎ + �̇ ℎ + �̇, (With heat Entropy Balance exchanger �̇ � + �̇ � + �̇ , = �̇ � + �̇ � + efficiency) ̇ , HX �� 1 4 �� 2 5 Exergy Balance �̇ �� + �̇ �� = �̇ �� + �̇ �� + 3 , 6 ��̇ , + ��̇ ̇ , �� = 1 − �̇, �� ̇ , = ��̇ , �� .

Valve Energy Balance �̇ ℎ = �̇ ℎ 1 Entropy Balance �̇ � + �̇ = �̇ ℎ 2 , Exergy Balance V �̇ �� = �̇ �� + ��̇ ,

Flash Energy Balance 2 �̇ ℎ + �̇ ℎ = �̇ ℎ + �̇ ℎ Entropy Balance �̇ � + �̇ � + �̇ , = �̇ � + �̇ � Exergy Balance 1 �̇ �� + �̇ �� = �̇ �� + �̇ �� + F ��̇ , 3

Vortex Tube Energy Balance �̇ ℎ + �̇ = �̇ ℎ 2 1 Entropy Balance �̇ ℎ + �̇ , = �̇ ℎ Exergy Balance VT �̇ �� + �̇ = �̇ �� + ��̇ ,

4.7 Energy and Exergy Efficiencies

The exergy efficiency for the liquefaction process of the system is expressed as follows:

̇ Ψ = (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

51 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:

� = (4.19)

� = = 1 − ∑ (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:

̇ � = (4.21) ̇

There are two other, commonly used exergy-based equations for the efficiency of a steady-state device:

�� = 1 − (4.22)

�� = (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:

̇ � = = (4.24) ̇ where � is the mass flow rate of the fraction of liquefaction. The total system efficiency is determined by:

.() �(%) = . 100 (4.25) .

The energy per unit mass liquefied in the system is calculated by:

̇ ̇ � = = (4.26) ̇ ̇

For the cooling cycle, the Coefficient of performance (COP) is determined by:

�� = (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]:

�� = (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.

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:

̇ ̇ ̇ ∑ ∑ ̇ �� = �� − �� + �̇ �� − �̇ �� + � (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 �̇ includes all financial costs correlated with ownership, operation and maintenance of any particular system component, and can be defined as [62]:

̇ �̇ = (4.30) × where ϕk is maintenance factor (typically equal to 1.06), �̇ 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:

�̇ = �̇ − �̇ + �̇ + �̇ (4.31) where different �̇ parameters represent the corresponding unit cost rates: �̇ = ��. Application of exergy cost balance equation for every system component yields a set of non-linear equations which can be solved for �̇ or C.

Table 4.2 Base system components exergy equations

Component Exergy Destruction Rate Exergy Efficiency

Cooler 1 Eẋ d,EX1 = ṁ 3ex3 + ṁ 5ex5 (ṁ 4ex4 − ṁ 6ex6) ηex,EX1 = − ṁ 4ex4 (ṁ 3ex3 − ṁ 5ex5) − ṁ 6ex6 Cooler 2 Eẋ d,EX2 = ṁ 8ex8 + ṁ 10ex10 (ṁ 9ex9 − ṁ 11ex11) ηex,EX2 = − ṁ 9ex9 (ṁ 8ex8 − ṁ 10ex10) − ṁ 11ex11 Cooler 3 Eẋ d,EX3 = ṁ 38ex38 + ṁ 40ex40 (ṁ 39ex39 − ṁ 41ex41) ηex,EX3 = − ṁ 39ex39 (ṁ 38ex38 − ṁ 40ex40) − ṁ 41ex41 Expansion Valve 1 Eẋ d,V1 = ṁ 18ex18 − ṁ 16ex16 (ṁ 16ex16) ηex,V1 = (ṁ 18ex18) Expansion Valve 2 Eẋ d,V2 = ṁ 23ex23 − ṁ 24ex24 (ṁ 24ex24) ηex,V2 = (ṁ 23ex23) Compressor 1 Eẋ d,C1 = ṁ 3ex3 + Ẇ in,C1 (ṁ 3ex3 − ṁ 2ex2) ηex,C1 = − ṁ 2ex2 Ẇ in,C1

Compressor 2 Eẋ d,C2 = ṁ 8ex8 + Ẇ in,C2 (ṁ 8ex8 − ṁ 7ex7) ηex,C2 = − ṁ 7ex7 Ẇ in,C2

Compressor 4 Eẋ d,C3 = ṁ 38ex38 + Ẇ in,C3 (ṁ 38ex38 − ṁ 37ex37) ηex,C3 = − ṁ 37ex37 Ẇ in,C3

Turbo Expander 1 Eẋ d,T1 = ṁ 31ex31 − ṁ 30ex30 Ẇ out,GT1 η = − Ẇ ex,T1 out,T1 (ṁ 31ex31 − ṁ 30ex30) Turbo Expander 2 Eẋ d,T2 = ṁ 32ex32 − ṁ 21ex21 Ẇ out,GT2 η = − Ẇ ex,T2 out,T2 (ṁ 32ex32 − ṁ 21ex21) =

Flash Drum 1 Eẋ d,F1 = ṁ s1exs1 + ṁ 25ex25 Ẇ out,ST η = − m ex ex,F1 ̇ s2 s2 ( ṁ s2exs2 − ṁ 24ex24) − ṁ 24ex24 − Ẇ out,F1

Flash Drum 2 Eẋ d,F2 = ṁ s4exs4 + ṁ s5exs5 Ẇ out,ST η = − m ex ex,F2 ̇ s3 s3 (ṁ 56ex56 − ṁ 54ex54) − ṁ s6exs6 − Ẇ out,F2

Heat Exchanger 1 Eẋ d,HX1 = ṁ 1ex1 + ṁ 29ex29 + ṁ 37ex37 + ṁ 42ex42 + ṁ 46ex46 + ṁ N2gasexN2gas + ṁ s5exs5 − ṁ 9ex9

− ṁ 28ex28 − ṁ 36ex36 + ṁ 39ex39 − ṁ 44ex44 − ṁ N2liqexN2liq − ṁ s7exs7

Heat Exchanger 2 ��̇ �,��2 = �̇ 15��15 + �̇ 19��19 + �̇ 27��27 + �̇ 34��34 + �̇ �4��4 − �̇ 6��6 − �̇ 13��13 − �̇ 14��14 + �̇ 26��26 − �̇ 49��49 Heat Exchanger 3 Eẋ d,HX3 = ṁ 18ex18 + ṁ 23ex23 + ṁ 26ex26 + ṁ 33ex33 + ṁ H2liqproexH2liqpro + ṁ s2exs2 − ṁ 15ex15

− ṁ 16ex16 − ṁ 22ex22 − ṁ 25ex25 − ṁ 32ex32 − ṁ s1exs1

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:

� = �̇⁄�̇ (4.32) where �̇ represents the overall financial cost of the system and �̇ is the total exergy loss in the system defined as:

�̇ = ��̇ + ��̇ (4.33) where ��̇ 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]:

̇ ̇ �, = = (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 56 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]:

� = (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)

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:

�� ⟶ � − �� + ℎ�� (4.36)

Conversion percentage for the equilibrium concentration of ortho and para hydrogen is temperature dependent the and follows:

�� (%) = � + � ∗ � + � ∗ � (4.37) 58 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 - -0.220755 0.0377215 0.223621 0.21599 3 Critical Pressure bar 13.13 12.928 34 73.83 Some of Entropies cal/mol- 31.2124 45.765 50.3696 K 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

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.

