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DISSERTATION Thermodynamic and Economic Evaluation of Existing and Prospective Processes for Liquefaction of Natural Gas in Malaysia M.Sc. Mohd Nazri Bin Omar Berlin
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DISSERTATION

ThermodynamicandEconomicEvaluationofExistingandProspectiveProcessesfor

LiquefactionofNaturalGasinMalaysia

M.Sc.MohdNazriBinOmar

Berlin

ThermodynamicandEconomic

EvaluationonExistingandPerspectiveProcessesforLiquefactionofNaturalGasinMalaysia

 

vorgelegt von 

Master of Science in Global Production Engineering 

Mohd Nazri Bin Omar geb. in Johor Bahru, 

Malaysia 

 

von der Fakultät III – Prozesswissenschaften 

der Technischen Universität Berlin 

zur Erlangung des akademischen Grades 

Doktor der Ingenieurwissenschaften 

– Dr.‐Ing. – 

 

genehmigte Dissertation 

 

Promotionsausschuss 

Vorsitzender: Prof. Dr.‐Ing. Felix Ziegler 

Berichter: Prof. Dr. Tetyana Morozyuk 

Berichter: Prof. Dr. Wojciech Stanek 

Berichter: Prof. Dr.‐Ing. George Tsatsaronis 

Tag der wissenschaftlichen Aussprache: 22. Oktober 2015 

 

Berlin, 2016

1

Foreword

ThisworkhasbeenconductedduringmyleavefromUniversitiMalaysiaPerlis UniMAP asa On‐Study Leave staff of School of Manufacturing Engineering, and hence as a Ph.Dstudent at the )nstitute for Energy Engineering at the Technische Universität Berlin TUBerlin .ThisstudywassponsoredbyUniMAPandbytheMinistryofEducationMalaysiainrefertoscholarship)DKPT BS SkimLatihanAkademikBumiputera SLAB .With this opportunity, ) would like to express my utmost praise and gratitude to theAlmightyAllah,andtothepeoplethatfacilitatedthecompletionofthiswork.)amespeciallygratefultoProfessorTetyanaMorozyuk,whosupervisedthisworkandwasalwayswilling,helpful,creativeandpresent.(ersupportwasthemostimportantmotivatortocompletethiswork.Professor George Tsatsaronis was always patient and willing to assist. (is energy andexcitementbrightenedsomecloudydays!)would like to thankProfessor Stanek for the reviewof this thesis. ) amalso thankful toProfessorZieglerforhiswillingnesstochairmythesisdefense.)amindebtedtothestudentsthatcontributedtothiswork:Adrian(iemann,RobertSelim(ill,MehmetÖzbek,andEmreTopal.And,specialthankstomycolleaguesfromtheinstitutefor the fruitful collaboration, eventfulmeetingsandniceexperiencesgathered.TheywereNidal Abboud, Max Sorgenfrei, Berit Erlach, Wanda Ali Akbar and MehrnoushSarcheshmepoor.FortheprovidedsoftwareAspenPlusofAspenTech,)amgrateful.) truly appreciate the help frommy parents and siblings, my wife and children, andmyextended families.Without their sacrifices, the completion of thisworkwould have beenimpossible.Finally, ) would like to thankmyMalaysian ikhwah, akhowat and all my friends here inGermanyandinMalaysiafortheirunderstanding,supportandencouragement.Only(im,Allah,couldrepayalltheabove‐mentioneddeedsinthisworldandhereafter.Berlin,June MohdNazriBinOmar

2

IfLiving’sSimplyToLive

EvenPigsInJungleLive

IfWorking’sSimplyToWork

EvenMonkeysWork(AMKA[ ]    

3

Synopsis

LiquefiedNaturalGas LNG intheenergysectorisseenasarealisticsourceforprovidingcleaner,smalltolargescalefuel,duetotheever‐increasingglobalenvironmentalprotectionandhigh demandof electricity at a competitive rate. )n thiswork, several LNGprocessesexistingandprospectiveonesareanalysedthermodynamicallyandeconomically.Refrigerationtechnologyisimportantinassessingtheconsideredenergyconversionplants.Thedifferentdesignsinprecooling,maincryogenicandsubcoolingcyclesamongtheLNG processes, with various pure‐ and mixed‐refrigerants, added the complexity butsignificant in their evaluations. While energetic analysis provides important informationthatcanleadstoplantgeneralperformance,forexamplefromthecumulativecoolingcurveschart, useful energy and irreversibilities of the system and system components can offerbetterassessment.Thisledtothedevelopmentofexergy‐basedanalyses,inwhichexergyoffuel and product at system and component levels derive their exergy destructions andexergetic efficiencies. )n cold production, physical exergy is split into thermal andmechanical exergies because of the wide range of working pressures and temperatures;betweenthehigh Cdownto‐ Cat theircorrespondingpressures,aswellas theirconditional ambient references. Sensitivity analysis helped against the challenges inobtaining the industrial plants datawhere confidentiality, costly information, and limitedreliable sources are in norm. Thus among the examined processes, themost efficient arethose working with the mixed‐refrigerant technology. The highest exergetic efficiencyamongallanalysedsystemsistheC MRprocessat % . mtpa followedbytheMR‐Xprocessat % . mtpa .EconomicanalysisisapproachedtoseetheLNGsystemsfurther.The main cryogenic multi‐flow heat exchangers remain the highest priced compared toothercriticalcomponentsintheirFixedCapital)nvestments.Thetotalcostsfortheanalysedsystems, from . to . $ bn/mtpa, are in rangewith available literatures consistent forgreenfieldpowerplants.AnewlydesignconceptofLNGprocess,theMR‐X,ispresentedtogetherwithsimilaranalyses.)topensupanewinsight,wheresuchdesignisfeasible,practicalandrealistictovariousclimaticandinfrastructurechallengeswhilehavingefficientandeconomichighLNGproductionstrongtocooperatewiththelatestLNGtechnologiesinthemarket. 

Zusammenfassung

In der Energiewirtschaft wird verflüssigtes Erdgas (LNG) weltweit als

umweltfreundlicher Energieträger für Klein- und Großanlagen betrachtet, wobei LNG

zur Deckung des stetig wachsenden Bedarfs an elektrischem Strom geeignet ist. In der

vorliegenden Arbeit werden sowohl bestehende als auch neue Prozesse zur LNG

Produktion thermodynamisch und wirtschaftlich analysiert.

Die analysierten Prozesse unterscheiden sich im Bereich der Vorkühlung, der

kryogenen Hauptkühlung durch Kreisprozesse sowie in der Verwendung verschiedener

Kältemitteln, die aus einem Reinstoff oder einem Stoffgemisch bestehen. Basierend auf

Energieanalysen werden Leistungen in den Anlagen zum Beispiel durch Summenkurven

für die Wärmeübertragung ermittelt. Andererseits werden die Komponenten und

Systeme durch Exergie bewertet. Die damit verbundenen exergiebasierten Analysen

nutzen den exergetischen Aufwand und Nutzen zur Ermittlung der Exergievernichtung

sowie der jeweiligen exergetischen Wirkungsgrade. Für die Erzeugung der notwendigen

Kälteleistung wird die physikalische Exergie in einen mechanischen und thermischen

Anteil unterteilt, da der große Bereich von Druck und Temperatur (142 bis -168°C) die

Umgebungsbedingungen überschreitet. Da Betreiber industrieller Anlagen viele

Betriebsparameter vertraulich behandeln und Veröffentlichungen über zuverlässige

Kostendaten kaum existieren, werden Sensitivitätsanalysen verwendet.

Generell stellten sich Prozesse mit Kältemittelgemischen als effizienter heraus.

Als exergetisch effizienteste Prozesse wurden der C3MR-Prozess mit 33% (4,5 mtpa)

gefolgt vom MR-X-Prozess mit 32% (7,8 mtpa) ermittelt. Zum weiteren Verständnis

wurden Wirtschaftlichkeitsanalysen durchgeführt. Im Besonderen sind die Kosten des

Hauptwärmeübertragers, in dem thermische Energie im kryogenen Temperaturbereich

übertragen wird, am höchsten. Die ermittelten Produktgestehungskosten liegen

zwischen 0,3 und 0,6 $bn/mtpa, was im üblichen Bereich für Neubauanlagen liegt. Unter

anderem wird ein neuer Prozess (MR-X) vorgestellt. Dabei wurde festgestellt, dass dieser

Prozess unter verschiedenen klimatischen und infrastrukturellen Rahmenbedingungen

effizient und wirtschaftlich LNG produziert.

4

Table of Contents

1. Introduction .................................................................................................................................................... 16

1.1. LNG chain ................................................................................................................................................ 18

1.2. LNG processes – global and Malaysian context ...................................................................... 19

2. Literature Review on Malaysian LNG Processes .............................................................................. 23

3. Overview on Energy, Exergy, and Economic Analyses .................................................................. 31

3.1. Energy Analysis .................................................................................................................................... 31

3.2. Exergy Analysis .................................................................................................................................... 33

3.2.1. Software requirements for simulation and exergy calculations ............................ 36

3.3. Economic Analysis .............................................................................................................................. 38

3.3.1. Estimation of Total Capital Investment (TCI) ................................................................ 38

4. Processes of Liquefaction of Natural Gas in Malaysia .................................................................... 44

4.1. Propane Pre-Cooled Mixed-Refrigerant (C3MR) LNG Process......................................... 44

4.1.1. Principle of Operation.............................................................................................................. 44

4.1.2. Simulation and Energy Analysis .......................................................................................... 45

4.1.3. Exergy Analysis .......................................................................................................................... 53

4.1.4. Economic Analysis .................................................................................................................... 54

4.2. AP-XTM LNG Process ........................................................................................................................... 56

4.2.1. Principle of Operation.............................................................................................................. 56

4.2.2. Simulation and Energy Analysis .......................................................................................... 57

4.2.3. Exergy Analysis .......................................................................................................................... 60

4.2.4. Economic Analysis .................................................................................................................... 62

4.3. MR-X LNG Process .............................................................................................................................. 67

4.3.1. Principle of Operation.............................................................................................................. 70

4.3.2. Simulation and Energy Analysis .......................................................................................... 71

4.3.3. Exergy Analysis .......................................................................................................................... 72

4.3.4. Economic Analysis .................................................................................................................... 74

5. Conclusion and Future Works ................................................................................................................. 82

6. References ........................................................................................................................................................ 85

5

Appendix A. Research Contributions ..................................................................................................... 91

Appendix B. Energy and Exergy Analyses – Data, Flow and Results ........................................ 92

B.1 General Information on Liquefaction Processes ................................................................... 93

B.2 System testing using PRICO® process [23,24,32] ................................................................ 99

B.3 C3MR Process [73] ........................................................................................................................... 103

B.4 AP-XTM Process [51]......................................................................................................................... 111

B.5 MR-X Process [56,94] ..................................................................................................................... 124

Appendix C. Economic Analysis Data and Flow – A Study Case on the C3MR Process ... 129

C.1 Purchased Equipment Costs (PEC) Estimates ...................................................................... 130

C.1.1 Heat Exchangers ...................................................................................................................... 131

C.1.2 Dissipative coolers ................................................................................................................. 131

C.1.3 Propane and mixed refrigerant compressors ............................................................. 132

C.1.4 Separators .................................................................................................................................. 133

C.1.5 Valves and mixers ................................................................................................................... 134

C.2 Estimation of Total Capital Investment................................................................................... 136

C.2.1 Calculation of startup costs (SUC) and working capital (WC) ............................. 136

C.2.2 Estimation of allowance for funds used during construction (AFUDC) ........... 138

C.3 Estimation of Operating and Maintenance (O&M) Costs ................................................. 141

C.4 Estimation of the Fuel Costs (FC) .............................................................................................. 142

C.5 Estimation of Revenue Requirements ..................................................................................... 143

C.5.1 Total capital recovery ........................................................................................................... 145

C.5.2 Returns on equity and debt ................................................................................................ 146

C.5.3 Taxes and insurance .............................................................................................................. 147

C.5.4 Fuel, operating and maintenance costs ......................................................................... 147

C.5.5 Total revenue requirement (TRR) ................................................................................... 148

C.5.6 Levelized Costs and the Cost of the Main Product .................................................... 150

6

List of Figures

Fig. 1.1. Global LNG demand [6]. ...................................................................................................................... 16

Fig. 1.2. Comparison of transportation cost [9]. ........................................................................................ 17

Fig. 1.3. The process chain for LNG - from Extraction, Processing and Transport to

Consumption [10]................................................................................................................................................... 18

Fig. 1.4 Liquefaction Capacity by Type of Technology, 2013-2018 [15]. ........................................ 19

Fig. 1.5. Malaysia is located in the Asia Pacific basin, color-coded reference by the IGU.

Modified from [15]. ................................................................................................................................................ 21

Fig. 1.6. The three Malaysian LNG process plants including their corresponding trains are

located in Bintulu, Sarawak, an eastern state of Malaysia [17]. .......................................................... 21

Fig. 3.1. A Single Cycle Liquefaction Process [60]. .................................................................................... 31

Fig. 3.2. High level view of refrigeration cycles within processes [60]. ........................................... 32

Fig. 4.1. A general schematic of the C3MR process. .................................................................................. 45

Fig. 4.2. Overall cooling curves for the simulated C3MR process [73]. ............................................ 50

Fig. 4.3. A similar curves for propane pre-cooled MR cycle versus the natural gas’ proposed

by Madhavan [81]. .................................................................................................................................................. 51

Fig. 4.4. Exergy destruction (MW) and exergy destruction ratio (%) for selected components

of the AP-X process. ............................................................................................................................................... 61

Fig. 4.5. Exergetic efficiency for selected components of the AP-X process. .................................. 62

Fig. 4.6. The estimation of direct costs for AP-X process ....................................................................... 65

Fig. 4.7. A general cooling curve for cascade type of LNG process. The smoother curve is the

NG-LNG curve and below it is the refrigerants curve [92]. ................................................................... 67

Fig. 4.8. Cooling curves for C3MR process [92]. ......................................................................................... 68

Fig. 4.9. Exergy destruction (MW) and exergy destruction ratio (%) for the components of

the MR-X process. ................................................................................................................................................... 73

Fig. 4.10. Exergetic efficiency of selected components of the MR-X process. ................................ 73

Fig. 4.11. PEC ($ mil) for selected components of the MR-X process and distribution of the

PEC among the components............................................................................................................................... 76

Fig. 4.12. Levelized total revenue requirement for the MR-X process using different

assumptions for the economic analysis: OMC as a function of CC - between 1% and 10% and

cost of the electricity – between 0.05 and 0.20 $/kWh. ......................................................................... 79

7

Fig. 4.13. Cost per unit of mass of the liquefaction process when different assumptions for

the economic analysis are used. ...................................................................................................................... 80

Fig. B.1. Classification of natural gas liquefaction processes [93]. .................................................... 93

Fig. B.2. A typical coil-wound MCHE for a C3MR process-based LNG plant [103]. .................... 98

Fig. B.3. Flow diagram of PRICO process: CM1 - Compressor 1; COL - Cooler; CM2 -

Compressor 2; CD - Condenser; HE - Heat exchanger; TV - Throttling Valve [24]. .................... 99

Fig. B.4. C3MR flowsheet using Aspen Plus [73]. .................................................................................... 105

Fig. B.5. A general schematic on the AP-X process. ................................................................................ 111

Fig. B.6. Flowsheet for AP-XTM process. ...................................................................................................... 112

Fig. B.7. �- T diagram for HEX0 (∆ ����ℎ = 0.3�).............................................................................. 113

Fig. B.8. �- T diagram for HEX1 (∆ ����ℎ = 6.4�) ............................................................................. 113

Fig. B.9. �- T diagram for HEX2 (∆ ����ℎ = 2.1�).............................................................................. 114

Fig. B.10. �- T diagram for HEX3 (∆ ����ℎ = 17.7�) ........................................................................ 114

Fig. B.11. PETRONAS FLNG to be commissioned in 2015 [15]. ........................................................ 123

Fig. B.12. A general schematic on the MR-X process. ............................................................................ 124

Fig. B.13 Process flow diagram for MR-X process .................................................................................. 125

Fig. B.14. Cumulative cooling curves for the MR-X process. .............................................................. 126

8

List of Tables

Table 1.1. Pounds of Air Pollutant Produced per Billion Btu of Energy [8] ................................... 17

Table 1.2. Typical LNG Compositions at Different Plant Locations [16]. ........................................ 20

Table 3.1. Elements of total capital investment [62]. .............................................................................. 39

Table 4.1. Estimation on Purchased Equipment Cost for selected AP-X process equipments

........................................................................................................................................................................................ 64

Table 4.2. The optimum composition of precooling refrigerants for DMR process analysed by

[93]. .............................................................................................................................................................................. 69

Table 4.3. Estimation of the fixed-capital investment. ............................................................................ 77

Table B.1. Malaysian LNG Plants. ..................................................................................................................... 94

Table B.2. Liquefaction Plants with specific LNG Technology, sorted by year of project start

[102]. ............................................................................................................................................................................ 95

Table B.3. Composition and concentration of natural gas and refrigerants. ................................. 99

Table B.4. Thermodynamic data for material streams at real operating conditions............... 100

Table B.5. Reference values for the exergetic analysis (state 0) for material streams. .......... 100

Table B.6. Detailed thermodynamic data of each chemical component in the streams within

mixed refrigerant. ................................................................................................................................................ 101

Table B.7. Definition of the exergy of fuel and the exergy of product for the components of

the PRICO® process. .......................................................................................................................................... 101

Table B.8. Results obtained from the exergetic analysis. .................................................................... 102

Table B.9. Composition for the C3MR process mixed-refrigerant in molar percentage. ....... 103

Table B.10. LMR and HMR compositions. .................................................................................................. 103

Table B.11. Boiling temperatures (in C) for refrigerants at different pressures [104]. ......... 103

Table B.12. Stream 7 molar fraction and its partial pressures ......................................................... 103

Table B.13. Enthalpy and Entropy Values Required for Exergy Calculation............................... 104

Table B.14. Thermodynamic data for the material streams at real operating conditions for

C3MR process [73]. ............................................................................................................................................. 106

Table B.15. Standard Molar Chemical Exergy Values for Selected Substances at Tref =

298.15K. Model II is referred. ......................................................................................................................... 107

Table B.16. Chemical exergy result for affected streams for C3MR ................................................ 107

9

Table B.17. Definition of the exergy of fuel and the exergy of product for the components of

the C3MR process. ............................................................................................................................................... 108

Table B.18. Exergy rate of product and fuel for the selected components of the C3MR process

..................................................................................................................................................................................... 110

Table B.19. Composition for the AP-X process in molar percentage. ............................................. 111

Table B.20. Thermodynamic data for the material streams at real operating conditions. ... 115

Table B.21. Mole flow rate of the mixed refrigerant. ............................................................................ 117

Table B.22. Definition of the exergy of fuel and the exergy of product for the selected

components of the AP-X process. .................................................................................................................. 119

Table B.23. Exergy rate of product and fuel for the AP-X process. ................................................. 121

Table B.24. Power net required by AP-X process components. ....................................................... 122

Table B.25. Composition of NG and refrigerants for MR-X process. ............................................... 124

Table B.26. Thermodynamic data for the material streams (at real operating conditions). 127

Table C.1. Parameters and assumptions used in TRR calculations [62] ....................................... 130

Table C.2. U and A values of the liquefaction heat exchangers [101]. ............................................ 131

Table C.3. The purchased equipment cost of liquefaction heat exchangers (106$). ................ 131

Table C.4. U and A values of dissipative coolers [106]. ........................................................................ 132

Table C.5. Purchase equipment cost of dissipative coolers (106$). ................................................ 132

Table C.6. The process work input (indicated and net required). ................................................... 133

Table C.7. The purchased equipment cost of the compressors (106 $). ........................................ 133

Table C.8 Sizing parameters of the separators. ....................................................................................... 134

Table C.9 Purchased equipment cost of separators (106$). ............................................................... 134

Table C.10 Purchased equipment cost of throttling valves (106$). ................................................. 135

Table C.11. Estimation of the total capital investment. ....................................................................... 137

Table C.12. The calculated values for the allowance for funds used during construction (106

$). ................................................................................................................................................................................ 139

Table C.13. Statutory percentages for use in the MACRS for a life period of 15 years, annual

tax depreciation and tax book at the end of each year for the LNG plant. ................................... 140

Table C.14. Year-by-year revenue requirement breakdown of the LNG liquefaction plant (106

$) ................................................................................................................................................................ ................. 144

Table C.15. Year by year capital recovery schedule for the LNG plant. (106 $) .......................... 146

Table C.16. Distribution of capital recovery for the LNG plant (106 $). ........................................ 149

Table C.17. LNG plant data set [98]. ............................................................................................................. 152

10

Table C.18. Economic data for the selected LNG plants [98]. ............................................................ 153

11

Nomenclature

Abbreviations

AC aftercooler

AFUDC allowance for funds used during construction

APCI Air Products and Chemicals Inc.

bn billion

C3MR propane pre-cooled mixed refrigerant cycle

CD condenser

CEPCI chemical engineering plant cost index

CI cost index

COMP compressor

COP coefficient of performance

CWHE coil wound heat exchanger

DMR dual mixed refrigerant cycle

EV evaporator

EXP expander

FCI fixed capital investment

FLNG floating liquefied natural gas

HEX pre-cooling heat exchanger

HEX heat exchanger

HMR heavy mixed refrigerant

HPN2 high pressure nitrogen compressor

IC intercooler

LMR light mixed refrigerant

LNG liquefied natural gas

LPG liquefied petroleum gas

LPMR low pressure mixed-refrigerant compressor

LPN2 low pressure nitrogen compressor

12

MACRS modified accelerated cost recovery system

MBtu million British thermal unit

MCE main cryogenic exchanger

MCHE main cryogenic heat exchanger

mil million

MIX mixer

MLHE main liquefaction heat exchanger

MLNG Malaysian Liquefied Natural Gas plant

mmtpa million metric tonne per annum

MPMR middle pressure mixed-refrigerant compressor

MPN2 middle pressure nitrogen compressor

MR mixed-refrigerant

MT million ton

mtpa million ton per annum

N2 nitrogen

NG natural gas

NGL natural gas liquids

PEC purchased equipment cost

PETRONAS Petroliam Nasional Berhad

PFHE plate fin heat exchanger

PHX pre-cooling heat exchanger

PMR parallel mixed refrigerant cycle

PPHE perforated plate heat exchanger

SEPA separator

SMR Single Mixed-Refrigerant

SRK Soave-Redlich-Kwong

TCI total capital investment

TV throttling valve

VALVE throttling valve

13

Symbols

A heat transfer surface area [m2]

C cost rate associated with exergy transfer [$/h]

c cost per unit of exergy [$/GJ]

E time rate of exergy transfer [kW]

e specific exergy [kJ/kg]

F future value of money [$]

f exergoeconomic factor [%]

H rate of enthalpy [kW]

h specific enthalpy [kJ/kg]

ieff effective annual discount rate [%]

k variable used in levelized cost calculations [-]

m mass flow rate [kg/s]

n number of time period [year]

Q heat transfer rate [kW]

p pressure [bar]

P present value of money [$]

r relative cost difference [-]

S rate of entropy [kW/K]

s specific entropy [kJ/kgK]

t income tax rate [%]

T temperature [K, ⁰C]

U overall heat transfer coefficient [W/m2K]

V volume [m3]

v specific volume [m3/kg]

W work [MJ]

W work rate [MW]

X variable represents the size of equipment [-]

Dy exergy destruction ratio [%]

Ly exergy loss ratio [%]

14

Z non-exergy related cost rate [$/h]

Greek letters

α capacity exponent [-] Δ difference [-] ε exergetic efficiency [-]

mechη mechanical efficiency [%]

isη isentropic efficiency [%]

δ variable used for cost difference in coolers [-]

γ activity coefficient [-] τ average annual plant operation hours [h]

Subscripts

BM bare module

ce common equity

CM compressor machine

d debt

D exergy destruction

F fuel

e outlet stream

el electricity

i specified state

i inlet stream

is isentropic

j stream of matter : year

k component of the plant

L levelized value : exergy loss

M material factor

mech mechanical

15

NG natural gas

out exiting exergy

OTXI other taxes and insurance

P product : pressure

PH physical exergy

ps preferred stock

r real escalation

ref reference state

tot overall system

x type of financing

0 ambient/environment state

1-2 control volume inlet and outlet

Superscripts

CH chemical exergy

CI capital investment

KN kinetic energy

M mechanical exergy or cost

OM operating and maintenance

PH physical exergy

PT potential exergy

T thermal exergy

TOT total cost of the stream

16

1. Introduction

Energy worldwide is in demand exponentially [2,3] and expectably will increase even

speedier in the near future while oil reserves are depleting and alternative sources are cost-

challenging. Furthermore, the world’s desire for cleaner type of fuel with numerous

governments regulating environmental policies and incentives for green technologies has

pushed natural gas as an exciting solution.

Natural gas is not new in providing solutions to mankind, as a record shows the

Chinese have applied it commercially some 2 400 years ago [4] and was highly consumed

during the post Second World War. Such abundant source is now providing 23% of the

world’s total energy supply. Its applications currently include but not limited to electricity

generation, grid heating as well as domestic needs. The liquefied natural gas (LNG) is an

enhanced type of such energy, yielding approximately 40% more heating value than any

liquid fuel derived from the chemical conversion of natural gas [5]. In February 2015 BP

showed as per Fig. 1.1 the current and projected global demand for LNG.

Fig. 1.1. Global LNG demand [6].

Among characteristics of LNG are odorless, colorless, shapeless, and lighter than air.

These are advantageous when compared to the solid coal or the liquid oil, especially on the

aspect of greenhouse gas emission (where oil and coal produced 1.4 – 1.75 times more of

CO2), as compiled in Table 1.1. Since January the 1st, the Europeans has had recently

regulated stricter emission control under the International Convention for the Prevention of

1 . I n t r o d u c t i o n

17

Pollution from ships (MARPOL), referring to Sulphur Oxides (SOx) pollution. Thus, LNG-

fueled vessel has become a preferable option due to its low sulphur emission [7].

Table 1.1. Pounds of Air Pollutant Produced per Billion Btu of Energy [8]

Pollutant Natural Gasa Oilb Coalc

Carbon

117 000 164 000 208 000

Carbon

40 33 208

Nitrogen

92 448 457

Sulphur

0.6 1 122 2 591

Particulates 7.0 8.4 2 744

Formaldehyde 0.750 0.220 0.221

Mercury 0.000 0.007 0.016

a Natural gas burned in uncontrolled residential gas burners.

b Oil is # 6 fuel oil at 6.287 million Btu per barrel and 1.03% sulphur

with no post-combustion removal of pollutants. c Bituminous coal at 12,027 Btu per pound and 1.64% sulphur with

no post-combustion removal of pollutants.

Fig. 1.2. Comparison of transportation cost [9].

Apart from yielding more heating, having natural gas in liquefied form means higher

significant volume of energy is transportable. It contains no traces of acid gas and water.

The natural gas is transformable into the commercialised LNG due to the natural gas’

density factor of 1/600th at same pressure conditions. In other words, LNG takes about

1/600 spaces required by natural gas. Such increment in density is profitable when

transporting the LNG via ships and trucks. Fig. 1.2 illustrates the transportation cost against

the distance for shipping and on- and offshore pipeline. As shipping the LNG is more

favorable than the other two options, especially when the distance is more than 3000 km,

more possibilities are opening up. Such opportunities are like exploring stranded gas fields

using floating LNG (FLNG) platform, and LNG ships could reach various customers at

different distances. While criminal activity is a risk that accounts for both shipping and

pipeline options, shipping is preferred when spot delivery, as well as peak shaving purpose,

are in demands.

1 . I n t r o d u c t i o n

18

1.1. LNG chain

LNG liquefaction and transportation (ship or truck) become economically reasonable when

the reserves size justifies the principal investment of an LNG process plant [4]. The

liquefaction process refrigerates in a particular plant; either peak shaving plant or baseload

plant depending on factors for example customer demand, reserves location, geographical

and economic challenges as well as political conditions. Storage applications are essential

prior to LNG transport and when receiving it for regasification (for sale). Fig. 1.3 shows the

LNG production chain.

Fig. 1.3. The process chain for LNG - from Extraction, Processing and Transport to

Consumption [10]

In the above process chain, after being pointed by geo-exploration team, the raw

natural gas is drill-extracted from its reserves in the earth. The raw gas is firstly fed into a

purification part of the liquefaction plant to remove unwanted materials. Treating the gas is

important to ensure it has as much methane as possible, and contaminants do not interfere

in the process of achieving the optimum temperature for LNG (at about -162 C). Treated

natural gas then enters into the main liquefying phase of the plant. Depending on the

installed refrigeration technology of plant, in modern system the natural gas undergone a

typical pre-cooling by selected refrigerant(s) before proceeds to main cryogenic heat

exchanger (MCHE) for a deeper liquefying and before being sub-cooled. Once the conditions

(pressure and temperature) are met, the liquefied natural gas is kept in storage due to be

shipped at designated times. The LNG ship that will carry the LNG is designed specially,

particularly its tanks, to cater the behaviors of natural gas such as boil-off and regasification

during transport. Based on contracts, the LNG is transported to worldwide customers at

mutually agreed prices. The transported LNG is then being stored before it is regasified into

commercial gas for various needs. Japan, for example, has been among the biggest world’s

LNG importer which Malaysia are among its trusted exporters for the past decades.

Energy Information Administration [11] has reported the costs throughout the LNG

chain in detailed, however, in brief:

1 . I n t r o d u c t i o n

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• Production of natural gas. The cost of carrying (including all related processing)

natural gas from the reserves to the LNG process plant is 15-20% of the total cost.

• LNG process plant. The cost for all related processes (treatment, liquefaction, by-

production, storage and ship-loading) is 30-45% of the total cost.

• Shipment of LNG. The shipping cost for LNG is 10-30% of the total cost.

• LNG receiving terminal. The cost for all related processes (ship-unloading, storage,

regasification and sales) is 15-25% of the total cost.

The liquefaction plant part of the LNG production chain takes the highest value compared to

the other parts. Among the factors contributing to such cost are the strict design and safety

standards as well as the remote location aspect. Jenson estimated (personal

communication) a base case LNG project particularly on liquefaction part is about $350

million in CAPEX (for a greenfield infrastructure plus $250 million/ton of LNG capacity) and

about $0.20/mmBtu in OPEX [12]. Referring to an Indonesian LNG facility [13,14], the cost

for three trains used to treat and liquefy LNG is $202 million from the total $869 million of

CAPEX, the highest cost contributor comparing to other project components. As such, the

selection of liquefaction technology is the most critical in the whole LNG chain. Proper

classification of the liquefaction technology or process is imperative to ensure the desired

LNG process plant is sustaining as long as possible. Fig. B.1 (Appendix 1) shows a general

classification of natural gas liquefaction processes.

1.2. LNG processes – global and Malaysian context

Fig. 1.4 Liquefaction Capacity by Type of Technology, 2013-2018 [15].

There are at least nine different LNG processes operating around the world. Technologists

are striving in giving better solutions to the increasing demand for better quality of LNG

1 . I n t r o d u c t i o n

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product while at the same time cheaper and long-lasting. Such efficiency is rapidly seen in

patents and in the industrial world hence competition to achieve higher efficiency of LNG is

inevitable. Fig. 1.4 describes the liquefaction capacity made and forecasted versus the type

of LNG processes.

It can be seen that several processes, such as the APC C3MR/Split MR process and

ConocoPhillips Optimized Cascade process are expected to increase in capacity per year,

whereas others are anticipated to be stagnant and decreasing over the years. These may be

contributed by the increment of LNG production capacity (C3MR/Split MR and Optimized

Cascade), and by the emergence of new LNG process namely Floating LNG process.

Individual processes are decreasing per world’s percentage capacity due to stagnancy. This,

for example refers to the 51% to 38% (estimated) LNG capacity for Air Products’ C3MR

process, in which it will still stay at about 150 mtpa for the next four years to come. The

various kinds of LNG processes provide higher chances for clients (gasfield owner) to have

the best technology that suits their conditions and limitations.