Exergy Efficiency 100 99 99 99 99 99 100 100 92 88 90 80 80 80 82 81 80 76 68 69 68 66 65 62 60 53 53 49 46 47 40 34 Efficiency (%) 24 26 20

Cooler (EX1)Cooler (EX2)Cooler (EX3) CompressorCompressor (C1) Compressor (C2) (C3) Flash Drum (F1) Expansion ValveExpansion (V1) Valve (V2) Turbo ExpanderTurbo (TE1) Expander (TE2)AdsorptionHeat Unit Exchanger (AO)Heat Exchanger (HX1)Heat Exchanger (HX2) (HX3) System Components

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

Compressor C1

Compressor C2 600 Compressor C3 500

400

300

200 Work Loss (kW) 100

0 -60 -40 -20 0 20 40 60 o H2 Feed temperature ( C)

(a)

3.9

3.8

3.7

3.6

3.5 Yield (Tons/Day) 3.4

3.3

3.2 -50 -40 -30 -20 -10 0 10 20 30 40 50 o H2 Feed Temperature ( C)

(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

20 Exergy efficiency

19 Energy Effeciency

Efficiency (%) 14

10 -60 -40 -20 0 20 40 60 Hydrogen Feed Tempreture (oC)

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

4000 Turbo Expander TE1

3500 Turbo Expander TE2

3000

2500

2000 Work (kW)

1500

1000

500

0 0 1000 2000 3000 4000 5000 6000 7000 H2 Mass flow rate (kg/h)

(a)

Work of Adsorber A0 0 -2 -4 -6 -8 -10 -12 -14 -16 -18 -20 0 1000 2000 3000 4000 5000 6000 7000 (b)

200

180 160

140

120 100

yield (tones/day) 80 2 60 40

Liquid Liquid H 20

0 0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000 5500 6000 6500 7000 7500

H2 Feed Mass Flow Rate (kg/h)

(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.

18 16 15.46 Exergy Effeciency 14 Energy Effeciency 12 10 8

Efficiency (%) 6 4 2 0.32 0 0 2000 4000 6000 8000 Hydrogen Feed Mass Flow Rate, kg/h

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.

990 Turbo Expander TE1 Compressor C1 Compressor C2

Compressor C3 Adsorption Unit A0 Turbo Expander TE2 790

590

390 Work Rate (kW)

190

-10 0 50 100 150 200 Turbo Expander TE1 Pressure, bar

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

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.

Exergy Effeciency Energy Effeciency 13 Efficiency (%) 12

10 0 50 100 150 200 TE1 Pressure (bar)

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 66 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

14 Exergy Effeciency Energy Effeciency 13 Efficiency (%) 12

10 0 1 2 3 4 5 6 7 Flash Drum Pressure (bar)

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

Work of Compressor C2 Work of Compressor C3 Work of compressor C1

800 236.4

700 236.38 600 236.36 500

400 236.34

300 Work (kW)

Work Rate (kW) 236.32 200 236.3 100

0 236.28 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 H2 Feed Pressure (bar) 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.

20 Exergy Effeciency 19 Energy Effeciency 18 17 16 15 14 Efficiency (%) 13 12 11 10 0 5 10 15 20 25 Hydrogen Feed Pressure (bar)

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

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.

16 Exergy efficiency Energy efficiency 15

13 Efficiency (%)

10 0 10000 20000 30000 40000 Nitrogen Gas Mass Flow Rate (kg/h)

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.

150 0

100 -50 C) C) o o 50 -100 0

Temperature ( Temperature ( Temperature -150 -50

-100 -200 -1.2E+08 -80000000 -40000000 0 -5E+08 -3E+08 -1E+08 Heat Duty (kJ/h) Heat Duty (kJ/h) (a) (b)

0 0

-50 -50 C) C) o o -100 -100 -150

Temperature ( Temperature -150 ( Temperature -200

-200 -250 -12000000 -8000000 -4000000 0 1000000 6000000 11000000 16000000 Heat Duty (kJ/h) Heat Duty (kJ/h) (c) (d) C) o -135

-185 Temperature ( Temperature

-235 0 2000000 4000000 6000000 8000000 10000000 12000000 14000000 Heat Duty (kJ/h) (e)

Figure 5.13 Heat Exchanger HX1 Heat composite curves

-180 -170

-190 -180 C) o C)

o -190 -200

-200 -210 -210 Temperature ( Temperature Temperature ( Temperature -220 -220 -230 -1.35E+08 -8.50E+07 -3.50E+07 1.50E+07 -230 -3.60E+07 -2.60E+07 -1.60E+07 -6.00E+06 Heat Duty (kJ/h) Heat Duty (kJ/h) (a) (b)

450

400 -195 C) o C) 350 o -205

300 -215 Temperature ( Temperature 250 ( Temperature -225

200 -235 0.00E+00 1.00E+08 2.00E+08 -2.00E+07 -1.50E+07 -1.00E+07 -5.00E+06 0.00E+00 Heat Duty (kJ/h) Heat Duty (kJ/h) (c) (d)

Figure 5.14 Heat Exchanger HX2 Heat composite curves

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.

-252.6 -234 C) o

C) -252.7 o -239

-252.8 -244 Temperature ( Temperature Temperature ( Temperature -252.9 -249

-254 -253 -1.00E+08 -6.00E+07 -2.00E+07 -1.80E+06 -1.30E+06 -8.00E+05 -3.00E+05 2.00E+05 Heat Duty (kJ/h) Heat Duty (kJ/h) (a) (b)

-228 -234 -232 C) o -239 C)

o -236

-244 -240 Temperature ( Temperature

-249 ( Temperature -244

-254 -248 -2.00E+07 -1.50E+07 -1.00E+07 -5.00E+06 0.00E+00 0.00E+00 2.00E+06 4.00E+06 6.00E+06 8.00E+06 Heat Duty (kJ/h) Heat Duty (kJ/h) (c) (e)

350

-175 250 C) o 150 -195 C) o 50 -215 -50

Temperature ( Temperature -235 -150 Temperature ( Temperature -255 -250 -1.20E+07 -7.00E+06 -2.00E+06 3.00E+06 -3.00E+07 2.00E+07 7.00E+07 1.20E+08 Heat Duty (kJ/h) Heat Duty (kJ/h) (d) (f) Figure 5.15 Heat Exchanger HX3 Heat composite curves

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 overall system exergy efficiency. Different catalyst types and performance enhancement of the heat exchangers could create a dramatic improvement to 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.

100 9999 99 98 92 89 90 83 82 81 85 80 74 69 70 63 64 65 62 60 53 49 50 41 40 34 31 28 Efficiency (%) 30 24 23 Exergy Efficiency 20 Energy Efficiency 10 0

CompressorCompressor (C1)Compressor (C2)Compressor (C3)Compressor (C3)Flash (C3) Drum (F1)

Heat ExchangerHeat ExchangerHeat (HX1) ExchangerExpansion (HX2)Expansion (HX3) Valve (V1) Valve (V2) Adsorption Unit (AO)

(a)

100 99 99 99 99 87 90 77 78 77 78 78 80 74 67 70 60 62 60 55 50 42 40 30 30 Efficiency (%) 30 23 23 23 Exergy Efficiency 18 20 Energy Efficiency 10 0

CompressorCompressor (C1)Compressor (C2)Compressor (C3)Compressor (C3)Flash (C3) Drum (F1)

Heat ExchangerHeat ExchangerHeat (HX1) ExchangerExpansion (HX2)Expansion (HX3) Valve (V1) Valve (V2) Adsorption Unit (AO)

(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%.

25 Exergy Efficiency Energy Efficiency 20 Efficiency (%) 15

10 0 5 10 15 20 25

H2 Feed Pressure (bar)

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

80 70 60

50 Energy Efficiency 40 Exergy Efficiency 30 Efficiency (%) 20 10 0 1.5 2 2.5 3 3.5 4 4.5 Compressor C1 Pressure (bar)

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

17 Energy Efficiency 15 Exergy Efficiency 13

11 Efficiency (%)

7 1 2 3 4 5 6 Compressor C2 Pressure (bar)

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

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.

22 Exergy Efficiency 21 Energy Efficiency 20

19 Efficiency (%)

17 0 5 10 15 20 25 Pressure (bar)

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

35 33 31 29 27 25 23 Exergy Efficiency

Efficiency (%) 21 Energy Efficiency 19 17 15 2 2.5 3 3.5 4 Pressure (bar)

Figure 5.21 Effects of Compressor C2 pressure on overall energy and exergy efficiencies for System 1B

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.

1200 -195.775

Total Exergy 1000 -195.78 Heat Load Temperature 800 -195.785

600 -195.79

400 -195.795 Temperature (˚C)

200 -195.8 Heat Load and Exergy Flow (kW)

0 -195.805 760.2 760.1 709.6 658.9 608.3 557.7 507.1 456.5 405.9 355.3 304.7 254.1 203.5 Specific Exergy Flow (kJ/kg)

Figure 5.22 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream N2LIQ inlet for System 1A

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.

30 0 25 -20 20 15 -40 10 Total Exergy -60 5 Heat Load -80 0 Temperature

-5 -100 (˚C) Temperature -10 -120 Heat Load and Exergy Flow (kW) -15 -20 -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 Specific Exergy Flow (kJ/kg)

Figure 5.23 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 28 inlet for System 1A

400 150

300 100 200 Total Exergy C)

Heat Load 50 o 100 Temperature 0 0

-100 -50 Temperature ( Temperature -200 -100 -300 Heat Load and Exergy Flow (kW)

-400 -150 3903.8 3822.1 3756.5 3709.2 3682.8 3680.5 3706.6 3766.2 3866.3 4016.2 4228.7 4522.1 Specific Exergy (kj/kg)

Figure 5.24 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 9 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

78 flow rises, heat load increases and exergy decreases. Additionally, higher specific energy flows correspond to higher inlet temperatures in both cases.

70 0 -20 60 -40 50 -60 Total Exergy -80 40 Heat Load Temperature -100 30 -120 -140 20 (˚C) Temperature -160 10 -180 Heat Load and Exergy Flow (kW) 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 Specific Exergy Flow (kJ/kg)

Figure 5.25 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 36 inlet for System 1A

45 0 40 -20 35 -40 -60 30 Total Exergy -80 25 Heat Load -100 20 Temperature -120 15 -140 Temperature (˚C) Temperature 10 -160 5 -180 Heat Load and Exergy Flow (kW) 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 Specific Exergy Flow (kJ/kg)

Figure 5.26 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 44 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 79 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.

100 0 -20 80 -40 -60 60 Total Exergy -80 40 Heat Load -100 Temperature -120 20 -140 Temperature (˚C) -160 0 -180 Heat Load and Exergy Flow (kW) -20 -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 Specific Exergy Flow (kJ/kg)

Figure 5.27 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream R inlet for System 1A

1200 -195.775

1000 -195.78

800 -195.785

600 Total Exergy -195.79 Heat Load 400 Temperature -195.795 Temperature (˚C)

200 -195.8 Heat Load and Exergy Flow (kW) 0 -195.805 760.2760.1709.2658.3607.3556.4505.5454.5403.6352.6301.7250.8199.8 Specific Exergy Flow (kJ/kg)

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.

30 0 25 -20 20 15 -40 10 Total Exergy -60 5 Heat Load 0 -80 Temperature -5 -100 -10 Temperature (˚C) -120 -15 -20 -140 Heat Load and Exergy Flow (kW) 396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2 Specific Exergy Flow (kJ/kg)

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.

400 150 300 100 200 50 100 Total Exergy 0 Heat Load 0 Temperature -100 -50

-200 Temperature (˚C) -100 -300 -400 -150 Heat Load and Exergy Flow (kW) 3042.8 2961.9 2897.0 2850.3 2824.5 2822.6 2849.0 2908.7 3008.6 3158.0 3369.6 3661.8 Specific Exergy Flow (kJ/kg)

Figure 5.30 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 9 inlet for System 1B

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 81 case, it can be noted that the relationship between specific exergy flow and all three other variables is linear.

60 -135

-140 50 -145

40 -150

Total Exergy -155 30 Heat Load Temperature -160 20 -165 Temperature (˚C)

-170 10 Heat Load and Exergy Flow (kW) -175

0 -180 396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2 Specific Exergy Flow (kJ/kg)

Figure 5.31 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 36 inlet for System 1B

40 -135

35 -140 -145 30 -150 25 Total Exergy -155 20 Heat Load Temperature -160 15 -165 Temperature (˚C) 10 -170

Heat Load and Exergy Flow (kW) 5 -175

0 -180 396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2 Specific Exergy Flow (kJ/kg)

Figure 5.32 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 44 inlet for System 1B

Figure 5.32 shows the variation of temperature, exergy flow and heat load for precooling heat exchanger HX1 at stream 44 inlet. With the increase in specific exergy

82 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.

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.

200 0 -20 150 -40 -60 100 -80 Total Exergy 50 -100 Heat Load Temperature -120 0 -140 Temperature (˚C) -160 -50 Heat Load and Exergy Flow (kW) -180 -100 -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 Specific Exergy Flow (kJ/kg)

Figure 5.33 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX2 at stream 49 inlet for System 1A

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. 83

500 0 -20 400 -40 -60 300 Total Exergy -80 200 Heat Load -100 Temperature -120 100 -140 Temperature (˚C) -160 0 Heat Load and Exergy Flow (kW) 401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1 -180 -100 -200 Specific Exergy Flow (kJ/kg)

Figure 5.34 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX2 at stream 13 inlet for System 1A

140 0

120 Total Exergy -50 100 Heat Load Temperature 80 -100

60 -150

40 Temperature (˚C) -200 20 Heat Load and Exergy Flow (kW) 0 -250 401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1 Specific Exergy Flow (kJ/kg)

Figure 5.35 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream S10 inlet for System 1A

Figure 5.36 illustrates the variation of heat load, exergy flow and temperature 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. 84

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.

140 0 -20 120 -40 100 -60 80 Total Exergy -80 Heat Load -100 60 Temperature -120

40 Temperature (˚C) -140 20 -160 Heat Load and Exergy Flow (kW) 0 -180 401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1 Specific Exergy Flow (kJ/kg)

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

85 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.

140 -182

120 -184 -186 100 -188 80 Total Exergy -190 60 Heat Load -192 Temperature 40 -194

-196 Temperature (˚C) 20 -198

Heat Load and Exergy Flow (kW) 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 -20 -202 Specific Exergy Flow (kJ/kg)

Figure 5.37 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 22 inlet for System 1A

70 -205

60 -210 50 Total Exergy -215 40 Heat Load Temperature 30 -220

20 -225 Temperature (˚C) 10 -230

Heat Load and Exergy Flow (kW) 0 401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1 -10 -235 Specific Exergy Flow (kJ/kg)

Figure 5.38 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.