Among the challenges gasfield and process owners are facing are the different

compositions of natural gas extracted. Other components of natural gas may be processed

into commercial by-products, like butane and propane gases. Nevertheless, the lesser these

elements contained in the natural gas the higher percentage of methane gas, and

consequently the cheaper it is for treatment and fractionation, as well as higher volume of

LNG could be produced for sales. Specific processes are necessary to handle the light and

heavy hydrocarbons. The final LNG product, therefore, has different compositions when

compared to other locations. Table 1.2 shows such typical comparison.

Table 1.2. Typical LNG Compositions at Different Plant Locations [16].

Component,

mole %

Das Island,

Abu Dhabi

Whitnell

Bay,

Australia

Bintulu,

Malaysia

Arun,

Indonesia

Lumut,

Brunei

Bontang,

Indonesia

Ras

Laffan,

Qatar

Methane 87.10 87.80 91.20 89.20 89.40 90.60 89.60

Ethane 11.40 8.30 4.28 8.58 6.30 6.00 6.25

Propane 1.27 2.98 2.87 1.67 2.80 2.48 2.19

Butane 0.141 0.875 1.36 0.511 1.30 0.82 1.07

Pentane 0.001 - 0.01 0.02 - 0.01 0.04

The different of LNG composition can be associated with the respective basins namely

Atlantic-Mediterranean basin, Middle East basin and Pacific basin (Fig. 1.5). From Table 1.2,

Malaysia which has the highest methane percentage is located within the Pacific basin.

1 . I n t r o d u c t i o n

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Fig. 1.5. Malaysia is located in the Asia Pacific basin, color-coded reference by the IGU.

Modified from [15].

Fig. 1.6. The three Malaysian LNG process plants including their corresponding trains are

located in Bintulu, Sarawak, an eastern state of Malaysia [17].

Malaysia since 1983 has started to export LNG to Japan under a 20-year contract

[18]. Japan, Taiwan and Korea are the leading importers of Malaysia’s LNG. The Malaysian

LNG’s plant located in Bintulu, Sarawak (Fig. 1.6) is among the biggest single concentrations

of LNG production capacities in the world [19]. The plant was a joint venture between the

national petroleum company PETRONAS, Shell B.V and Japan’s Mitsubishi Corporation.

MALAYSIA

1 . I n t r o d u c t i o n

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Several characteristics of Malaysian LNG plants are illustrated in Table B.1. As PETRONAS is

expecting their ninth onshore production train completion (adding 3.6 mtpa in capacity),

their PFLNG 1 and 2, built in South Korea, will start producing LNG in 2015 and 2018,

respectively. These will make the global petroleum conglomerate the first company in the

world to bring an FLNG facility to the market. The PFLNG1 will operate at the Kanowit field

offshore Sarawak, and the PFLNG2 will run in the Rotan field offshore Sabah, Malaysia

[20,21]. Both plants will increase Malaysia’s total LNG production capacity from 25.7 mtpa

to 26.9 mtpa and is expected to rise to over 30 mtpa in total by 2017. The industry players

are paying full attention to see how the world’s first Malaysia’s PFLNGs and their

technologies will perform.

23

2. Literature Review on Malaysian

LNG Processes

There are not many specific studies involving MLNG processes which are openly accessible

and at the same time reputable. Although there are numerous publications pertaining to the

market state of MLNG, technical reports about its liquefaction processes are limited.

Following Fig. B.1, a classical cascade type of liquefaction of natural gas is one of the

first LNG processes. It is a three-cascade refrigeration system with single refrigerant for

each cascade. Beside several publications exist in 1960s, this process has been examined by

Morosuk et al. [22] in view of Malaysian operating and environmental conditions. The

analysis includes up to advanced exergetic analysis to reveal the interactions among system

components. The advanced analysis also aided the authors to show potential for improving

the components and overall system thermodynamic efficiencies. Three-cascade

refrigeration system is used. The selection of cascade type of LNG plant may not represent

accurately the existing MLNG plants, but provides bases for a proper and reliable flow of

analyses and methodologies. It is also advantageous for understanding the liquefaction

process system and their process equipment, using explicit EES simulation software, and for

learning necessary assumptions and system boundaries for analyses. Important and less

important liquefaction equipment (for improvement) were found based on their exergy

destructions and exergetic efficiencies. Three components were established as the most

important components for improvements (compressors CM1, CM2, and heat exchangers

CD1-EV1, CD2-EV2). Summarily, conventional exergetic analysis provides useful

information but an advanced exergetic analysis makes such information more precise,

useful and supplies additional information that cannot be supplied by the conventional

analysis. The advanced exergetic analysis is also able to correct the ranking of a certain

component that was initially found from the conventional analysis result (throttling valves

TV2 and TV3). Only CH4 is used as the composition for natural gas thus creates potential for

further improvement where one may replace such composition with a more mixed-type of

natural gas, such as presented in this thesis. The mixture of several components in a gas, or

in a refrigerant, proves, however, more complicated for analyses and, therefore, requires

careful selection of simulation software, data sources, and results interpretations. Apart

from the single-material NG composition, other initial data and assumptions are in the range

of existing MLNG plants data. Among valuable data that validates and support this thesis

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data are the outlet temperature for cooling water, the exergetic efficiency and the system

coefficient of performance.

An insight over a type of LNG process called Single Mixed Refrigerant (SMR) is

necessary. This is because the current process for MLNG plants is among the close

successors to the SMR process. Experiences learned from the plants that used the single-

cycle SMR process helped technologists to develop the two-cycle C3MR process. PRICO is a

popular process that uses SMR technology. This process has been drawn up by the

Black&Veatch Company, and the industrial applications of PRICO started in the year 1955,

when it was applied to one of the first world’s LNG plants [23–25]. Three U.S./international

patents cover the PRICO process. At present at least 21 LNG plants use this process while 16

more plants are in the design and/or construction phase. The PRICO process is famous for

LNG peak-shaving units. In the year 2010, 25% of the LNG plants in the U.S. used this

process. Within two years after that, design and construction for the world's first offshore

LNG project started [25]. The following advantages are associated with the PRICO process

[26]: 1) proven process that achieves the promised performance, 2) relative simple

operation, 3) minimal refrigerant inventory, 4) reduced number of equipment items, 5) low

capital cost and operating cost, 6) high flexibility, 7) high reliability, and 8) rapid startup.

There are not many research publications dealing with liquefaction processes;

however, the PRICO (SMR) process recently became quite popular among researchers. Four

processes for small-scale LNG plants were evaluated by Remeljej and Hoadley, 2006 [27].

The PRICO process was selected there as a reference process. An exergy analysis was

performed in a simple way, and only relative data are given. The exergy destructions

(thermodynamic inefficiencies) are distributed as follows: 21% - within both compressors,

30% - within both coolers, 46% - within a heat exchanger, and 3% - within the throttling

valve. The paper concluded that, among all studied processes, the SMR process has the

lowest exergy consumption for the compressors, and that the main difference between the

processes was caused by efficiency differences of the expander-driven compressors. Jensen

and Skogestad, 2009 [28] discussed eight compositions of the mixed refrigerant that usable

for the PRICO process; the effect of properties of the mixed refrigerant to the main

characteristics of the PRICO process were reported. The authors demonstrated that

increasing the concentration of nitrogen within the mixed refrigerant leads to an

improvement in the heat-transfer performance of all heat exchangers. An application of the

gradient-free optimization-simulation method to processes modeled with the simulator

Aspen HYSYS is reported by Aspelund et al., 2010 [29]. The PRICO process was selected as

an academic example for the optimization for two reasons: firstly, this process is a simple

LNG process with seven independent variables (opted by the authors). This number is too

large for the optimization routine but small enough to be optimized with an optimization-

simulation tool. The second reason is that it is possible to verify the results by investigating

the hot and cold composite curves. The paper focused on the number of iterations required

to get an optimal concentration of the mixed refrigerant. Mokarizadeh Haghighi Shirazi and

Mowla, 2010 [30] discussed the simulation of SMR concepts and the properties that are

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used in MATLAB to generate the objective function. A genetic algorithm was used for

optimization. The energy consumption of the process was minimized. Depending on the

concentration of the refrigerant, the specific energy consumption can be reduced from 1485

kJ/kg LNG to 1186.6 kJ/kg LNG or from 1126.7 kJ/kg LNG to 1092.4 kJ/kg LNG. The smaller

values were taking from Lee, 2001 [31] (as a reference publication). The authors applied

also an exergy analysis, in order to calculate the values of the exergy destruction within the

components: 31% - within both compressors, 33% - within both coolers, 27% - within the

heat exchanger, and 9% - within the throttling valve. Hiemann, 2011 [32] conducted a

detailed exergy analysis of the PRICO process. Here the approach “exergy of fuel/exergy of

product” has been used taking into account a splitting of the physical exergy into thermal

and mechanical parts. Marmolejo-Correa and Gundersen, 2012 [33] selected the PRICO

process as an academic example to demonstrate the effect of using different approaches in

the exergy analysis (“inlet exergy/outlet exergy” versus “exergy of fuel/exergy of product”

as well as splitting of the physical exergy into thermal and mechanical parts) on the

obtained results. The authors assumed the operation conditions without necessarily a

reference to real plants. Xu et al., 2013 [34] reported the results of the optimization of the

concentration of the refrigerant as a function of the inlet temperature to the heat exchanger

(263.15 K through 313.15 K). For the optimization, a genetic algorithm coupled with the

process simulation software Aspen Plus has been used. The results show that when the

ambient temperature increases, the concentrations of methane, ethylene and propane

should decrease while the concentration of isopentane should increase. In this way, the

overall exergetic efficiency can be increased from 30% (calculated by the authors for the

commercial concentration of the refrigerant) up to 39.6-42.3%. In this paper, the exergetic

efficiency is a function of COP and of a “correlation factor”. In their follow-up paper, the

effect of concentration on each working fluid within the mixed refrigerant was investigated.

Such is to minimize the specific power consumption (the value of 1003.6 kJ/kg LNG was

reached), i.e. maximize the values of COP and exergetic efficiency. The reported value of

COP=0.782 is surprisingly high in comparison with results reported in other publications;

however, the definition of COP is not given. The exergetic efficiency was calculated as 43.9%,

which is in the range of other available data for the PRICO process. The distribution of the

exergy destruction within the components is as follows: 36% - within both compressors,

27%- within both coolers, 26% - within the heat exchanger, and 11% - within the throttling

valve. Sequential quadratic programming was also applied to the optimization of the PRICO

process (Morin A. et al., 2011) [35]. The research focused on the method used for

optimization. The optimization results related to the liquefaction process itself were

discussed very briefly for the two study cases, in which the mixed refrigerant is with and

without pentane. Through the energetic optimization, the specific energy supply decreases

by 3.12%. Again the same optimization procedure for the PRICO process was reported by

Wahl et al., 2013 [36]. The optimal composition of the mixed refrigerant was a function of

the composition of the natural gas (so-called “lean natural gas” and “rich natural gas”). The

heat-transfer characteristics for the multi-flow heat exchanger are also discussed. The main

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goal of the authors was to get the results of the optimization within a short period of

execution time, in comparison with the optimization procedure discussed by Aspelund et al.,

2010 [29] that required 12 h. Castillo and Dorao, 2012 [37], discussed economic issues

related to LNG processes. They reported the application of Decision-Making (using a Genetic

Algorithm binary coding and Nash-GA) for the PRICO process. The LNG markets were also

considered in the optimization of the PRICO process. Only relative economic data are

reported, for example, the cost of the multi-flow heat exchanger is approximately 10-15% of

the total investment cost and the cost associated with the compression process is always the

dominating factor for all approaches used in the optimization. Khan et al., 2012 [38]

discussed the optimal composition of the mixed refrigerant for the SMR process from the

energetic point of view, i.e. through the minimization of energy consumption for the

compression process (from 1600 to 1528 kJ/kg LNG). The log mean temperature difference

within the multi-flow heat exchanger is 7.8 K. The SMR process was modeled in the UniSim

Design simulator, and the model was optimized with nonlinear programming. The exergy

analysis was implemented into the described optimization methodology (Khan et al., 2013

[39]) and more complex mixed refrigerant processes were optimized. Heldt, 2011 [40]

developed and tested a mathematical model for control strategies, in order for the SMR

processes to operate at optimal conditions. High attention was given to the modeling of the

multi-flow heat exchanger based on industrial experimental data. The literature review for

the evaluation of the PRICO (SMR) process shows that mainly energetic optimizations were

discussed using different methods for the mathematical optimization and corresponding

algorithms. Sometimes the selected method for optimization and its

improvement/robustness were more important to the authors that the obtained results

related to the PRICO process. The objective function of the optimization refers mainly to the

composition of the mixed refrigerant. An economic analysis is not very common for the

evaluation of the PRICO liquefaction process. Morosuk et al. [24] reiterated the variety

advantages that PRICO process has, and added that recently it has become popular among

researchers. As the simulated process was a two-staged compression ( =1CMW 44.70 MW

and =2CMW 46.51 MW), the energetic analysis showed among others that 5.6 MW more

would be required if one-stage compression was applied. From their evaluation, the PRICO

process productive components (compressors, heat exchanger and throttling valve) have

high exergetic efficiencies (80-90%). More interesting and useful information can be

obtained on the interdependence between the components and the real potential for their

improvement through advanced exergetic analysis. Conclusively, PRICO, in general, is well-

designed in terms of thermodynamics and economics while its heat exchanger could be

improved further due to the intensiveness of energy, cost, and environmental impact. These

are, after all, common challenges to all LNG process plants. The evolution from a single-cycle

to the two-cycle proved to be significant worldwide as process owners continuously seek

possibilities for higher capacity production. Since 1970s, the SMR process plants with

capacities about 1 mtpa were quickly replaced by the C3MR process [41], with Brunei

Lumut 1 plant as the first utilising the C3MR (in 1972) [42].

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The C3MR process was developed for improving the SMR process. The

improvements were cycle efficiency and increasing capacity potential per LNG train. AP

reported [43] that the SMR train at Libya had a capacity of around 0.6 mtpa per train, and

the first C3MR process plant at Brunei is more than 1 mtpa per train of capacity. This

improvement of at least 0.4 mtpa was mostly caused by the reduction in MR volumetric flow

due to the propane pre-cooling configuration. The reduction also effectively debottlenecked

the MCHE and MR compression equipment design. From the experiences gained in the

Brunei Lumut 1 C3MR process as well as in other parts of the world, process developers and

owners with Malaysian Government started the MLNG complex. Kasmuni et al. [44]

reported the historical growth of MLNG Complex. The three plants - MLNG Satu, MLNG Dua,

and MLNG Tiga, and their expansions (1983) are detailed down to the utilities and facilities.

The type and arrangement of coolers, staged-compressors, and cryogenic heat exchangers

are revealed. All three MLNG plants have common C3MR process as their primary LNG

production technology. The authors compared installed components between the C3MR

plants and their advantages and disadvantages. This is because each MLNG plant and train

has a different arrangement of drivers and compressors, different in type of turbines, and

some have been debottlenecked for more production. For example, the plants use a similar

kind of heat exchanger for the Main Cryogenic Heat Exchanger (MCHE) that is of spiral

wound type. However, with different ‘bundle design’ (the MLNG Satu with a 3-bundle design

and the MLNG Dua with 2-bundle design), the former has 4 warm-end feed circuits (NG,

light MR, heavy MR and a low pressure LPG reinjection circuit) and the latter has 3 warm-

end feed circuits. This warm-end feeds help researchers to simulate the MCHE more

accurately in the simulation software. Incorrect match between warm and cold feeds in the

simulation interface results in crossover of the temperature progress for the particular heat

exchanger. Crossover issue prevents simulating system to converge fully. Using the

experiences gained from MLNG Satu and MLNG Dua, MLNG Tiga has the preferred 3-bundle

design with 3 warm-end circuits (warm, middle and cold). Kasmuni et al. also show the

number of propane compression stages that differs between the plants, from 3 stages to 4

stages of compression. The useful information about the (maximum) temperature of

seawater cooling was found. Related economic views are presented, and some challenges

due to the continuous modifications of the MLNG plants are given. The information about

the mixed refrigerant and the reason behind the selection of C3MR process are, however,

unavailable. It may be helpful if the authors also present the different design schematics

between the existing MLNG plants. The schematics should vary from plants to plants over

the past decades of installation, retrofitting and debottlenecking. The starters and turbine

drivers are distinguished as well. MLNG Satu uses the steam turbine to power the

compressors while MLNG Dua and MLNG Tiga use gas turbine driven compressors. The

application of gas turbine on the LNG plants proved better efficiency in fuel and reduced the

complexity, compared to the steam-based plant. This is because the gas turbine is directly

connected to the compressor shaft, and the absence of boilers (steam generation)

equipment, steam cycle water treating water facilities, as well as overall cooling

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requirements for the plant, were significantly reduced. While using steam turbine as the

driver has the advantages of high reliability and availability, it may require a large space for

its operation and maintenance [45]. Meanwhile, the intensive maintenance and high CO2

emissions related to the gas turbine are still challenges need to be further improved.