200 -244.98

0 -244.99 -200 -245 -400 Total Exergy -600 -245.01 Heat Load Temperature -800

-245.02 Temperature (˚C) -1000

Heat Load and Exergy Flow (kW) -245.03 -1200

-1400 -245.04 401.4 431.1 462.2 494.8 528.9 564.6 601.9 641.0 681.9 724.6 769.4 816.1 Specific Exergy Flow (kJ/kg)

Figure 5.39 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream S1 inlet for System 1A

12 0

Total Exergy -10 10 Heat Load Temperature -20 8

-30

-40

4 Temperature (˚C) -50

Heat Load and Exergy Flow (kW) -60

0 -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 Specific Exergy Flow (kJ/kg)

Figure 5.40 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 17 inlet for System 1A

140 -198.5

-199 120 -199.5 100 Total Exergy -200 80 Heat Load -200.5 Temperature 60 -201 -201.5 40 Temperature (˚C) -202 20 Heat Load and Exergy Flow (kW) -202.5

0 -203 396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2 Specific Exergy Flow (kJ/kg)

Figure 5.41 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream S10 inlet for System 1B

120 0

-20 100 Total Exergy Heat Load -40 80 Temperature -60

-80 60 -100

40 -120 Temperature (˚C) -140 20

Heat Load and Exergy Flow (kW) -160

0 -180 396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2 Specific Exergy Flow (kJ/kg)

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. 88

300 -122

-123 250

-124 Total Exergy 200 Heat Load Temperature -125

150 -126

100 -127 Temperature (˚C) -128 50 Heat Load and Exergy Flow (kW) -129

0 396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2 -130

-50 -131 Specific Exergy Flow (kJ/kg)

Figure 5.43 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 13 inlet for System 1B

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.

160 0 140 -20 120 -40 100 -60 80 -80 60 Total Exergy -100 40 Heat Load -120 Temperature 20 -140 Temperature (˚C) 0 -160 -20 -180 Heat Load and Exergy Flow (kW) -40 -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 Specific Exergy Flow (kJ/kg)

Figure 5.44 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 49 inlet for System 1B

300 0 250 -50 200

150 Total Exergy -100 100 Heat Load Temperature 50 -150

0 Temperature (˚C) -200 -50

Heat Load and Exergy Flow (kW) -100 -250 396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2 Specific Exergy Flow (kJ/kg)

Figure 5.45 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 22 inlet for System 1B

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 0 Total Exergy Heat Load 150 Temperature -50

100 -100

-150

0 Temperature (˚C)

Heat Load and Exergy Flow (kW) -200 -50

-100 -250 396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2 Specific Exergy Flow (kJ/kg)

Figure 5.46 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 17 inlet for System 1B

200 -244.994

0 -244.996

-200 Total Exergy -244.998 Heat Load -400 Temperature -245

-600 -245.002

-800 -245.004

-1000 -245.006 Temperature (˚C)

Heat Load and Exergy Flow (kW) -1200 -245.008

-1400 -245.01

-1600 -245.012 396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2 Specific Exergy Flow (kJ/kg)

Figure 5.47 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream S1 inlet for System 1B

70 -205

60 -210 50 Total Exergy -215 40 Heat Load Temperature 30 -220

20 -225 Temperature (˚C) 10 -230

Heat Load and Exergy Flow (kW) 0

-10 -235 396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2 Specific Exergy Flow (kJ/kg)

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 overall system exergy efficiency. Different catalyst types and performance enhancement of the heat exchangers could create a dramatic improvement to the process efficiency.

100 92 99 92 90 90 77 77 76 77 80 75 70 57 61 60 44 47 46 44 50 38 40 37 40 32 33 35 30 Efficiency (%) 13 Exergy Efficiency 20 810 11 9 12 912 10 Energy Efficiency 0

CoolerCooler EX1 Cooler EX2 Cooler EX3 Cooler EX4 Cooler EX5 EX6

CompressorCompressor (C1)Compressor (C2)Compressor (C3)Compressor (C4)Flash (C5)Drum (F1)

Heat ExchangerHeat ExchangerHeat (HX1) Exchanger (HX2) (HX3) (a) 92 91 92 94 100 86 90 77 76 77 78 78 80 70 60 48 50 42 39 42 40 40 30 32 30 32 30 25 23

Efficiency (%) 18 20 11 Exergy Efficiency 78 56 9 10 Energy Efficiency 0

CoolerCooler EX1 Cooler EX2 Cooler EX3 Cooler EX4 Cooler EX5 EX6

CompressorCompressor Compressor(C1) Compressor(C2) Compressor(C3) (C4)Flash (C5)Drum (F1)

Heat ExchangerHeat ExchangerHeat (HX1) Exchanger (HX2) (HX3) (b) 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%.

60.0% Exergy Efficiency

50.0% Energy Efficiency

40.0%

30.0%

Percentage 20.0%

10.0%

0.0% 1.0 3.9 6.7 9.6 12.4 15.3 18.1

H2 Feed Pressure (bar)

(a)

15.50% Exergy Efficiency Energy Efficiency 15.00%

14.50%

14.00% Percentage

13.50%

13.00% 1.0 3.9 6.7 9.6 12.4 15.3 18.1 H2 Feed Pressure (bar)

(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.

80.00%

70.00%

60.00%

50.00%

40.00% Exergy Efficiency

Efficiency 30.00% Energy Efficiency 20.00%

10.00%

0.00% 1.50 2.00 2.50 3.00 3.50 4.00 4.50 Compressor 1 (C1) pressure (bar)

(a)

70.00%

60.00%

50.00%

40.00%

30.00% Exergy Efficiency Efficiency Energy Efficiency 20.00%

10.00%

0.00% 1.50 2.00 2.50 3.00 3.50 4.00 4.50 Compressor 1 (C1) pressure (bar)

(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.

70.0% Exergy Efficiency

60.0% Energy Efficiency

50.0%

40.0%

30.0% Efficiency

20.0%

10.0%

0.0% 2.00 2.22 2.44 2.67 2.89 3.11 3.33 3.56 3.78 4.00 20.00 Compressor 5 (C5) pressure (bar)

(a)

25.00%

20.00%

15.00% Exergy Efficiency Energy Efficiency 10.00% Efficiency

5.00%

0.00% 2.00 2.22 2.44 2.67 2.89 3.11 3.33 3.56 3.78 4.00 20.00 Compressor 5 (C5) pressure (bar)

(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

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 5.53 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. 96

Figure 5.55 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 as can be seen in Figure 5.54 and Figure 5.55.

80000 -195.775

70000 TOTAL EXERGYFL -195.78 60000 Heat duty

50000 Temperature -195.785

40000 -195.79 30000

20000 -195.795 Temperature (˚C) 10000 -195.8 0

Heat Duty and Exergy Flow (kW) 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 -10000 -195.805 Spesific Exergy Flow (kJ/kg)

Figure 5.53 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream N2LIQ inlet for System 2A

120 0

100 -20

80 TOTAL EXERGYFL -40 60 Heat duty -60 40 -80 20 -100 0

134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Temperature (˚C) -120 -20

-40 -140 Heat Duty and Exergy Flow (kW) -60 -160 Spesific Exergy Flow (kJ/kg)

Figure 5.54 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 28 inlet for System 2A

At inlet 36 and 44 the exergy flow drops at an almost constant rate as shown in Figure 5.56 and Figure 5.57. It can be seen that in both cases, as the specific exergy flow rises, heat load increases and exergy decreases. Additionally, higher specific energy flows correspond to higher inlet temperatures in both cases. 97

1000 0

-20 800

-40 600 TOTAL EXERGYFL -60 Heat duty 400 Temperature -80

200 -100

-120 0

-140 Temperature (˚C) -200 -160

Heat Duty and Exergy Flow (kW) -400 -180

-600 -200 891.4 901.9 915.9 934.0 956.8 984.9 1019.3 1061.1 1111.6 1172.8 1247.4 1339.3 Spesific Exergy Flow (kJ/kg)

Figure 5.55 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 9 inlet for System 2A

180 0 TOTAL EXERGYFL 160 Heat duty Temperature -50 140

120 -100 100

80 -150 60 Temperature (˚C)

40 -200

20 Heat Duty and Exergy Flow (kW)

0 -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 Spesific Exergy Flow (kJ/kg)

Figure 5.56 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 36 inlet for System 2A

10000 0 TOTAL EXERGYFL 9000 Heat duty Temperature 8000 -50

7000

6000 -100

5000

4000 -150

3000 Temperature (˚C)

2000 -200

1000 Heat Duty and Exergy Flow (kW)

0 -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 Spesific Exergy Flow (kJ/kg)

Figure 5.57 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 44 inlet for System 2A

80000 -1.96E+02 TOTAL EXERGYFL 70000 Heat duty Temperature -1.96E+02 60000

50000 -1.96E+02

40000 -1.96E+02 30000

20000 -1.96E+02 Temperature (˚C) 10000 -1.96E+02 0 Heat Duty and Exergy Flow (kW)

-10000 -1.96E+02 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Spesific Exergy Flow (kJ/kg)

Figure 5.58 Heat Load, Exergy flow vs Temperature for Precooling Phase heat exchanger HX1 at stream N2LIQ inlet for System 2B

Figure 5.59 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.

60 0.00E+00 TOTAL EXERGYFL 50 Heat duty -2.00E+01 Temperature 40 -4.00E+01

30 -6.00E+01 20 -8.00E+01 10 -1.00E+02 0 Temperature (˚C) -1.20E+02 -10

-20 -1.40E+02 Heat Duty and Exergy Flow (kW)

-30 -1.60E+02 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Spesific Exergy Flow (kJ/kg)

Figure 5.59 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 28 inlet for System 2B

700 0

600 -20

500 -40 TOTAL EXERGYFL 400 -60 Heat duty 300 Vapor fraction -80

200 -100

100 -120

0 -140 Temperature (˚C)

-100 -160

Heat Duty and Exergy Flow (kW) -200 -180

-300 -200 3732.7 3775.8 3833.7 3908.4 4002.1 4117.8 4259.0 4430.3 4637.5 4888.7 5195.0 5572.4 Spesific Exergy Flow (kJ/kg)

Figure 5.60 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 9 inlet for System 2B

Figure 5.60 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. 100

Figure 5.61 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 case, it can be noted that the relationship between specific exergy flow and all three other variables is linear.

250 0.00E+00 TOTAL EXERGYFL Heat duty Temperature 200 -5.00E+01

150 -1.00E+02

100 -1.50E+02 Temperature (˚C)

50 -2.00E+02 Heat Duty and Exergy Flow (kW)

0 -2.50E+02 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Spesific Exergy Flow (kJ/kg)

Figure 5.61 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 36 inlet for System 2B

Figure 5.62 shows the variation of temperature, exergy flow and heat load for precooling heat exchanger HX1 at stream 44 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.

5.3.2 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 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 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. 101

180 0.00E+00 TOTAL EXERGYFL 160 Heat duty Temperature -5.00E+01 140

120 -1.00E+02 100

80 -1.50E+02 60 Temperature (˚C) 40 -2.00E+02

20 Heat Duty and Exergy Flow (kW)

0 -2.50E+02 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Spesific Exergy Flow (kJ/kg)

Figure 5.62 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 44 inlet for System 2B

1000 50 TOTAL EXERGYFL Heat duty 800 Temperature 0

600

-50 400

200 -100

0 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 -150 Temperature (˚C) -200

-200

Heat Duty and Exergy Flow (kW) -400

-600 -250 Spesific Exergy Flow (kJ/kg)

Figure 5.63 Heat Load, Exergy flow and Temperature for liquefaction Phase heat exchanger HX2 at stream 49 inlet for System 2A

Figure 5.64 illustrates the variation of heat load, exergy flow and temperature 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 102 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.

1000 0 TOTAL EXERGYFL Heat duty 800 Temperature -50

600 -100

400

-150 200 Temperature (˚C)

-200 0

Heat Duty and Exergy Flow (kW) 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5

-200 -250 Spesific Exergy Flow (kJ/kg)

Figure 5.64 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 13 inlet for System 2A

400 0 TOTAL EXERGYFL Heat duty 200 Temperature -50

0 -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

-200 -150

-400 -200 Temperature (˚C)

-600 -250 Heat Duty and Exergy Flow (kW)

-800 -300 Spesific Exergy Flow (kJ/kg)

Figure 5.65 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 26 inlet for System 2A

103

Figure 5.65 illustrates the variation of heat load, exergy flow and temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 2A. It can be 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.

The variation brought about by pre cooling in heat exchanges HX2 can be seen in Figure 5.67, where the heat load, exergy flow and temperature for Liquefaction Phase heat exchanger HX2 at stream 22 inlet for System 2A 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.

600 800 TOTAL EXERGYFL Heat duty 700 500 Temperature 600

500 400 400

300 300 200

100 200

0 Temperature (˚C)

100 -100

Heat Duty and Exergy Flow (kW) -200

0 -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 Spesific Exergy Flow (kJ/kg)

Figure 5.66 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 2A

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 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.

104

With variations at S1, it is seen in Figure 5.69 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.

200 -2.24E+02 TOTAL EXERGYFL 100 Heat duty -2.26E+02 Temperature 0 -2.28E+02

-100 -2.30E+02 -200 -2.32E+02 -300 -2.34E+02 -400 -2.36E+02 -500 Temperature (˚C) -2.38E+02 -600

-700 -2.40E+02 Heat Duty and Exergy Flow (kW)

-800 -2.42E+02 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Spesific Exergy Flow (kJ/kg)

Figure 5.67 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 22 inlet for System 2A

400 -205 TOTAL EXERGYFL 350 Heat duty Temperature -210 300

250 -215

200 -220 150

100 -225 Temperature (˚C) 50 -230 0 Heat Duty and Exergy Flow (kW)

-50 -235 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Spesific Exergy Flow (kJ/kg)

Figure 5.68 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 32b inlet for System 2A

105

120000 0.00E+00 TOTAL EXERGYFL 100000 Heat duty -5.00E+01 Temperature

80000 -1.00E+02

60000 -1.50E+02 40000

-2.00E+02

20000 Temperature (˚C)

-2.50E+02 0 Heat Duty and Exergy Flow (kW)

-20000 -3.00E+02 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Spesific Exergy Flow (kJ/kg)

Figure 5.69 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream S1 inlet for System 2A

Figure 5.70 shows the variation of heat load, exergy and temperature for HX3 at stream inlet 17 for system 2A. It can be seen that heat load increases with increase in specific exergy flow while total exergy flow decreases. At lower temperatures, the specific exergy flow is higher and vice versa. The next set of graphs illustrate the variation of exergy, heat load and temperature for system 2A.

400 8.00E+02 TOTAL EXERGYFL 300 Heat duty 7.00E+02 Temperature 200 6.00E+02

100 5.00E+02 0 4.00E+02 -100 3.00E+02 -200 Temperature (˚C) 2.00E+02 -300

-400 1.00E+02 Heat Duty and Exergy Flow (kW)

-500 0.00E+00 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Spesific Exergy Flow (kJ/kg)

Figure 5.70 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 17 inlet for System 2A

106

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.