MLNG Satu initially commissioned at a total capacity of 8.1 mtpa, has since de-

bottlenecked to yield a further 30% of its initial capacity. This significant increment was

done using operating experiences at other plants existed at that time, including the Brunei

LNG plant [46]. The steam-based plant is cooled using seawater heat exchangers while

MLNG Dua and MLNG Tiga use hybrid (sea water and air) heat exchangers to cool the gas

turbine based plant.

Awang [47] in the LNG Journal explained the growing challenges of heat exchanger

operations at MLNG plants. Some relevant information for future simulation works is

shown, for example, the detailed components specifications and characteristics. The

methods and modifications of the seawater cooling line were presented. The report focused

mainly on maintenance and inspection especially on the main cryogenic heat exchanger.

Hence, only a few important aspects are significant to this thesis though one may see from

the author’s recommendations the complexity of analysing the internal side of LNG heat

exchangers. Potential for future works from this thesis is seen possible by referring to the

subject of Awang’s report.

Norrazak et al. (1998) [48] define in their report the fundamental features of the

MLNG Tiga project during its development with particular emphasis on the integration

aspects of the previous existing two plants. The particular LNG process is mentioned

together with its cooling and driving types. The train fuel efficiency is claimed to be

enhanced by the use of gas turbine exhaust heat waste, but specific value or percentage of

such efficiency cannot be found. Although stated as simplified, the C3MR process flowsheet

revealed in their paper is comprehensive comparing to other available sources. It is

significant to note the mention of the Main Cryogenic Heat Exchanger (MCHE) that was

supplied by APCI. This confirmation is important because it came from the client itself

(PETRONAS), and, therefore, all data from APCI websites and white papers became more

reliable for reference purpose. There are other MCHEs produced by different cryogenic-

based component companies with various efficiencies and qualities. Norrazak et al. also

confirmed the 4-stage propane precooling compressor (while the previous two plants are 3-

stage precooling compression [44]), and the liquefaction and endflash compressors are

driven by an average specific power of 12.4 kW/tpd (or around 34 MW/tpy), in which they

claimed to be low.

Norrazak et al. (2004) [49] in another paper presented the usage of dynamic

simulation for research of plant design and verification of the plant performance, claiming it

to be effective. The simulation model included a full representation of the propane, mixed

refrigerant and the natural gas circuits. At the particular year mentioned, the MLNG Tiga

trains have been among the largest ever built with an annual capacity of 3.9 mtpa. The type

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and number of compressor used in this plant are revealed; (apart from one 4-stage

centrifugal type mentioned above) one low pressure axial type MR compressor and one high

pressure centrifugal type MR compressor (connected both by a common shaft). The

arrangement of compressors and drivers is presented in a schematic. An example of simple

propane refrigeration system was illustrated however this was not part of the MLNG Tiga

plant, but only to show the use of dynamic simulation. From there the compressor

performance was shown. The authors reported in detail on how to simulate a real

complicated system, component by component design and head curves. Interestingly, for

heat exchanger models they used a fixed value for UA in the LMTD calculation. This value

was retrieved through the physical property from their database. The value was not shown

in the report. In estimating the costs for LNG heat exchangers, among important factors are

the UA values. In this paper, several parameters for example MR composition and LP MR

compressor are permitted to be manipulated while others are restricted during the quest of

searching the maximum value LNG production. Among the restrictions are all temperature

approach in the MCHE should be higher than a predetermined minimum, typically 1 or 2

degree Celsius, and compression power must be less than or equal to the maximum

allowable power. This information is valuable for researchers to manipulate parameters

while restricting other variables during software simulation, especially whenever data

availability and reliability are difficult to have or to validate. In the same lengthy report,

Norrazak et al. established a set of steps for verification of system and failure scenario. The

final section may as well be another potential for future works in furthering LNG process

research and development particularly in health, safety and environment aspects.

It is widely known about the C3MR to AP-XTM system evolution, but it is not so for

the AP-XTM to AP-NTM evolution. PETRONAS will be using the latter process in their floating

LNG production entirely soon, and it is worth to understand the AP-XTM first due to the

implication similar to SMR-C3MR process evolution mentioned earlier. The AP-XTM has been

reported significantly only by its investors [26,50] and Omar et al. [51]. In the latter study,

the system was simulated using Aspen Plus and analysed energetically and exergetically in

detail. The AP-XTM used Propane (pre-cooling), MR (liquefying) and Nitrogen (sub-cooling).

The addition of the third cycle (the Nitrogen cycle) is the main evolution from the two-cycle

C3MR process. The process is presented in this thesis after the C3MR process section. While

the COP of the AP-XTM was found to be 0.14, and its overall exergetic efficiency was 15%, the

exergetic efficiency for each component of the system is in between 65% to more than 95%

individually, which is a high percentage range for such type of efficiency. Of the total exergy

destruction, 28% is associated with the main heat exchanger and two multi-stage

compressors. A further evaluation of this process should be conducted, suggested by the

authors, using advanced exergy analysis where the interdependence between the

components as well as the real potential for improving the overall system will be

discovered. AP-XTM so far is claimed by reputable sources to be the largest operating

liquefaction capacity production. Based on such success and experience, the company built

the latest AP-NTM, which has been confirmed for PETRONAS’ FLNG application [21].

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The Borneo Post [52] recently reported about the two PETRONAS’ FLNGs (PFLNG)

that will increase the current Malaysia LNG production output of 25.7 mmtpa. The FLNG1

will contribute 1.2 mmtpa and FLNG2 1.5 mmtpa. The FLNG vessels will use AP-NTM

technology [53–55] while Dual-Mixed Refrigerant (DMR) technology is being seen as a

favorite to be integrated into this newest area of LNG production. However, note that the

DMR process was developed for Arctic climate operation conditions - low average annual

temperature but relatively high-temperature differences during the year [56]. Therefore,

the use of mixed refrigerant instead of propane refrigerant as their pre-cooling refrigerant

proved to be critical due to the negative effect of propane, for example when operating

below -30 C. Meanwhile, the AP-NTM which evolved from the AP-XTM, (the former) used only

Nitrogen (which has very low boiling temperature) to liquefy and sub-cool NG. With this,

PETRONAS will be the world pioneer in the LNG production via FLNG. Such move is

expected to transform Malaysia into an ‘LNG import-free nation’ by 2016 [57] as currently

Malaysia regasifies short-term contracted LNG from overseas for domestic usage. The

PFLNG is expected to accommodate the regasification storing capacity of 3.8 mtpa. The

FLNG market is young and has enormous potentials to be explored in terms of process

integration and innovation. It is interesting to explore and design a new concept of

liquefaction of natural gas that could integrate or retrofit C3MR, AP-X, and DMR together, for

better capacity and efficiency. A new design is presented in the final part of this thesis as

MR-X with simulations, analyses and discussions.

Weems and Sullivan [58] in 2014 presented the “LNG at 50 - History and Projected

Future for Liquefied Natural Gas Exports in an Unconventional Era” at an annual meeting of

Rocky Mountain Mineral Law Institute. They set their report based on decades of LNG

history, and Malaysia’s are shown grouped together with Australia, as major LNG exporters.

Classic and on-going contracts are detailed with relations to the U.S. and European markets.

Current and planned LNG technologies are provided, especially the Malaysian floating

storage and regasification vessels (FSRUs). It is expected that the report is not technically

explored due to the wide breadth of scopes covering many countries. Nevertheless, charting

prices of LNG per producing/importing countries for the past 50 years may prove

important, particularly when reduction of capital investment for plant construction is taken

into consideration. The absence of such prices may due to the issue of confidentiality.

Nonetheless, Damon Evans [59] reported some gas prices in which the Malaysian

power producer Tenaga Nasional Berhad (TNB) pays at 4.53 $/mmBtu (or 0.016 $/kWh).

Even though in the process chain the final LNG product is several steps behind the

commercially gas product, the pricing may not be far in range of the said price due to factors

for instance government subsidy, transport distance, and long-term contract incentives.

Lately it is reported that Malaysia is projected to be fully independent of LNG (or rather

having no LNG import at all) in a year thus such condition reduces the risk of global LNG

price volatility [57] at least for Malaysia.

31

3. Overview on Energy, Exergy,

and Economic Analyses

3.1. Energy Analysis

In discussing the LNG technologies used in Malaysia, it is important to understand first the

basics of a liquefaction process. The success of a liquefaction process is influenced much by

the number of cycles it has. An example of a cycle is shown in Fig. 3.1.

Fig. 3.1. A Single Cycle Liquefaction Process [60].

Pre-treated natural gas is fed at warm temperature into the cycle, and the cycle cools

it until it becomes liquefied. For the cycle to have the cooling, Work, � is put into it using

compressor, and heat must be rejected from it using air or water cooler. The compressor

usually has its dedicated working fluid or refrigerant. Hence the compressor (size, type), its

refrigerant (flowrate, composition), and its driver (size, type) that contribute to the amount

of work � are key factors for a cooler LNG in such process.

While there are peak-shaving LNG plants that use single-cycle process, almost all

base-load LNG plants use either two- or three-cycle process. For example, the widely used

Propane Pre-Cooled Mixed Refrigerant (C3MR) process has two cycles; the first is the

Propane cycle that pre-cools feed gas and mixed-refrigerant, and the second is the mixed-

refrigerant cycle that cools and sub-cools up to the final product. Each of this cycle has their

own refrigerant, compressors and heat exchangers.

Among existing LNG processes that use three cycles are AP-XTM, Linde MFC, and the

ConocoPhillips Optimized Cascade (refer Table B.2 for all existing LNG processes). Fig. 3.2

shows the macro level view of refrigeration cycles within processes.

3 . O v e r v i e w o n E n e r g y , E x e r g y , a n d E c o n o m i c A n a l y s e s

32

Fig. 3.2. High level view of refrigeration cycles within processes [60].

This thesis discusses in detail the two-cycle process (C3MR) and three-cycle process

(AP-X and MR-X) only, through software simulations and their analyses that follow.

3 . O v e r v i e w o n E n e r g y , E x e r g y , a n d E c o n o m i c A n a l y s e s

33

3.2. Exergy Analysis

Exergy is the maximum theoretical useful work (shaft work or electrical work) obtainable

from a thermal system as it brought into thermodynamic equilibrium with the environment

while interacting with the environment only [61]. The environment is a large equilibrium

system, in which the state variables (T0, p0) and the chemical potential of the chemical

components contained in it remain constant, when, in a thermodynamic process, heat and

materials are exchanged between another system and the environment. It is important that

no chemical reactions can take place between the environmental chemical components. The

environment is free of irreversibilities and the exergy of its amounts to zero. In any thermal

system surroundings, the environment is part of the surroundings.

For an energy conversion system, the total exergy can be divided into four main

parts (neglecting nuclear, magnetic, electrical and surface tension effects): physical,

chemical, kinetic and potential exergy [62], that is ���� = ������ + � � + � � + ��� (3.1)

The rate of physical exergy ������ associated with a material stream (subscript j) is ��� = �� − � � − � − � (3.2)

where � and denote the rates of enthalpy and entropy, respectively. The subscript 0

denotes values at temperature and � of the environment. The physical exergy, where

similarly ��� = � ∙ �� (3.3)

= �� �ℎ − ℎ � − ∗ �� − � � �. At per unit kilogram of mass, the physical exergy of a working fluid can be further

split into its thermal (eT – due to its temperature) and mechanical (eM – due to its pressure)

exergy components: �� = � + (3.4)

�� = ��ℎ − ℎ ,�� − (� − � ,���=�����.+ ��ℎ ,� − ℎ , � − (� ,� − � , ��0=�����.

(3.5)

where X is the state where the temperature is at ambient temperature, and the pressure is at

specified temperature, or specifically [62]

3 . O v e r v i e w o n E n e r g y , E x e r g y , a n d E c o n o m i c A n a l y s e s

34

�� = ��ℎ − ℎ @�0,��� − �� − � @�0,��� �

+ ��ℎ @�0,�� − ℎ @�0,�0� − �� @�0,�� − � @�0,�0� � (3.6)

Exergy destruction � , and exergy loss � , are measures of inefficiencies

associated with the irreversible processes taking place in k-th system component. Exergy

analysis additionally often involves the calculation of measures of performance: exergy

destruction ratios, exergy loss ratio and exergetic efficiencies. As dimensionless criteria are

used for performance evaluations, an appropriate defined exergetic efficiency

unambiguously characterizes the performance of a system or system component from the

thermodynamic viewpoint. The rate of exergy destruction in the kth component is given by � , = � , − ��, − � , (3.7)

Here, � , represents the exergy loss in kth component, which is usually zero when the

component boundaries are at . For the overall system, � includes the exergy flow rates of

all streams leaving the system.

Exergy of product, �� is the desired result expressed in exergy terms, and achieved

by the system (the k-th component) being considered. �� is defined to be equal to the sum

of:

• all the exergy values to be accounted at the outlet (including the exergy of energy

streams generated in the component) plus

• all the exergy increases between inlet and outlet (i.e. the exergy additions to the

respective material streams) that are in accord with the purpose of the component.

Exergy of fuel, � is the exergetic resources expended to generate the exergy of the

product. Similarly, � is defined to be equal to:

• all the exergy values to be accounted at the inlet (including the exergy of energy

streams supplied to the component) plus

• all the exergy decreases between inlet and outlet (i.e. the exergy removals from the

respective material streams) minus

• all the exergy increases (between inlet and outlet) that are not in accord with the

purpose of the component.

The performance evaluation and the design optimization of thermal systems require

a proper definition of the exergetic efficiency and an appropriate costing approach for each

component of the system. The exergetic efficiency of a component and fuel are defined by

��

��

3 . O v e r v i e w o n E n e r g y , E x e r g y , a n d E c o n o m i c A n a l y s e s

35

considering the desired result produced by the component and the resources expended to

generate the result. � =�,��,� = 1 − �,� �,� (3.8)

In addition to � and � , , the thermodynamic evaluation of a system component is

based on the exergy destruction ratio , , which compares the exergy destruction in the

kth component with the fuel supplied to the overall system � ,��� :

y =� ,� ,���

(3.9)

This ratio expresses the percentage of the decrease in the overall system exergetic efficiency

due to the exergy destruction in the kth system component:

� =��,���� ,��� = 1−� , − � ,���� ,���

(3.10)

����� =��,���� ,���

(3.11)

� , is an absolute measure of the inefficiencies in the kth component, whereas � and ,

are relative measures of the same inefficiencies. In � the exergy destruction within a

component is related to the fuel for the same component, whereas in , the exergy

destruction in a component is related to the fuel for the overall system.

Alternatively, the component exergy destruction rate can be compared to the total exergy

destruction rate within the system, � ,��� , giving ratio

∗ =� ,� ,���

(3.12)

For the comparison of dissimilar components, the exergy destruction ratio may be

used. The exergy loss ratio can be defined similarly, by comparing the exergy loss to the

exergy of the fuel provided to the overall system:

=� ,���� ,���

(3.13)

Where and when necessary, chemical exergy is considered as referred to the

correlation in equation (3.1). The tabulated values for standard molar chemical exergy for

substances at normal environment conditions [61], help chemical exergy calculation. Such

calculation, for an ideal mixture of N ideal gasses,

�� ��������� ������ = � � ∗��= � � + � ∗ ∗� � ∗�

�= ln ( �) (3.14)

3 . O v e r v i e w o n E n e r g y , E x e r g y , a n d E c o n o m i c A n a l y s e s

36

and, for liquids (where coefficients of activity, �� is involved),

��������� = � � ∗��= � � + � ∗ ∗� � ∗�

�= ln (γ� ∗ �) (3.15)

The coefficients of activity, �� is represented as ‘gamma’ in Aspen Plus. Aspen Plus also

checked the enthalpy balances. Most components had zero error, while the maximum error

among components was in the scale of around 10-12 W. There are two models for the

tabulated standard chemical exergy values: The first one is suggested by [63] and uses pref =

1.019 atm, while the second one is suggested by [64] and uses pref = 1 atm [65]. The second

model was used because for this system pref was also taken as 1 atm.

3.2.1. Software requirements for simulation and exergy calculations

Aspen Plus [66] simulates LNG process system, components and streams. To prepare for

simulation, the following are the necessary inputs:

1. Components – Specifications: All involved materials must be identified here,

depending on the desired system to be simulated.

2. Methods – Specifications – Global – Property methods & options – Base method: SRK

(Soave-Redlich-Kwong). The Methods Assistant helps to suggest user to choose

property method based on user’s input [67].

3. Methods – Specifications – Global – Petroleum calculation options – Free-water

method: STEAMNBS. This method is selected due to its correlation with the SRK

method above.

4. Methods – Specifications – Global – Referenced – Reference methods: PENG-ROB,

CHAO-SEA, PSRK, REFPROP, RK-SOAVE, RKSWS, SRK, SR-POLAR. These are chosen

based on their specific approach towards natural gas, mixtures and non-ideal system

equation of state.

5. Property Sets – Properties: GAMMA. This property is chosen only when chemical

exergy is concern, where it is referred as the activity coefficient, �� [68]. It is strongly

recommended to remove this property should chemical exergy is not considered to

avoid multiple error messages in the Aspen Plus’ Control Panel.

To calculate the two split exergies, �� and � to have more accurate physical

exergy, the equation (3.6) could be executed using common formulas in Microsoft Excel

spreadsheet [69]. However, the spreadsheet requires from Aspen Plus unit-customized

results of each stream’s:

i. mass enthalpy (J/Kg), ℎ

ii. mass enthalpy (J/Kg), ℎ @�0,�� iii. mass enthalpy (J/Kg), ℎ @�0,�0

iv. mass entropy (J/Kg.K), �

3 . O v e r v i e w o n E n e r g y , E x e r g y , a n d E c o n o m i c A n a l y s e s

37

v. mass entropy (J/Kg.K), � @�0,�� vi. mass entropy (J/Kg.K), � @�0,�0

vii. mass flow (Kg/s), �

viii. ambient temperature (K),

Hence the calculated thermal and mechanical parts are in Watts units. As ��� = �� + � (3.16)

the addition of both results should produce the physical exergy. This could be validated

using physical exergy result produced from Fortran-Aspen subroutines calculations in

Megawatts (MW) term. Such MW term is useful for the next step calculation – exergies for

fuel and product of system and system components.

3 . O v e r v i e w o n E n e r g y , E x e r g y , a n d E c o n o m i c A n a l y s e s

38

3.3. Economic Analysis

An important aspect of analysing a process is economy. The success of a process design is

always, but not limited to, measured by monetary terms. To accomplish such, a complete

design of a process relies on the estimation of costs – of major things for example important

equipment, capital investment including fuel, operation and maintenance, and the

estimation of the final cost of the product. With proper cost estimations through the project

design, important basis for decision making can be achieved.

An initial practical economic analysis is conducted by means of engineering

economics with the help of several assumptions and estimations, in view of the type of the

process, technological and legal environments [29, 32]. For such analysis, it is vital to have

various factors such as the process flow, its flow rates, the process conditions, its

thermodynamic states of each material streams as well as the specifications of necessary

unit operations and equipment. To help the analysis with regards of the said factors, the

Total Revenue Requirement (TRR) method for estimating the Total Capital Investment (TCI)

is referred throughout the economics.

3.3.1. Estimation of Total Capital Investment (TCI)

The TCI of a plant is measured as a one-time cost, unlike cost of fuel or operating and

maintenance costs that are categorized as the continuous expenses, especially when the

plant’s design, construction and start-up phase is concerned. The criteria of the plant TCI are

shown in Table 3.1.

3 . O v e r v i e w o n E n e r g y , E x e r g y , a n d E c o n o m i c A n a l y s e s

39

Table 3.1. Elements of total capital investment [62].

I. Fixed Capital Investment (FCI)

A. Direct costs (DC)

1. Onsite costs (ONSC)

• Purchased equipment cost (PEC)

• PEC installation (45% of PEC)

• Piping (35% of PEC)

• Instrumentation and control (20 % of PEC)

• Electrical equipment and materials (20% of PEC)

2. Offsite costs (OFSC)

• Land (10% of PEC)

• Civil, structural and architectural work (50% of PEC)

• Service facilities (65% of PEC)

B. Indirect costs (IC)

1. Engineering and supervision (35% of PEC)

2. Construction costs (15% of DC)

3. Contingencies (10% of IC)

II. Other Outlays (OO)

A. Startup costs

B. Working capital

C. Cost of licensing, R&D

D. Allowance for funds used during construction (AFUD)

From the TCI criteria there are two main components

1. Fixed Capital Investment (FCI), and

2. Other Outlays (OO)

The FCI is applied for money needed for land and all necessary constructions,

purchases and installations of facilities and equipment of the plant. This is executed by

assuming the period for such FCI is a zero-time period of design, construction and

installation for the total cost of the plant. Properly the FCI is further divided into direct costs

and indirect costs categories. The former are the costs covering the main permanent

components, labor, materials and other means used throughout the construction,

fabrication and installation of the plant’s fixed facilities. The latter are the costs covering

other remaining expenses including temporary facilities needed for the project [29, 32].

Considering the information provided, the total capital investment is calculated as

OOICDCTCI ++= (3.17)

FCI

3 . O v e r v i e w o n E n e r g y , E x e r g y , a n d E c o n o m i c A n a l y s e s

40

Estimation of purchased equipment cost (PEC)

Estimating PEC is the first step in economic analysis. To do that, the construction materials

and the operating conditions of the process are specified. With as much quality data as

possible, the estimation and calculations implemented here will be better in accuracy and

reliability. Estimating the FCI is best when experience of such in the real industrial scene is

taking into account thus vendor’s quotation of any criterion of cost is valuable. The bigger

the scale of the project, the more important it is to consider advices, consultations or at least

price tags from vendors for costly equipment. Otherwise should manufacturer’s quotation

be unavailable due to confidentiality or whatnot, estimations taken from qualified and

seasoned professionals are preferred. These include calculations done from extensive cost

databases of reliable firms. Then, if the two most preferred options are not obtainable,

perhaps due to time or budget limitations, the purchased equipment cost (PEC) can be also

estimated via available charts in the literature [29].

The charts are usually built using the help of a high volume of data in cost and

design. When parameters necessary to get values from the charts are known, such as heat

exchanger’s heat transfer area or compressor’s power, it is easily possible to study the

charts for the desired cost of the equipment. The charts also allow effects of the equipment

characteristics to be taken into account, for example temperature and pressure. The

equipment base cost (CB) which obtainable from the charts is corrected by having these

effects as factors such as material factor (fm), temperature factor (fT) or pressure factor (fp)

[29]. Furthermore, bare module factors (fBM) can be included as well to the equipment final

module cost of, as shown below.

BMpTmdB fffffCPEC = (3.18)

Normally for stainless steel, fM = 2.5 and for electric motor driver, fD =1. For a compressor,

the base cost, referring to the year 2009 (CEPCI 394) is given by [70] � = exp {7.2223 + 0.80[ln(� )]} (3.19)

Typically, the charts available in the literature are costs log-plotted against the

equipment size, resulting in a straight line across the chart. The line’s slope, α, or the scaling

factor, is used for cost estimation of a specific equipment (CPE,Y) at a certain specific

parameter, for instance size (XY), in the presence of the purchase equipment cost of the same

equipment (CPE,W) at a different size (XW), as shown in the Equation (3.20).

α

W

YWPE,YPE,

X

XCC

= (3.20)

The change in reference year and size could fluctuate the scaling factor α. These two

parameters are searchable in the literature for various types of equipment and sizes. Unless

reliable data is available to be used, α may be assumed as 0.6 commonly known as the six-

tenths rule [29].

3 . O v e r v i e w o n E n e r g y , E x e r g y , a n d E c o n o m i c A n a l y s e s

41

Another key parameter for the economic analysis is the cost index. As the charts may

only have data referring for a specific period of time, all data gained from these charts have

to be carried to the same reference year (in which the year used as a basis for all economic

analysis). The Cost Index (CI), or referable as inflation indicator, is stated as

=

yearoriginal

yearreference

originalyearreference

CI

CI

CostCost (3.21)

The objective of this cost index is to correct the costs for the equipment, supplies, labor and

material for the estimation. The indexes typically referred to are Marshall and Swift (M&S)

Equipment Cost Index which is based on the construction expenses for numerous chemical

process industries, and Chemical Engineering Plant Cost Index (CEPCI), which is based on

the construction expenses for chemical plants [29].

Estimation of direct costs of FCI

There are two core divisions for direct costs as they are a component of FCI. Firstly, there is

the onsite costs (ONSC) and secondly the offsite costs (OFSC). While it is probable to

estimate these divisions through assessments based on the detailed flow diagrams of the

system, a factor method is applied for the rest of the calculations. Factor method is used for

respective component in terms of a percentage of the purchased equipment costs (% of

PEC). These percentage values given for each component are established as a product of an

experience from various plants in the chemical process industry [29].

Purchased equipment installation consists of the costs for the shipping and insurance

to bring equipment from the manufacturer to the site, the unloading, handling, ground

works, supports and labor and all other expenses that are significant to install fully the

equipment. The average value of 45 % may be used if any other information is not provided

[29].

Piping represents the expenses that relate to the material and labor costs required

for the construction of the whole piping grid in the plant. The piping costs 10% to 70% of

the purchased equipment costs.

The instrumentation and control rely heavily upon the intricacy of the purchased

equipment. The higher degree of automation and control are resulting more sophistication

in the equipment design, the more these instrumentation and control then costs. Typically, it

covers around 6-40% of the purchased equipment cost.

Electrical equipment and materials is defined as the expenses related to the

materials, installation and labor for distribution lines, power centers, substations,

emergency control supplies, switch gears and area lightning. Average value of 20% of the

PEC is applicable should the information for such section is unavailable.

3 . O v e r v i e w o n E n e r g y , E x e r g y , a n d E c o n o m i c A n a l y s e s

42

Land for the plant relies highly on the location of the plant. Typically, if need be

around 10% of the PEC is expected to cover such cost.

Service facilities are consisting of but not limited to the utilities (water, electricity,

steam and fuel) supply. For such expenses 65% is accounted from the PEC [29].

Estimation of indirect costs of FCI

The costs associated with supervision, construction and contingencies of the plant are

categorized as the indirect costs (IC).

Engineering and supervision expenses consist of the expense for the complete plant

design, drawings and other necessities such as supervision and inspection, scale models,

administration and procurement, travel and advisor fees. This engineering and supervision

may cost 25-75% of the PEC.

The construction fee itself which usually have the charges for movable facilities and

operations, equipment and tools, personnel home office on-site as well as insurances is

assumable at 15% of the PEC.

Contingencies accounts all uncertainties and risks in the actual costs calculations in

view of changing weather, challenges in transportation and sudden price changes, as well as

work stoppages. Contingencies take 5-20% of the FCI relying on the difficulty, size and

uniqueness of the plant [29].

Other outlays

The charges involved in the working funds, startup investment and allowance for funds used

during construction are referred as Other Outlays. These charges are second part of the total

capital investment (TCI).

Startup costs mostly involve the expenses of equipment, materials and overheads

which are funded only throughout the startup phase of the plant prior to its operation. The

startup costs of a thermal system are representable as a total of the unescalated costs such

as one month of permanent O&M charges, one month of unfixed operating costs calculated

at full load, one week of fuel at full load and 2% of the plant facilities investment.

Additionally, to these outlays, working capital is the capitals needed for the period of

the plant operation. It is essential for the operating costs prior to receiving the payment

from the product sale. Working capital includes the investment for [29]

a) raw materials, fuels and provisions carried in stock

b) finished goods in stock and semi-finished goods in the process of being

manufactured

3 . O v e r v i e w o n E n e r g y , E x e r g y , a n d E c o n o m i c A n a l y s e s

43

c) monies kept on hand for operating expenses, taxes and other current

obligations, and

d) accounts receivable

e) accounts payable

The allowance for funds used during construction (AFUDC) characterizes the time

value of the money through the construction, based on an interest rate equivalent to the

weighted cost of capital. Taking into account the construction period of a plant, portion of

the outlay is necessary to cover design studies, civil, acquisition, engineering effort and

setting up of equipment without having any income from the plant [29].

44

4. Processes of Liquefaction of

Natural Gas in Malaysia

Three LNG processes are discussed here in details: Propane Pre-cooled Mixed-Refrigerant

(C3MR), AP-XTM, and newly-developed MR-X.

The PRICO process however is not discussed in detail because it was used exclusively for

testing the simulation and mathematical models. The testing, executed prior to the three

former-mentioned LNG processes, was necessary to see the possibility (a) to simulate them

in suitable platforms, and (b) to analyse them using energy and exergy analyses especially

when involving the mixed-refrigerant cases. To regard such contributions and their

importance, the PRICO system is shown in Appendix B.1 (System testing using PRICO®

process).

4.1. Propane Pre-Cooled Mixed-Refrigerant (C3MR)

LNG Process

4.1.1. Principle of Operation

To facilitate the understanding of C3MR process, Fig. 4.1 is used for simplification and Fig.

B.4 is used for simulation. Mainly there are two main blocks (or cycles per Fig. 3.2). The first

block is the pre-cooling block, which uses Propane as its refrigerant (refer Legend). The

second block uses mixed-refrigerant to liquefy and successively sub-cool, in the central heat

exchanger known as the Main Cryogenic Heat Exchanger (MCHE) (or sometimes referred as

“coldbox”).

4 . P r o c e s s e s o f L i q u e f a c t i o n o f N a t u r a l G a s i n M a l a y s i a

45

Fig. 4.1. A general schematic of the C3MR process.

Typically, wet natural gas (NG) undergoes the pre-treatment process. In the pre-

treatment process, acid gas is firstly being removed from the NG. It then proceeds to

dehydration (water removal) and mercury removal process. After that, natural gas liquids

(NGL) are withdrawn from the NG. The NGL normally is sent for fractioning various by-

products such as Ethane and LPG among others. The pre-treatment and fractionation

processes for the C3MR process plant, however, are not analysed in detail in this thesis.

At (above) near ambient temperature, the cleaned NG from pre-treatment is fed into

the pre-cooling block at about 65 bar. The mixed-refrigerant coming from the coldbox is also

fed into this pre-cooling block, at about 48 bar. Both feeds’ temperatures produced from this

block are about -33°C. Throughout the cycles (including liquefaction), heat is removed using

air and/or seawater through inter- and after-coolers, to the environment. The pre-cooled

NG is fed into the coldbox to be liquefied and sub-cooled. The MR that received refrigerating

effect from the pre-cooling block is separated into gas and liquid phase through a phase

separator before going into the coldbox, at about 84 kg/s and 218 kg/s respectively, to

liquefy and subsequently sub-cool NG. A typical coldbox for a C3MR process-based LNG

plant is shown in Fig. B.2. The vaporized- and liquid-phase mixed-refrigerants recombine at

the exit (bottom catchment) of the coldbox. It is looped back afterwards into the pre-cooling

block driven by axial compressors [71] of several stages. The sub-cooled NG, which at this

state as liquefied, exits the coldbox. About temperature -162 C and slightly above ambient

pressure, through valve-throttling, the final LNG product, is ready for storage for shipment

purpose. Other targets may be implemented such as recycling flashed LNG into the system,

local electricity sources or further by-products.