700 0.00E+00 TOTAL EXERGYFL 600 Heat duty Temperature 500 -5.00E+01

400

300 -1.00E+02

200

100 -1.50E+02

0 Temperature (˚C)

-100 -2.00E+02

-200 Heat Duty and Exergy Flow (kW)

-300 -2.50E+02 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Spesific Exergy Flow (kJ/kg)

Figure 5.71 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 49 inlet for System 2B

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.

107

140 -198.5

-199 120 -199.5 100 Total Exergy -200 80 Heat Load -200.5 Temperature 60 -201 -201.5 40 Temperature (˚C) -202 20 Heat Load and Exergy Flow (kW) -202.5

0 -203 396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2 Specific Exergy Flow (kJ/kg)

Figure 5.72 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 26 inlet for System 2B

120 0

-20 100 Total Exergy Heat Load -40 80 Temperature -60

-80 60 -100

40 -120 Temperature (˚C) -140 20

Heat Load and Exergy Flow (kW) -160

0 -180 396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2 Specific Exergy Flow (kJ/kg)

Figure 5.73 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 2B

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.

108

300 -122

-123 250 -124 200 Total Exergy -125 150 Heat Load -126 Temperature 100 -127

-128 Temperature (˚C) 50 -129

Heat Load and Exergy Flow (kW) 0 -130 396.7 427.2 459.1 492.6 527.7 564.5 603.1 643.5 685.9 730.2 776.6 825.2 -50 -131 Specific Exergy Flow (kJ/kg)

Figure 5.74 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 13 inlet for System 2B

100 -2.24E+02 TOTAL EXERGYFL 50 Heat duty -2.26E+02 Temperature 0 -2.28E+02

-50 -2.30E+02

-100 -2.32E+02

-150 -2.34E+02

-200 -2.36E+02 Temperature (˚C) -250 -2.38E+02

-300 -2.40E+02 Heat Duty and Exergy Flow (kW)

-350 -2.42E+02 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Spesific Exergy Flow (kJ/kg)

Figure 5.75 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 22 inlet for System 2B

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

109 trend changes, and the total; exergy flow decreases in a nonlinear manner with increase in specific exergy flow.

400 0.00E+00 TOTAL EXERGYFL 350 Heat duty Temperature -5.00E+01 300

250 -1.00E+02

200

-1.50E+02 150 Temperature (˚C) 100 -2.00E+02 50 Heat Duty and Exergy Flow (kW)

0 -2.50E+02 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Spesific Exergy Flow (kJ/kg)

Figure 5.76 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 17 inlet for System 2B

1000 0.00E+00 TOTAL EXERGYFL Heat duty Temperature -5.00E+01 500

-1.00E+02 0

-1.50E+02

-500 -2.00E+02 Temperature (˚C)

-1000 -2.50E+02 Heat Duty and Exergy Flow (kW)

-1500 -3.00E+02 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Spesific Exergy Flow (kJ/kg)

Figure 5.77 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream S1 inlet for System 2B

Figure 5.77 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

110 nonlinear relationship, with lower values of specific exergy flow corresponding to 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.

300 -2.05E+02 TOTAL EXERGYFL Heat duty 250 Temperature -2.10E+02

200 -2.15E+02

150 -2.20E+02 100

-2.25E+02

50 Temperature (˚C)

-2.30E+02 0 Heat Duty and Exergy Flow (kW)

-50 -2.35E+02 134.9 144.4 154.3 164.6 175.5 186.9 198.8 211.3 224.4 238.1 252.5 267.5 Spesific Exergy Flow (kJ/kg)

Figure 5.78 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat 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

111 enhancement of the heat exchangers could create a dramatic improvement to the process efficiency.

100 97 95 95 93 84 85 90 81 77 80 78 78 80 7… 70 60 50 36 36 40 35 30 2323 Efficiency (%) 20 78 99 7 Exergy Efficiency 10 2.23 45 0 Energy Efficiency

CompressorCompressor (C1)Compressor (C2)Flash (C3) Drum (F1) Vortex TubeVortex (VT1) Tube (VT2) Heat ExchangerHeat ExchangerHeat (HX1) Exchanger Expansion(HX2) Expansion(HX3) ValveExpansion (V1) Valve (V2) Valve (V3) Adsorption Unit (AO)

93 96 94 100 99 96 100 85 86 90 99 90 81 77 80 70 60 44 50 3539 40 31 34 2324 30 1619

Efficiency (%) 15 20 99 Exergy Efficiency 10 33 0 Energy Efficiency

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

5.4.1 Pre-cooling phase at systems S3A and S3B

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. shows the Heat Load, Exergy flow and Temperature against specific 112 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 5.80 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 as can be seen in Figure 5.81 and Figure 5.82

1200 0 TOTAL EXERGYFL Heat duty 1000 Temperature -50

800 -100

600

-150 400 Temperature (˚C)

-200 200 Heat Duty and Exergy Flow (kW)

0 -250 760.2 760.1 687.9 615.6 543.4 471.1 398.8 326.6 254.3 190.4 182.2 125.0 84.8 56.1 Spesific Exergy Flow (kJ/kg)

Figure 5.80 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream N2LIQ inlet for System 3A

At inlet 36 and 44 the exergy flow drops at an almost constant rate as shown in Figure 5.83 and Figure 5.84. It can be seen that in both cases, as the specific exergy flow rises, heat load increases and exergy decreases. Additionally, higher specific energy flows correspond to higher inlet temperatures in both cases

113

60 -100

50 -105

40 TOTAL EXERGYFL Heat duty -110 30 Temperature

20 -115

10 -120 Temperature (˚C) 0

-125 -10 Heat Duty and Exergy Flow (kW)

-20 -130 643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8 Spesific Exergy Flow (kJ/kg)

Figure 5.81 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 28 inlet for System 3A

60 -100

50 -105

40 TOTAL EXERGYFL Heat duty -110 30 Temperature

20 -115

10 -120 Temperature (˚C) 0

-125 -10 Heat Duty and Exergy Flow (kW)

-20 -130 643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8 Spesific Exergy Flow (kJ/kg)

Figure 5.82 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 9 inlet for System 3A

114

50 -140 TOTAL EXERGYFL Heat duty 40 Temperature -142

30 -144

20 -146

10 -148 Temperature (˚C)

0 -150 Heat Duty and Exergy Flow (kW)

-10 -152 643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8 Spesific Exergy Flow (kJ/kg)

Figure 5.83 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 36 inlet for System 3A

35 -140 TOTAL EXERGYFL 30 Heat duty Temperature -142

-144 20

15 -146

10 -148 Temperature (˚C) 5

-150 0 Heat Duty and Exergy Flow (kW)

-5 -152 643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8 Spesific Exergy Flow (kJ/kg)

Figure 5.84 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 44 inlet for System 3A

115

1200 -195.775 TOTAL EXERGYFL Heat duty 1000 Temperature -195.78

800 -195.785

600 -195.79

400 -195.795 Temperature (˚C)

200 -195.8 Heat Duty and Exergy Flow (kW)

0 -195.805 760.2 760.1 711.5 662.8 614.1 565.5 516.8 468.1 419.5 370.8 322.1 273.4 224.8 Spesific Exergy Flow (kJ/kg)

Figure 5.85 Heat Load, Exergy flow vs Temperature for Precooling Phase heat exchanger HX1 at stream N2LIQ inlet for System 3B