4.1.2. Simulation and Energy Analysis

Aspen Plus [66] is chosen to simulate this process. For initialization, the Soave-Redlich-

Kwong (SRK) property method, recommended by the Aspen Property Method Assistant is

selected for this C3MR process simulation. Generally, SRK is recommended for gas

applications, as it can calculate the enthalpy and entropy values of the process streams, it is

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appropriate for a mixture of non-polar or mildly-polar, consistent even in the critical region,

and reasonable results can be found at all pressures and temperature [67].

Assumptions and specifications for the process, are as follow:

• The ambient temperature and pressure are set as 25 C (298. 15 K) and 1.01325

bar, respectively.

• Pressure drops throughout all heat exchangers are assumed to be 3% (in side

pipes), and no pressure loss for propane cycle for simplicity sake [22,72].

• It is also assumed no pressure loss within phase separators and mixers, and all

phase separations are specified to have no heating duty.

• The isentropic efficiencies are assumed to be 78% (propane compressors) and

75% (mixed-refrigerant compressors).

• The mechanical efficiencies are assumed to be 90% for all compressors.

• The discharge pressures for propane compressors are specified as 2.5 – 5.1 – 7.2

– 14.3 bar, and for mixed-refrigerant compressors are as 7.5 – 17.5 – 48.6 bar

[73].

• The initial values necessary to simulate NG feed, propane and mixed-refrigerant

mass flows are regulated according to a common LNG train capacity that used

C3MR process, which is at 4.5 mtpa. As such, the propane mass flow is specified

as 442.7 kg/s, the mixed-refrigerant mass flow as 301.8 kg/s, and the NG feed

mass flow is specified as 158.4 kg/s at 65 bar and 300 K (26.85 C)

• The composition for this process is as Table B.9 [74].

• The NG feed gas is pre-treated, pre-dried, and all unwanted materials including

acid gas are assumed to be removed before NG is being fed into the pre-cooling

block.

• The final LNG product is set to be slightly above ambient pressure, 1.2 bar

(113.15 K or -160 C) for further purpose such as storage.

Fig. B.4 shows the detailed C3MR process flowsheet used in the simulation using Aspen Plus.

For this particular simulation, full vapor NG at 65 bar and 26.85 C enters the first

pre-cooling heat exchanger, HEX1 and is pre-cooled by two-phased Propane at 14.19 C

through stream 19. The colder NG (17.85 C) leaves HEX1 and enters another three pre-

cooling heat exchangers HEX2, HEX3, and finally HEX4. Each of them has a similar pre-

cooling arrangement as the first; that is Propane that has been expanded pre-cools NG. NG

exits HEX4 at -33.15 C (57.55 bar) leaving the pre-cooling block and enters the liquefaction

and sub-cooling block through stream 43.

Each of these pre-cooling heat exchangers produces two-phase Propane (stream 20,

23, 26 and 29) which increases in temperature as the stage increases. They are separated

into vapor and liquid state using separators, where the former is used for compression, and

the latter is valve-expanded for pre-cooling. The vaporized Propane exiting from HEX4 is

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however recompressed. Such arrangement in series of stages, as the Propane’s pressure

being let down, reduces its temperature further and provides refrigeration to the NG.

Meanwhile in view of the refrigerant, the vaporized Propane is recompressed by

centrifugal compressor COMP1 from 1.3 bar (-34.65 C) to 2.5 bar (-7.73 C). It is then mixed

with separated Propane gas (stream 38 at 2.5 bar and -19.24 C) for the next series of

compression stages. Similar compression arrangement as the first one for the next COMP2,

COMP3 and COMP4. The Propane gas leaves COMP4 and enters its aftercooler at 14.3 bar

and 58.37 C. After expanded, the partially-condensed Propane pre-cools the first pre-cooling

heat exchanger HEX1 through stream 19 (7.2 bar).

For the mixed-refrigerant coming from the liquefaction and sub-cooling blocks into

the pre-cooling block, it enters the first pre-cooling HEX1 at 31.85 C and 48.2 bar. It then

continues to receive the refrigerating effects at all pre-cooling heat exchangers that follow.

The mixed-refrigerant leaves HEX4 at -33.15 C similar to that of NG and enters phase

separator SEPA4 for gas-liquid separation.

NG at gas state enters the liquefaction and sub-cooling block at MCE1. The MCE1 and

MCE2 heat exchangers with its valves and mixers represent the Main Cryogenic Heat

Exchanger (MCHE). The arrangement is as such so that the industrial MCHE flow and the

main reason behind the flow arrangement could be replicated as close as possible,

considering all the limitations that exist. To simulate the industrial MCHE as per its original

structure would be too complex and out of this thesis’ scope. Inside the MCHE, the NG is sent

through a tube circuit where it is liquefied (from vapor to liquid state, in this case at stream

44) and sub-cooled to -139.15 C at the exit of MCE2, by mixed-refrigerant.

The mixed-refrigerant (vapor at stream 8 and liquid at stream 14) coming from the

phase separator SEPA4 enter the MCHE simultaneously at -33.15 C and 42.63 bar. It is found

when the two-phase mixed-refrigerant is split into its individual vapor and liquid streams,

they formed new compositions of mixed-refrigerant. These mixed-refrigerant vapor stream

and mixed-refrigerant liquid stream are called Light Mixed-Refrigerant (LMR) and Heavy

Mixed-Refrigerant (MHR), respectively. Table B.10 [73] shows the newly formed

compositions.

Both streams climb through separate tube circuits where they are liquefied and sub-

cooled [75]. As they are relieved in pressure reducing their temperatures further, they

vaporize and flow downward on the shell side of the MCHE. This provides the liquefaction

(stream 13) and sub-cooling (stream 11) to the NG. Then, exiting the MCHE, the vaporized

mixed-refrigerant is sent to the compressors for recompression.

For the mixed-refrigerant coming from MCHE to be recompressed, a series of axial

compressors is used starting with COMP5 compressing the vaporized mixed-refrigerant

from 2.82 bar (-270.33 C). COMP6 and COMP7 with intercooler further compress it up to

48.6 bar (109.12 C), before being after-cooled and reinserted into the pre-cooling block.

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For the final product, LNG at stream 45 (54.15 bar, -139.15 C) is expanded through

VALVE7 and is sent to a high pressure phase separator. Due to the expansion any vaporized

NG is separated into stream 47. The vast expansion expands the LNG down to a slightly-

above ambient pressure (1.22 bar) and further reduces the LNG temperature to -160 C. Such

pressure and temperature is necessary for later processes such as LNG storage for shipping

or local transportation. A selected simulation results for this C3MR process is shown in

Table B.14.

In energetic analysis of the C3MR process, it is crucial to analyse the flow of energy

conversion process occurred in the particular system.

In the pre-treatment process for example, apart from bulk components such as

water and CO2, hydrogen sulphide, carbon dioxide and mercaptans are also removed from

the NG to ensure the NG that will be inserted into the liquefaction block has heavier

hydrocarbon components [76]. These heavy hydrocarbon components allow them to freeze

at very low temperature, condense, and consequently removed from the main gas (in the

liquefaction block). As a treated condensate, they can be shipped for commercial targets, or

being further refractionated into clean marketable products such as butane and propane.

The NG proceeds to the pre-cooling block as a clean and treated gas.

In the C3MR pre-cooling block, the propane cycle cools both MR and NG feeds at

several pressure levels or stages depending on the propane compressors arrangement.

Particularly for MLNG Dua plant, these are centrifugal type of compressors [60]. Different

designs of pre-cooling mean different number of pre-cooling stages, depending on the

environmental conditions. The number of compressors in C3MR system could be reduced,

compared to another system with the different type of refrigerant [77]. This is because MR

presents an extensive temperature glide up to the sub-cool temperature.

The vital system cooling using seawater has several important points. Awang [47]

stated about the early operation of heat exchangers in MLNG Satu and MLNG Dua facing

seawater tube failures. There were more than 150 documented cases due to pin hole

perforations initiated internally. Using seawater also required the Malaysian LNG producer

to obey the minimum standard for chlorination (level of 0.2 – 0.5 ppm) to keep the heat

exchangers free from marine fouling. While using seawater may hinder growth of barnacle

inside the heat exchangers, the higher level of chlorination proves to be challenging as

various materials of the exchangers reacts to such chlorination, which in this case important

to be accounted when selecting C3MR LNG heat exchangers, with further research.

Moreover, to avoid deposits and erosion-corrosion, all heat exchangers are advised to

operate at seawater velocities. As such, the seawater cooling systems require more

maintenance and should be designed specifically to the location’s environmental conditions.

Air cooling system may not be as cheap as the seawater cooling system in the long run, but it

is easier to operate and lesser to maintain.

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Meanwhile, in the main liquefaction refrigerant tubes, MR is preferred for its

efficiency. Reviewing the cooling curve of NG and its refrigerant is one of various methods to

determine the efficiencies of the liquefaction system and of the refrigerants applied. For

these two (NG and refrigerants) curves progressing from pre-cooling towards sub-cooling,

the closer the gap of these two curves, the more efficient is the system. As such, the

refrigerant curve should depict as much as possible the NG curve, and as close as possible.

By increasing the process efficiency, the work done by the heat exchangers would be

decreased [78]. As seen for cascade-type (such as of ConocoPhillips) and C3MR-type

liquefaction systems, the gap between the curves indicates the amount of exergy destruction

within the heat exchangers.

Fig. 4.2 shows the overall cooling curves for the C3MR process. Basically NG has a

cooling curve needs to be matched by the refrigerant system’s, and therefore their

compositions are important for the system to achieve high efficiency [79,80]. As such, the

smaller the gaps between NG curve and its refrigerant curve, the more efficient the process

will be. Here, it is notable the use of a single component (or pure) pre-cooling refrigerant

with a staged pressure let-down provides an easy to manage pre-cooling phase [75].

Nevertheless, the pressure levels permit unnecessary gaps which could be tackled through

for example the use of multi-component pre-cooling refrigerant. S. Madhavan [81] in his

presentation paper showed a similar curve for propane precooled mixed refrigerant cycle

against the natural gas’ (Fig. 4.3).

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Fig. 4.2. Overall cooling curves for the simulated C3MR process [73].

125

145

165

185

205

225

245

265

285

305

Te

mp

era

ture

(K

)

Heat Duty (MW)

NG cooling curve

Propane pre-cooling

Mixed-refrigerant

liquefaction and sub-

cooling

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Fig. 4.3. A similar curves for propane pre-cooled MR cycle versus the natural gas’ proposed

by Madhavan [81].

Meanwhile, the gaps located along the mixed-refrigerant curve could be improved

through transferring a partial of its distribution load to an additional consequent cycle (such

as per AP-XTM and MR-X processes) and/or readjustment of the multi-component

composition. In any case, the multi-component composition partial pressures and their

individual boiling temperatures are important to be understood first, for the sake of

advancements. Castillo et al. [77] however argues that when aspects like equipment size,

efficiency, final investment and other factors are considered, following the cooling curve

may not necessarily result in an optimal process.

The mixed-refrigerant used to liquefy and sub-cool LNG is a multi-component type of

refrigerant, which means several components are lumped together to be one working fluid.

The C3MR refrigerant components (Table B.9), are an unmixed mixture. It is not as water

and sugar which blends and becomes one entity, but take it simply as water and oil that is

together but unmixed. Of course, each of these non-azeotropic components affects each

other as pressure changes.

In this C3MR process condensation, the Propane substance of the mixed-refrigerant

starts the condensation. Propane, compared to other substances, has the highest boiling

temperature, as shown in Table B.11.

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This occurs when the (partial) pressure of the substance, �� is higher than its

saturated pressure, at a particular temperature. That is �� > ���� (4.1)

resulting condensation.

Taking stream 7 for example, at -33.15 C the saturation pressure ���� for pure

Propane is 1.467 bar. But, the partial pressure for Propane substance �������� (of the

mixed-refrigerant) at that particular temperature, can be calculated through �� = �� ∙ ���� (4.2)

where ���� is the total pressure of the mixed-refrigerant, and �� is the mole fraction of the

specific substance. Hence for Propane substance, �������� = 0.213 ∙ 42.63 bar = 9.08 bar (4.3)

therefore larger than the ����,������� (for pure Propane 1.467 bar at -33.15 C) or simply �������� > ����,�������(−33.15 �) (4.4)

which means the Propane substance condenses. Table B.12 shows the mixed-refrigerant

molar fraction and its partial pressure at stream 7 (as a sample for this process).

Note the effects affected the substances in the multi-component refrigerant affects

the compositions of the vapor and liquid phase of the partially condensed mixed-refrigerant.

However, as such changes occur over time, the overall mole fraction of a mixed-refrigerant

stream does not change. The partially condensed mixed-refrigerant may be separated into

LMR and HMR, where each has different composition, but when it recombines (even though

the LMR and HMR have undergone expansion and/or heat exchange) the overall mole

fraction is as before its separation.

Aspen Plus calculates entered user specifications, including the property method. It

will only simulate if all required specifications, and mass and properties balances are

correct. Error or warning messages will be shown to user based on the severity of

simulation issues. The software however cannot analyse whether a design is good, or bad.

Only the user could see and compare different designs, thus further execute (for example)

sensitivity or exergetic analysis to see which design is better.

As C3MR process involved cycles that use pure and mixed refrigerants, specially

designed diagrams are necessary [73]. These include their process and saturation curves (T-

s diagram), and cooling curves (T-Q diagram). The diagrams are important for analysing the

flow of a process and the relationships between involving properties. For such purpose,

additional Aspen Plus simulations are required based on their unique properties, and the

results are plotted using Microsoft Excel. Aspen Plus calculates enthalpies and entropies as ℎ = (�������� ,�ℎ�� , ,�) (4.5)

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� = (�������� ,�ℎ�� , ,�) (4.6)

4.1.3. Exergy Analysis

Based on the theory explained in the Overview on Energy, Exergy, and Economic Analyses

chapter, it is important to analyse the C3MR process exergetically through specific exergies.

To have the necessary enthalpies and entropies for thermal exergy calculation, each

stream of the C3MR process were resimulated at ambient temperature, and specified

pressure, �� . Also, prior to the resimulation, other data required by Aspen Plus were

reconciled from the previous result.

Similarly, to calculate mechanical exergy, the streams were reconciled and

resimulated but at ambient temperature, and ambient pressure, � . Table B.13 shows the

enthalpy and entropy values required for exergy calculation. Equations (3.3), (3.5) and (3.6)

enable the exergy components such as TE , M

E and PHE to be calculated (Table B.14).

These three components are particularly important to define the product and fuel for

conduction of the component-by-component exergy analysis as well as the total C3MR

system.

The C3MR process involved chemical exergy in its evaluation, as such the foundation

equation (3.1) is applied. The components need to be considered for chemical exergy here

are MIX4, a mixer, and SEPA4, a phase separator. These components are affected because of

the separation and mixing processes of the LMR and HMR. The LMR and the HMR, due to

their different compositions chemical reaction occurs and thus creates difference between

the chemical exergies of the streams related to the mentioned components.

Aspen Plus ‘gamma’ helps in calculating chemical exergy through equations (3.14)

and (3.15). The standard tabulated data mentioned in the chapter of ‘Overview on Energy,

Exergy, and Economic Analyses’ is also used, and values for chemical exergy are shown in

Table B.15 and Table B.16.

For other process components, chemical exergies which also occur remains similar,

and thus nullify each other when the components’ exergies of product and fuel are defined.

Consequently, the definitions of the exergy of fuel and the exergy of product for the C3MR

process are illustrated in Table B.17. The definitions are achieved through the splitting and

calculations of thermal and mechanical exergies. And hence the process exergy rate of

product and fuel is calculated, as well as the exergy destruction, exergetic efficiency and

exergy destruction ratio, as shown in Table B.18.

The total C3MR exergetic efficiency, ����� is 33.6%.