1200 -195.775

1000 -195.78

TOTAL EXERGYFL 800 Heat duty -195.785 Temperature

600 -195.79

400 -195.795 Temperature (˚C)

200 -195.8 Heat Duty and Exergy Flow (kW)

0 -195.805 760.2 760.1 711.5 662.8 614.1 565.5 516.8 468.1 419.5 370.8 322.1 273.4 Spesific Exergy Flow (kJ/kg)

Figure 5.86 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 28 inlet for System 3B

Figure 5.85 shows the changes in heat load, Exergy flow and Temperature for 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. 116

1200 -195.775

1000 -195.78

TOTAL EXERGYFL 800 Heat duty -195.785 Temperature

600 -195.79

400 -195.795 Temperature (˚C)

200 -195.8 Heat Duty and Exergy Flow (kW)

0 -195.805 760.2 760.1 711.5 662.8 614.1 565.5 516.8 468.1 419.5 370.8 322.1 273.4 Spesific Exergy Flow (kJ/kg)

Figure 5.87 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 9 inlet for System 3B

1200 -195.775 TOTAL EXERGYFL Heat duty 1000 Temperature -195.78

800 -195.785

600 -195.79

400 -195.795 Temperature (˚C)

200 -195.8 Heat Duty and Exergy Flow (kW)

0 -195.805 760.2 760.1 711.5 662.8 614.1 565.5 516.8 468.1 419.5 370.8 322.1 273.4 Spesific Exergy Flow (kJ/kg)

Figure 5.88 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 36 inlet for System 3B

Figure 5.85 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

117 specific exergy flow increases, the total exergy flow also increases, but the two have a non-linear relationship.

Figure 5.88 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 case, it can be noted that the relationship between specific exergy flow and all three other variables is linear.

40 -135 TOTAL EXERGYFL 35 Heat duty -140 Temperature -145 30

-150 25 -155 20 -160 15 -165 Temperature (˚C) 10 -170

5 -175 Heat Duty and Exergy Flow (kW)

0 -180 760.2 760.1 711.5 662.8 614.1 565.5 516.8 468.1 419.5 370.8 322.1 273.4 Spesific Exergy Flow (kJ/kg)

Figure 5.89 Heat Load, Exergy flow and Temperature for Precooling Phase heat exchanger HX1 at stream 44 inlet for System 3B

Figure 5.89 shows the variation of temperature, exergy flow and heat load for precooling heat exchanger HX1 at stream 44 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.

5.4.2 Liquefaction Phase at systems S3A and S3B

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 exergy flow while heat load decreases. The variation of heat load is linear, while total 118 exergy flow has a nonlinear relationship. Higher specific exergy flows are achieved at lower temperatures.

140 0

120 -20

100 -40

80 TOTAL EXERGYFL -60 60 Heat duty Temperature -80 40 -100 20 Temperature (˚C) -120 0

-20 -140 Heat Duty and Exergy Flow (kW)

-40 -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 Spesific Exergy Flow (kJ/kg)

Figure 5.90 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 49 inlet for System 3A

350 -118 TOTAL EXERGYFL 300 Heat duty Temperature -120

250 -122

200 -124 150 -126 100

-128 Temperature (˚C) 50

0 -130 Heat Duty and Exergy Flow (kW)

-50 -132 643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8 Spesific Exergy Flow (kJ/kg)

Figure 5.91 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 13 inlet for System 3A

Figure 5.91 illustrates the variation of heat load, exergy flow and temperature 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 119 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.

160 -198.5 TOTAL EXERGYFL 140 Heat duty -199 Temperature -199.5 120

-200 100 -200.5 80 -201 60 -201.5 Temperature (˚C) 40 -202

20 -202.5 Heat Duty and Exergy Flow (kW)

0 -203 643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8 Spesific Exergy Flow (kJ/kg)

Figure 5.92 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream S10 inlet for System 3A

Figure 5.93 illustrates the variation of heat load, exergy flow and temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 2A. It can be 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.

The variation brought about by pre cooling in heat exchanges HX2 can be seen in Figure 5.94 where the heat load, exergy flow and temperature for Liquefaction Phase heat exchanger HX2 at stream 22 inlet for System 3A 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.

120

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 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 5.96 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.

120 0 TOTAL EXERGYFL Heat duty -20 100 Temperature -40

80 -60

-80 60 -100

40 -120 Temperature (˚C)

-140 20

Heat Duty and Exergy Flow (kW) -160

0 -180 643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8 Spesific Exergy Flow (kJ/kg)

Figure 5.93 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 3A

Figure 5.97 shows the variation of heat load, exergy and temperature for HX3 at stream inlet 17 for system 3A. It can be seen that heat load increases with increase in specific exergy flow while total exergy flow decreases. At lower temperatures, the specific exergy flow is higher and vice versa. The next set of graphs illustrate the variation of exergy, heat load and temperature for system 3A.

121

140 0

120 -20

100 TOTAL EXERGYFL -40 Heat duty 80 Temperature -60

60 -80 40

-100 20 Temperature (˚C)

-120 0 Heat Duty and Exergy Flow (kW) -20 -140

-40 -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 Spesific Exergy Flow (kJ/kg)

Figure 5.94 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 22 inlet for System 3A

80 -205

70 -210 60

50 TOTAL EXERGYFL -215 Heat duty

40 Temperature -220 30

20 -225 Temperature (˚C)

10 -230 Heat Duty and Exergy Flow (kW) 0

-10 -235 643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8 Spesific Exergy Flow (kJ/kg)

Figure 5.95 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 32b inlet for System 3A

122

140 -244.955

-244.96 120

-244.965 100 TOTAL EXERGYFL Heat duty -244.97

80 Temperature -244.975

60 -244.98

-244.985 40 Temperature (˚C) -244.99 20 -244.995 Heat Duty and Exergy Flow (kW)

0 -245

-20 -245.005 643.7 659.6 675.7 692.2 709.0 726.0 743.4 761.0 779.0 797.3 815.9 834.8 Spesific Exergy Flow (kJ/kg)

Figure 5.96 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream S1 inlet for System 3A

140 0

120 -20

100 -40 TOTAL EXERGYFL 80 Heat duty -60 60 Temperature -80 40 -100 20 Temperature (˚C) -120 0

-20 -140 Heat Duty and Exergy Flow (kW)

-40 -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 Spesific Exergy Flow (kJ/kg)

Figure 5.97 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 17 inlet for System 3A

123

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.

160 0

140 -20

120 -40

100 -60

80 TOTAL EXERGYFL -80 Heat duty 60 Temperature -100 40 -120

20 -140 Temperature (˚C)

0 -160

-20 -180 Heat Duty and Exergy Flow (kW)

-40 -200 3381.5 3457.7 3541.2 3632.7 3732.8 3842.5 3962.6 4094.3 4238.8 4397.7 4572.6 4765.7 Spesific Exergy Flow (kJ/kg)

Figure 5.98 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 49 inlet for System 3B

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.

124

140 -198.5

-199 120

-199.5 TOTAL EXERGYFL 100 Heat duty -200 Temperature

80 -200.5

-201 60 Temperature (˚C)

-201.5

40 Heat Duty and Exergy Flow (kW)

-202

20 -202.5

0 -203 2858.3 2848.0 2837.7 2827.4 2817.2 2807.1 2797.0 2786.9 2776.9 2766.9 2756.9 2747.0 Spesific Exergy Flow (kJ/kg)

Figure 5.99 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream S10 inlet for System 3B

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.

125

140 -198.5

-199 120

-199.5 TOTAL EXERGYFL 100 Heat duty -200 Temperature

80 -200.5

60 -201

-201.5

40 Temperature (˚C) -202

Heat Duty and Exergy Flow (kW) -202.5

0 -203 2858.3 2848.0 2837.7 2827.4 2817.2 2807.1 2797.0 2786.9 2776.9 2766.9 2756.9 2747.0 Spesific Exergy Flow (kJ/kg)

Figure 5.100 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 14 inlet for System 3B

300 -122

-123 250

-124 TOTAL EXERGYFL 200 Heat duty -125 Temperature 150 -126

100 -127

-128

50 Temperature (˚C) -129

Heat Duty and Exergy Flow (kW) -130

-50 -131 3661.8 3668.5 3675.2 3681.9 3688.6 3695.4 3702.3 3709.2 3716.1 3723.1 3730.1 3737.2 Spesific Exergy Flow (kJ/kg)

Figure 5.101 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX2 at stream 13 inlet for System 3B

Figure 5.104 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.

126

300 0

250

-50 200

150 TOTAL EXERGYFL -100 Heat duty 100 Temperature

-150 50 Temperature (˚C) 0 -200

-50 Heat Duty and Exergy Flow (kW)

-100 -250 3737.2 3842.2 3956.8 4081.9 4218.6 4368.2 4532.1 4712.3 4910.7 5129.8 5372.8 5643.2 Spesific Exergy Flow (kJ/kg)

Figure 5.102 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 22 inlet for System 3B

200 0

-50 -200 TOTAL EXERGYFL -400 Heat duty Temperature -100 -600

-800 -150 -1000 Temperature (˚C) -1200 -200

-1400 Heat Duty and Exergy Flow (kW)

-1600 -250

-13880.1 -13876.6 -13873.1 -13869.5 -13866.0 -13862.5 -13859.0 -13855.5 -13852.0 -13848.5 -13845.0 -13841.5 Spesific Exergy Flow (kJ/kg)

Figure 5.103 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 17 inlet for System 3B 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.

127

200 -244.994

0 -244.996

-200 TOTAL EXERGYFL -244.998 Heat duty -400 Temperature -245

-600 -245.002

-800 -245.004

-1000 -245.006 Temperature (˚C) -1200 -245.008

-1400 -245.01 Heat Duty and Exergy Flow (kW)

-1600 -245.012

-13880.1 -13876.6 -13873.1 -13869.5 -13866.0 -13862.5 -13859.0 -13855.5 -13852.0 -13848.5 -13845.0 -13841.5 Spesific Exergy Flow (kJ/kg)

Figure 5.104 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream S1 inlet for System 3B

300 0

250

-50

200

150 TOTAL EXERGYFL -100 Heat duty 100 Temperature

-150 50 Temperature (˚C)

0 Heat Duty and Exergy Flow (kW) -200

-50

-100 -250 3737.2 3842.2 3956.8 4081.9 4218.6 4368.2 4532.1 4712.3 4910.7 5129.8 5372.8 5643.2 Spesific Exergy Flow (kJ/kg)

Figure 5.105 Heat Load, Exergy flow and Temperature for Liquefaction Phase heat exchanger HX3 at stream 32b inlet for System 3B

128

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

-245.0 x 1.0133 bar y 1.0133 bar -245.5

-246.0

-246.5

-247.0

-247.5

-248.0

-248.5

-249.0 Temperature, C Temperature, -249.5

-250.0

-250.5

-251.0

-251.5

-252.0

-252.5

-253.0 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 Liquid/vapor mole fraction, HYDRO-01

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.

129

T-x diagram for HYDRO-01/HYDRO-02 -245.0

1.0133 bar -245.5

-246.0

-246.5

-247.0

-247.5

-248.0

-248.5

-249.0

Temperature, C Temperature, -249.5

-250.0

-250.5

-251.0

-251.5

-252.0

-252.5

K-values for HYDRO-01/HYDRO-02 8.0

Liquid1 HYDRO-01 1.0133 bar 7.5 Liquid1 HYDRO-02 1.0133 bar

7.0

6.5

6.0

5.5

5.0

4.5

4.0 K-values

3.5

3.0

2.5

2.0

1.5

1.0

0.5

0.0 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 Liquid/vapor mole fraction, HYDRO-01 Figure 5.108 K-values for Vapor-liquid vs fraction of Para-hydrogen and Ortho Hydrogen

130

y-x diagram for HYDRO-01/HYDRO-02 1.00

0.95 1.0133 bar

0.90

0.85

0.80

0.75

0.70

0.65

0.60

0.55

0.50

0.45

0.40 Vapor Vapor molefraction, HYDRO-01

0.35

0.30

0.25

0.20

0.15

0.10

0.05

0.00 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 1.00 Liquid/vapor mole fraction, HYDRO-01

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

Activity coefficients for HYDRO-01/HYDRO-02 2.40

2.35 Liquid1 HYDRO-01 1.0133 bar 2.30 Liquid1 HYDRO-02 1.0133 bar

2.25

2.20

2.15

2.10

2.05

2.00

1.95

1.90

1.85

1.80

1.75

1.70

1.65

Activitycoefficients 1.60

1.55

1.50

1.45

1.40

1.35

1.30

1.25

1.20

1.15

1.10

1.05

1.00 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 Liquid/vapor mole fraction, HYDRO-01 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.

131

50 42.0 40

30 24.0 22.4 21.0 19.3 18.9 17.7 17.5 17.4 17.2 16.8 16.5 16.2

20 15.4 13.4 12.0 12.0 11.9 11.9 10.6 10.4 9.4 8.0 6.7 6.0

10 5.0 4.0 Exergy EfficiencyExergy (%) 2.3 0 Main S1A S1B S2A S2B S3A S3B

Exergy effeciency (0°C) Exergy effeciency (10°C) Exergy effeciency (25°C) Exergy effeciency (45°C)

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.

80 76 74 70

60 54 50

40 32 30

Energy EfficiencyEnergy (%) 20 10 12 10 7

0 Main S1-A S1-B S2-A S2-B S3-A S3-B

Figure 5.112 Energy efficiency for the hydrogen liquefaction systems.

132

20000.0

18000.0 17216.1 15941.2 16000.0 15017.9

14000.0 12688.1 12000.0

10000.0

Work (kw) 8000.0

6000.0

4000.0 3403.9 1657.8 2000.0

0.0 S1A S1B S2A S2B S3A S3B

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 �̇ + �

max �� subject to �̇ = �� = 3.628 ��/ℎ�

� ≤ 0

Here, �̇ is the sum of all compressor powers (kW). � ≤ 0 represents 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

o Feed hydrogen gas stream: normal hydrogen gas enters with P = 1 bar and T = 25 C after gas. 133

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, �0 -10 50 °C Compressor Pressure -10 100 °C LN Pressure 400 1500 Bar Heat exchanger minimum approach 5 10 °C temperature

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 134

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).

30% 27

25% 22

20% 19 17 17 16 15%

Efficiency 10 10%

0% S1A S1B S2A S2B S3A S3B OPTIMUM

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.

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Figure 5.115 Optimized system

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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

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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 systems 2A and 2B are found to be 22.55 % and 10.44 % respectively.

• The work done per liquefaction mass of systems 2A and 2B are found to be 12688.07 kJ/kg and 3403.89 kJ/kg respectively.

• The overall energy and exergy efficiencies of systems 3A and 3B are found to be 17.55 % and 16.44 % respectively.

• The work done per liquefaction mass of systems 3A and 3B are found to be 1657.76 kJ/kg and 17216.12 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.

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• 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.

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