To improve the system exergetic efficiency, the components such as mixers, coolers

and throttling valves are excluded, it is imperative to see other more important components

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and their priorities. Only the components with highest exergy destruction values and lowest

exergetic efficiencies should be improved. For example, component MLHE1 has the highest

exergy destruction rate possesses very high exergetic efficiency. Similarly, component HEX1

has the lowest exergetic efficiency holds very low exergy destruction rate. Moreover, heat

exchangers type of component is difficult to improve significantly due to the manufacturer’s

confidentiality and complexity. The next possible components to be improved are COMP5

and COMP6 compressors. This is because they have relatively high exergy destruction (over

10 MW) as well as their exergetic efficiencies still have rooms for further improvement.

4.1.4. Economic Analysis

Economic analysis on the C3MR process is carried out based on the TRR (Total Revenue

Requirement) method [62]. To conduct a detailed economic analysis, it is important to

estimate the purchased equipment cost (PEC) as accurate as possible, because the

remaining cost contributors are calculated based on this value. As such, the most significant

assumptions and results are given for each component group:

• Multi-flow, low-temperature heat exchangers are considered as quite challenging due

to their unique design. Only a few companies in the world produce such heat

exchangers; therefore, the data required for the economic analysis of the heat

exchangers used for LNG processes are confidential and not available in the open

literature. Before cost estimation, it is necessary to assess the heat transfer area (A)

based on the operation conditions. Due to the complexity of calculations of the

overall heat transfer coefficient (U) [70,82,83] for each heat exchanger, it is

necessary as well to estimate such values. The heat transfer area A and the overall

heat transfer coefficient U are shown in Table C.2. Aspen Plus simulation of the

C3MR process provides ‘UA’ values at each heat exchangers, which in turn help

significantly to estimate the necessary parameters as accurate as possible. The PECs

for the heat exchangers are estimated per equation (3.20). The base cost is referred

from [83] and the scaling factor, α is assumed to be . . )n addition, the relation for cost index (CI) as per equation (3.21) is used for years between 2009 and 2012, due

to the base costs that are evaluated at the year 2009 with CEPCI of 521.9. In 2012,

the CEPCI is 584.6. The bare module factor (fBM) is assumed to be equal to 3. These

parameters help to produce the PECs shown in Table C.3.

• Turbomachineries is the second most important group of equipment within the

C3MR process. Centrifugal compressors are selected. All compressors are driven by

electric motors. For the cost calculations of these equipment-items, the net required

power (PC) is used as a sizing factor [70]. Using equations (3.18) and (3.19), the

purchase equipment cost correlation is expressed as

BMD CFFPEC = (4.7)

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and a CEPCI of 584.6 in 2012 is applied. The PC and their PEC values are illustrated in

Table C.6 and Table C.7.

• Remaining equipment items – interstage- and aftercoolers, separators, mixers and

throttling valves.

Interstage coolers are considered to be shell-and-tube heat exchangers. The UA

values of the coolers were obtained from the AspenPlus simulation software, and the

overall heat transfer coefficients (U) were taken from the literature. Derived from

the estimating charts, such following equation � = �{11.0545− 0.9228[ln(�)] + 0.09861[ln (�)] } (4.8)

is applied to its purchased equipment cost. One may avoid the error of reading the

estimating charts by using the above correlation.

When the detailed economic analysis is conducted, it is also essential to compare the results

of the economic analysis with the real data of the industry. The C3MR process that is

analysed in this work is assumed to be in Malaysia. Therefore, the initial unit costs such as

electricity, labor cost or the cost of natural gas are selected according to this criterion. In

addition, the offsite costs which are mainly consist of the cost of land, civil, structural and

architectural work and service facilities, are not considered during the during the analysis

as no reliable information is available for the costs related to the property of the plant. Table

C.17 demonstrates the data set of some LNG plants that are still in operation. Consistent

with this table, the unit cost is assessed based on the total capital investment. The unit cost

of LNG is approached as follow: � �� =�����

(4.9)

� �� =. � 9� $

. ( 6) ��� = 0.344 $ bn/mtpa.

When the unit cost with a different approach is compared with the real existing

plants operating in Malaysia, it can be seen that the measured value is in the range of the

actual plant values at the capacity of 4.5 mtpa.

Technical and economic data correlated to the C3MR production facilities (the initial

start-up time is between 1994 and 2007) is illustrated in Table C.18. In this table, only LNG

plants that operate in warm climatic conditions are shown. It can be seen that the capital

unit costs of C3MR plants has decreased by a factor of . − . during the last years. For the analysis it is safely assumed that the minimum capital unit cost should be close to . − 0.30 $ bn/MTA for one- and two-train LNG plants and 0.37 – 0.40 $ bn/MTA for LNG plants

with three or more trains.

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4.2. AP-XTM LNG Process

A more suitable LNG process is preferable as the popular C3MR process, even with the use

of multiple drivers and large motors, could feasibly produce up to 5 mtpa [84]. This is true,

for example, for centrifugal MR and propane compressors (of the C3MR process)

approaching single casing flow limits at current world production scale. Furthermore,

customers continuously demand a lower unit cost.

The AP-XTM process was analyzed in this thesis with several objectives. With bigger

capacity production and design than the previous two processes (PRICO and C3MR), as well

as different in climatic conditions and compositions, more interesting results and

perspectives could be achieved by such analyses. Also, the AP-X process is confirmed to

influence the latest Malaysian AP-NTM process FLNG plant. Owned by PETRONAS, the latter

process plant is in construction phase ready to be commissioned in 2015 and 2018 as the

world’s first floating LNG plant offshore state of Sarawak, Malaysia [85]. In this thesis the

suggestions for improving AP-XTM process given in [50] are implemented.

4.2.1. Principle of Operation

The principle of operation for the AP-X process is almost similar to the C3MR process. It

differs mainly on 1) the subcooling section, and 2) the production capacity. Other

parameters such as the refrigerant compositions, drivers, and climatic conditions may also

differ, but they are less important compared to the two mentioned points as far as both

operations are concerned.

Fig. B.5 shows a simplified schematic of the AP-X process. The C3MR process is a

part of the AP-X process. Both, the C3MR and the AP-X processes use a multicomponent

mixture of the refrigerants as a working fluid. This mixture can contain hydrocarbons such

as methane, ethane, propane, butane, and other light hydrocarbons, and also nitrogen

[26,50].

There are mainly three main blocks (Fig. B.5). The first block is the pre-cooling

block, which uses Propane as its refrigerant (refer Legend). The second block uses mixed-

refrigerant to liquefy in large heat exchanger known as the Main Cryogenic Heat Exchanger

(MCHE) (“coldbox”). The third block is the sub-cooling block, which uses nitrogen as its sole

refrigerant.

Typically, wet natural gas (NG) undergoes the pre-treatment process. In the pre-

treatment process, acid gas is firstly being removed from the NG. It then proceeds to

dehydration (water removal) and mercury removal process. After that, natural gas liquids

(NGL) are withdrawn from the NG. The NGL normally is sent for fractioning various by-

products such as Ethane and LPG among others. The pre-treatment and fractionation

processes for the AP-X process plant, however, are not analyzed in detail in this thesis.

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At (above) near ambient temperature, the cleaned NG from pre-treatment is fed into

the pre-cooling block at about 65 bar. The mixed-refrigerant coming from the coldbox is also

fed into this pre-cooling block, at about 38 bar.

Throughout the cycles (including liquefaction), heat is removed using air and/or

seawater through inter- and after-coolers, to the environment.

The pre-cooled NG is fed into the coldbox to be liquefied and sub-cooled. The MR

that received refrigerating effect from the pre-cooling block is separated into gas and liquid

phase through a phase separator before going into coldbox to liquefy and subsequently sub-

cool NG. As AP-XTM process is a C3MR process-based, a rather similar typical coldbox for

such LNG process plant is shown in Fig. B.2. This however may depend on the larger size of

AP-X process – almost doubled the capacity of the C3MR’s.

The vaporized- and liquid-phase mixed-refrigerants recombine at the exit (bottom

catchment) of the coldbox. It is looped back afterwards into the pre-cooling block driven by

axial compressors [71] of several stages.

The final liquefied NG in the MCHE exits the coldbox, and enters the nitrogen

expander loop at -111 C to be further sub-cooled up to -166 C. After a valve-throttling, the

LNG is ready to be stored for shipping and other purposes including recycling flashed LNG

into system, local electricity sources or further by-products.

4.2.2. Simulation and Energy Analysis

Aspen Plus [66] is used to simulate this process. For initialization, the Soave-Redlich-Kwong

(SRK) property method, recommended by the Aspen Property Method Assistant is selected

for this AP-X process simulation. Generally, SRK is recommended for gas applications, as it

can calculate the enthalpy and entropy values of the process streams, is appropriate for a

mixture of non-polar or mildly-polar, is consistent even in the critical region, and reasonable

results can be found at all pressures and temperature [67].

Assumptions and specifications for the process, are as follow:

• The mass flow rate of the natural gas feed is 224.23 kg/s, in accordance with the 7.8

mtpa production.

• The spiral-wound Main Cryogenic Heat Exchanger (MCHE) or coldbox is modeled

using two multi-flow heat exchangers and a refrigerant catchment.

• Scrubbing and fractionation effects for the mixed refrigerant working fluid are

neglected.

• In general, the pressure drops across the process are 0 bar.

• The volumetric flow of the nitrogen expander outlet is 35% of the volumetric flow of

the inlet of separator for the mixed refrigerant [50].

• The isentropic efficiency of the nitrogen expander is 85% [50].

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• The mass flow rate of stream 11 is equal to zero.

• Based on Qatari conditions, the ambient temperature and pressure are set as 43 C

and 1.01325 bar, respectively.

• The composition for this process is as per Table B.19.

• The NG feed gas is pre-treated, pre-dried, and all unwanted materials including acid

gas are assumed to be removed before NG is being fed into the pre-cooling block.

• The final LNG product is set to be slightly above ambient pressure for further

purpose such as storage.

Fig. B.6 shows the detailed AP-X process flowsheet used in the Aspen Plus simulation.

The AP-X process can be divided into four circuits in general: a) natural gas (streams

1 through 10), b) mixed refrigerants (streams 100 through 120), c) nitrogen (streams 1000

through 1006), and d) propane (streams 500 through 521).

Natural gas at state 1 passes through four in-series propane loop heat exchangers

and exits as stream 5. Pre-cooled natural gas at state 5 goes through (a) the COLDBOX that

consists of the two multi-flow heat exchangers HEX0 and HEX1, and mixer MIX1, and (b) the

HEX2. Then it is throttled down (V3) to the required pressure (state 8) and at state 9 it is

transferred for storage and further transportation.

The mixed-refrigerants at state 100 is separated in the separator SEP2 to its liquid

and gas phases. Both streams, 101 and 102, enter HEX0 and exit it as stream 112 and stream

103, respectively. Stream 112 continues to HEX1 and after that (stream 113) is split into

stream 114 and stream 118. Stream 114 through V1 returns to the coldbox (stream 115),

and after passing through HEX1 and HEX0 (stream 117), it is mixed with streams 105 and

120 (in MIX1 and MIX 2) to obtain stream 107. Stream 107 then is compressed in low- and

high-pressure compressors (LPMR and MPMR). During the interstage cooling process

between LPMR and MPMR, liquid may have formed and condensed under the multi-

components refrigerant partial pressure (stream 109). Stream 109 passes the pump (P1) to

reach the pressure that is equal to the pressure of stream 108. Both streams, 108 and 109,

are mixed, and the resulting stream 111 is pre-cooled in the four ‘in-series’ propane loop

heat exchangers to reach the condition of state 100.

The propane refrigerant at state 500 takes the heat from both MR and NG streams

from PHX1, the first precooling heat exchanger of four ‘in-series’, to state 501 into PSEP1 to

separate a) gas phase at stream 519 to combine with compressed stream 515 for the fourth

compression in COMP4, and b) liquid phase to take the heat again from both NG and MR

streams (2 and 97, respectively) and exits as stream 504. The PSEP2 separates again stream

504 into a) gas phase stream 502 so to combine with compressed stream 513 for the third

compression in COMP3, and b) liquid phase stream 505 to be valve-throttled at state 506 to

take the heat again from both NG and MR streams (3 and 98, respectively) and exits as

stream 507 for another phase separation. PSEP3 separates the phases to a) gas phase

stream 521 to combine with compressed stream 511 for the second compression in COMP2,

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and b) liquid phase stream 508 to be throttled to state 509 to a final pre-cooling heat

exchange at PHX4 before exit to stream 510 for the first compression in COMP1.

The discussion of the nitrogen cycle starts with stream 1001 that passes the LPN2,

MPN2 and HPN2 (multi-stage compressor with corresponding interstage coolers), followed

by the after-cooler AC, and the multi-flow heat exchanger HEX3. At state 1004 nitrogen is

expanded in expander EXP1 (stream 1005) and through HEX2 (stream 1000) to return to

HEX3.

The thermodynamic data for the material streams at real operating conditions is

illustrated in Table B.20 for all type of streams (1-11, 97-120, 500-521, and 1000-1005).

The process to produce 7.8 mtpa in a single train uses the multi-component refrigerant in

the main liquefaction section. Each stream carries the partial pressure of every component

that depends on their boiling temperature, saturation and stream pressures per molar

fraction. Table B.21 shows the mole flow rate of the mixed refrigerant at the different phases

for each substance.

Ideally the final exit pressure should at or slightly below the dew-point pressure of

the nitrogen - a temperature cold enough to affect the cooling of the LNG to the desired

temperature.

The coefficient of performance (COP) of the AP-X process, ��� =� − �����

(4.10)

is 0.41.

In the energetic analysis of AP-XTM process, it is crucial to analyse the flow of energy

conversion process occurred in the particular system. Almost similar to the C3MR process,

in the pre-treatment process for example, apart from bulk components such as water and

CO2, hydrogen sulphide, carbon dioxide and mercaptans are also removed from the NG to

ensure the NG that will be inserted into the liquefaction block has heavier hydrocarbon

components [76]. These heavy hydrocarbon components allow them to freeze at very low

temperature, condense, and consequently removed from the main gas (in the liquefaction

block). As a treated condensate, they can be shipped for commercial targets, or being further

refractionated into clean marketable products such as butane and propane. The NG

proceeds to the pre-cooling block as a clean and treated gas.

In the AP-X pre-cooling block, the propane cycle cools both MR and NG feeds at

several pressure levels or stages depending on the propane compressors arrangement.

Depending on the environmental conditions, different designs of pre-cooling mean different

number of pre-cooling stages. The power net required by the AP-X process components is

given in Table B.24. The expander power is shown to have -66.93 MW. The ‘minus’

represents the power produced by the expander.

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In view of ∆ ����ℎ of liquefying multi-flow heat exchangers namely HEX0, HEX1,

HEX2, and HEX3, their hot-cold temperature against cumulative duty (MW) are shown in

Fig. B.7, Fig. B.8, Fig. B.9, and Fig. B.10, respectively.

While small-scale LNG plant may show insignificant inefficiency due to the sole-

Nitrogen usage, mid- and larger-size LNG plants are recommended to couple nitrogen as

their sub-cooling refrigerant with another type of refrigerant(s) for pre- and main

liquefaction of LNG. [43] claimed that the AP-X process in Qatar has the best refrigerants

selection and arrangement due to having Propane (sub-cooling section), mixed-refrigerant

(main liquefaction) and Nitrogen (sub-cooling section) as their refrigerants. Nevertheless,

the latest AP-NTM technology for the first FLNG in the world, located in Malaysia, uses only

nitrogen as their primary refrigerant. Further study on the latter technology is highly

recommended to see the nitrogen’s capability, as well as its particular system efficiency.

The pre-cooling block for AP-X process uses pure propane as the refrigerant. The selection

of propane as the refrigerant, as explained earlier is due to the advancements made by

integrating both Single MR process and cascade-type process. In view of pure propane pre-

cooling, Castillo and Dorao [77] highlighted that in some cases, the application of pre-

cooling refrigerant seems to be related to patentability issues of the technology more than

due to engineering or efficiency factors. The authors mentioned that it was still not well

understood the advantages and disadvantages of considering an MR or pure refrigerant for

pre-cooling. The propane refrigerant however when being compared between C1, C2, mix1

(50% C1, 50% C2), and mix2 (25% C1, 25% C2, 50% C3), showed it consumed the least power.

And, while low power consumption could be a plus factor for propane preference, one must

aware about the limitation it has for lower temperature reaching capability at -36°C [86],

thus the possibility of vacuuming if it is lower than such.

While the C3MR process uses its similar loop from the MCHE for subcooling, the AP-X

process uses a nitrogen expander cycle to subcool the NG. As such, the cycle makes

significantly bigger capacity achievable [84]. This is possible through the reduction of both

propane and mixed-refrigerant flow. The MR volumetric flow at the low-pressure

compressor suction is around 60% of that used by the C3MR process for the same capacity.

The propane mass flow is around 80% of that used by the C3MR process. Therefore, the AP-

X subcooling process allows train capacities to achieve eight mtpa – almost doubled the

production of C3MR process. This is done without repeating compression equipment, in the

existing compressor frame sizes and using a single spool-wound MCHE, produced in a single

train.

4.2.3. Exergy Analysis

The definitions of the exergy of fuel and the exergy of the product for the AP-X process are

illustrated in Table B.22. The definitions are achieved through the splitting and calculations

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of thermal and mechanical exergies. Consequently, the process exergy rate of product and

fuel is calculated, as well as the exergy destruction, exergetic efficiency and exergy

destruction ratio, as shown in Table B.23 and is described simpler as per Fig. 4.4 and Fig. 4.5

for selected components.

Fig. 4.4. Exergy destruction (MW) and exergy destruction ratio (%) for selected components

of the AP-X process.

Referring to Fig. 4.4, the pre-cooling heat exchangers (PHX1-4) produces

insignificant exergy destruction when compared to the MCHE. Nevertheless, in the same

precooling section, the final COMP4 is the second highest component in destroying exergy.

Two compressors ‘in-series’ (LMPR and MPMR) carry the highest exergy destruction, when

combined, if compared to the HEX0 exergy destruction. The HEX0 exergy destruction is

understandably high due to maximum mixture of refrigerants and natural gas. The mixtures

in the MCHE therefore have high chemical potentials and reactions. In structural term, its

enormous column and long tubes which easily possess pressure drops, frictions and other

factors contribute as well to the highest-single-component exergy destruction.

Meanwhile Fig. 4.5 demonstrates the exergetic efficiency for selected components of

the AP-X process. Except for the PHX1, the PHX2-4 is seen to have the lowest efficiency even

though it belongs to the less-priority pre-cooling heat exchangers group of low exergy

COMP1

2%COMP2

3% COMP3

5%

COMP4

15%

EXP1

5%

HEX0

23%

HEX1

1%

HEX2

1%

HPN2

2%

LPMR

9%

LPN2

2%

MPMR

15%

MPN2

5%

P1

0%

PHX1

0%

PHX2

1% PHX3

3%

PHX4

7%

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destruction. This opens to a reevaluation of the PHXs, as well as for COMP4 that has only

around 30% exergetic efficiency but significant exergy destruction.

Fig. 4.5. Exergetic efficiency for selected components of the AP-X process.

Both low-pressure compressors for the MR circuit and the nitrogen expander cycle

achieve more than 70% of exergetic efficiency, a good mark for high LNG capacity

production. This is also true for both pre-cooling compressors COMP1 and COMP2. The

pump P1 while shows high efficiency produces only less than 500 kW for this 7.8 million tpa

plant.

On another note, between liquefaction and subcooling, supplemental refrigeration is

provided to heat exchanger HEX3 by a small steam 119 of the mixed refrigerant - to reduce

the irreversibility in the process by causing the cooling curves heat exchanger HEX3 to be

more closely aligned.

The implemented suggestions for improving AP-X process [50] showed only

reducing in the specific values of exergy destruction. Since the mass flow rate the working

fluid increased significantly, the total values became very large. As the result, the exergetic

efficiency of the overall system is 6%.

4.2.4. Economic Analysis

DiNapoli et al. [87] compared direct costs of five LNG plants (QatarGas, Nigeria LNG, Atlantic

LNG, RasGas, and Oman LNG) in detailed. The database is among the first available

econometric analysis of LNG plant capital cost development. In comparison to the cost

study, Table 4.1 and Fig. 4.6 show the estimation of the Purchased Equipment Cost and

direct costs for the AP-X process, respectively. The PEC is executed using equations (3.20)

and (3.21) where the previous C3MR process values are taken as its reference for

0

10

20

30

40

50

60

70

80

90

100

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benchmark assessment. The PHX1-4 have to take the maximum size data available, that is of

the MCHE, due to no publicly available data for them. It should be noted that since the MCHE

carries the highest cost in all LNG liquefaction components, it is therefore the PHX1-4 should

actually costs lower and therefore reducing the total direct cost.

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Table 4.1. Estimation on Purchased Equipment Cost for selected AP-X process equipments

No. Type Name Variable

X Value

XY Unit

Exponent α

C3MR ref. Value

XW Unit

C3MR ref. CPE,W (106 $)

AP-X

CPE,Y (106 $)

1 Expander EXP1 Power 66.93 MW 0.6 COMP6 36.11 MW 35.43 51.31

2 Heat Exchanger HEX0 & HEX1 Area 20,0001 m2 0.58 MLHE1&2 8,718 m2 92.1 149.08

3 Heat Exchanger HEX2 Area 2,0982 m2 0.58 MLHE 2 329 m2 12.2 35.73

4 Heat Exchanger HEX3 Area 10,0583 m2 0.58 MLHE 1 8,389 m2 79.9 88.76

5 Heat Exchanger PHX1-4 Area 20,0004 m2 0.58 MLHE1&2 8,718 m2 92.1 149.08

9 Compressor LPMR Power 504.36 MW 0.6 COMP7 42.87 MW 40.64 178.36

10 Compressor MPMR Power 495.35 MW 0.6 COMP6 36.11 MW 35.43 170.51

11 Compressor LPN2 Power 112.23 MW 0.6 COMP4 21.43 MW 18.73 50.58

12 Compressor MPN2 Power 173.00 MW 0.6 COMP4 21.43 MW 18.73 65.58

13 Compressor HPN2 Power 94.35 MW 0.6 COMP4 21.43 MW 18.73 45.58

14 Compressor COMP1 Power 100.70 MW 0.6 COMP1 3.38 MW 4.28 32.80

15 Compressor COMP2 Power 198.48 MW 0.6 COMP2 10.66 MW 10.72 61.97

16 Compressor COMP3 Power 203.82 MW 0.6 COMP3 7.77 MW 8.32 59.08

17 Compressor COMP4 Power 308.88 MW 0.6 COMP4 21.42 MW 18.73 92.88

18 Pump P1 Power 0.45 MW 0.6 COMP1 3.38 MW 4.28 1.28

TOTAL 1232.57

1 Based on Suprapto, 2007 [109]. 2 The usage of coil-wound heat exchanger (CWHE) [84]. 3 The usage of brazed aluminium type of plate fin heat exchanger (PFHE) [84]. 4 No data available thus is assumed similar to of the MCHE (HEX0 and HEX1).

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The total direct costs for the process, 4.252 $ bn for the 7.8 mtpa, or 0.545 $ bn per

tons per annum according to the DiNapoli’s reference shows the AP-X process (7.8 mtpa)

lies slightly higher than the QatarGas process plant (4.5 mtpa). Take into consideration that

the report from [87] is compiled by a consulting firm that has limited public access. The

costs for some specific precooling heat exchangers (PHX1-4) are represented by the MCHE

(HEX0 and HEX1) cost due to the unavailability of data. The estimation of such is done

according to the [62,73,83] references. The scaling factor, α is assumed to be 0.58. In

addition, the relation for cost index (CI) as per equation (3.21) is used for years between

2009 and 2012, due to the base costs that are evaluated at the year 2009 with CEPCI of

521.9. In 2012, the CEPCI is 584.6.

Fig. 4.6. The estimation of direct costs for AP-X process

The cost of building LNG plant varies over time. In the 1980s, the building cost

reported by the Economist [88] was 0.350 $ bn per year, while in 2000 it was 0.200 $ bn per

tpa. This is due to technological advancement, high competition among newly emerged and

seasoned LNG key players among other factors. However, in 2012, the costs could go further

as high as 1 $ bn per tpa partly because of the steel price hike as per the Economist report.

The AP-X process at this level is positioned at the near end of the range that is

approximately 0.700 $ bn per tpa. This value is the Total Fixed-Capital Investment (FCI)

which includes the total direct and indirect costs. This value also is consistent with a

consultant firm report [89] on the cost for the liquefaction part in the LNG chain at around 1

$ bn.

2711,65

1232,57

554,66431,40

246,51 246,51

1540,71

123,26

616,28801,17

0

500

1000

1500

2000

2500

3000

Est

ima

tio

n o

f C

ost

s

Elements of Direct Costs (Onsite + Offsite Costs)

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It was common to accept that this was a “learning curve” effect, and it would carry

on into the future. This is as recent as 2003. However, this insight of steadily falling costs for

LNG plant construction has been avoided in the past couple of years [90]. The construction

charges of greenfield LNG projects started to rise steeply from 2004 afterward and has

amplified from about 0.400 $ mil per tpa of capacity to 1 $ bn per tpa of capacity in 2008.

The following describe the main reasons for the sharply increased costs in the LNG industry

[91]:

• US dollar devaluation.

• High raw material prices due to surge in demand for such.

• Low availability of EPC contractors as a result of the extraordinary high level of

ongoing petroleum projects globally.

• Lack of skilled and experienced workforce in LNG industry.

The global financial crisis between 2007–2008 affected a general decline in raw

material and equipment prices, which somewhat lessened the construction cost of LNG

plants. However, by the year 2012, this was more than a counterbalance to the increasing

demand for materials and labor for the LNG market.

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4.3. MR-X LNG Process

To understand the newly design MR-X concept, the integration between the LNG processes

(C3MR, AP-X and DMR) are necessary to be described which leads to the creation of the MR-

X process. The new process is designed to replace, or at the least counterbalance, any

deficiency of the previous LNG systems. Therefore, it is also necessary to see the pros and

cons of them in general.

The AP-X process is known so far capable to produce the biggest capacity per train

(up to 10 mtpa [51]). While this may mainly be achieved by the nitrogen subcooling cycle,

the AP-X process still precools using pure Propane. There are pros and cons about it, but in

terms of cumulative cooling curves which can represent the system efficiency, it is rather

disadvantageous. A similar type of inefficiency is found in the previous systems; the C3MR

and the cascade systems; due to the application of pure refrigerant(s). This inefficiency is

seen using the cooling curves. Fig. 4.7 and Fig. 4.8 show the respective curves for both

mentioned process.

Fig. 4.7. A general cooling curve for cascade type of LNG process. The smoother curve is the

NG-LNG curve and below it is the refrigerants curve [92].

Work done on a liquefaction system is based on the function of heat duty ( �) and

temperature of the process. The smaller the work done for a similar LNG production, the

more efficient the process is. For example, for a Carnot cycle running between a certain low

temperature and the ambient temperature, hence � = �(��� − ����� ) (4.11)

Given a differential area by

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� = �( ��� − �� ) (4.12)

therefore integrating it produces � =���� (4.13)

As a result, the larger area between ��� and the curve, the more work the process

has to carry out.

Nevertheless, it is impossible to have refrigerants to exactly follow the NG-LNG

cooling curve in the real process. As ideally as possible, the refrigerants curve needs to

follow the NG-LNG curve since the closer it is, the more efficient the system is. Fig. 4.8 below

shows the existing C3MR process’ cooling curves.

Fig. 4.8. Cooling curves for C3MR process [92].

The mixed refrigerant, although more complicated than the pure ones, has the

option to manipulate its compositions (in the multi-components working fluid). This

adjustment enables it to match the NG-LNG curves as best as possible, therefore, increases

the system efficiency. The C3MR process has achieved so far due to such, except for its

precooling section (which the process owner may has proper and fair justification for it).

Apart from system inefficiency, having pure Propane as refrigerant means the

respective refrigeration system has to endure the limited low temperature. The precooling

cycle that uses Propane is not recommended to perform below -35 C [92]. Executing

Propane to refrigerate below such temperature will cause vacuum suction in the propane

compressor. This is more significant as typically the C3MR precooling section is the area

where most of the power reduction occurs.

The C3MR process is the most installed system globally, thus to have an improved

system (by removing the above disadvantages), a practical approach is essential. The

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improvement should increase the system efficiency and flexibility while maintaining

worldwide accepted and success-proven infrastructure. Malaysia is suitable to have such

newly improved system as the C3MR system infrastructure exists for at least three plants,

and train expansions have been progressing well [44].

The AP-X process is an extended version of the C3MR process. It still has similar

infrastructure of the latter system, thus by improving the AP-X process particularly on the

precooling section will both have the increased efficiency and flexibility, as well as the AP-

X’s advantage on largest production per train capability.

From here, scientific research is vital in finding the best pre-cooling system available

publicly, and reliable, for a new LNG process. Shell, has claimed that its Dual-Mixed

Refrigerant (DMR) process overcame the C3MR propane limitations [71]. Venkatarathnam

[93] also showed that the DMR exergetic efficiency is higher than the C3MR’s. In the same

analysis, a detailed comparison also was done between two DMR processes; DMR-1 and

DMR-2. Another variety is shown in Table 4.2 below where the author, in conclusion,

showed the DMR-1 is better than the DMR-2 overall. The fraction of Propane (of the mixed-

refrigerant) is very different between the two systems, which has influenced the process

performance.

Table 4.2. The optimum composition of precooling refrigerants for DMR process analysed by

[93].

Component

mol (%)

DMR-1 DMR-2

Ethane, C2H6 45.47 24.82

Propane, C3H8 4.94 64.16

nButane, C4H10-1 49.59 11.03

The DMR process was developed for Arctic climatic operation conditions: low

average annual temperature but relatively high-temperature differences during the year.

For such operating conditions, the precooling load varies significantly during the year. The

DMR process was built for a single train only so far referring to Sakhalin project in Russia.

Therefore, referring comments from [92], and based on the reasons mentioned

above, the DMR-1 is selected to replace the precooling Propane in the AP-X process.

Although it may be more intricate when compared to single-component refrigerant system,

the benefit is that two mixed-refrigerant systems (precooling using DMR-1 and central

liquefaction using MR of C3MR process) allow more flexibility in selecting the precooling

temperature. Furthermore, this permits a more favorable choice of drivers and compressors

for certain feeds or ambient conditions.

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Due to the arctic condition design initially for the DMR process, by employing it in

the AP-X pre-cooling process, the design condition is reset to the Equator temperature (in

this case referring to Qatar’s 40 C ambient temperature). This energetic conversion from a

freezing condition to a very hot climate enables LNG producers that have existing

infrastructures, such as Malaysia, to have improvement and expansion more convenient.

Furthermore, with regards to the many C3MR process-based plants operating for decades

compared to the Sakhalin DMR process that operates recently, designing the improved LNG

process using a C3MR-based process is seen as more preferable. The AP-X nitrogen

subcooling section that uses C3MR-based process also put more weight here.

Therefore, the author [56,94] redesign an improved LNG process that combines all said

advantages as well as removes the adverse effects. Such newly LNG system is named MR-X5.

The MR-X process is hoped to be a favorite due to its adjustability on mixed refrigerant

composition at seasonal basis as the temperature changes to have the best efficiency.

“As an alternative to propane, precooling can also be accomplished by using a separate MR

loop in a variant known as a dual mixed refrigerant (DMR) process. While operating an MR

precooling system may be more complicated than the simple single component propane

system, the advantage is that two MR systems in series allow more flexibility in selecting the

precooling temperature. With certain feeds or ambient conditions, this may allow a more

optimum selection of compressors and drivers.”

Bronfenbrenner et al., page 3 [92].

Hence, the novelty concept of MR-X is developed from the above reference, with the

support of previous experience in analysing LNG processes and their advantages and

disadvantages.

4.3.1. Principle of Operation

The principle of operation for the MR-X process is almost similar to the AP-X process. It

differs mainly only on the pre-cooling section. This includes the type of pre-cooling

refrigerant in which the MR-X process used mixed-refrigerant. Other parameters such as the

MCHE mixed-refrigerant compositions, drivers, and climatic conditions may also differ, but

they are less important compared to the two mentioned points as far as both operations are

concerned.

Fig. B.12 shows a simplified schematic of the MR-X process. It consists of three

blocks: pre-cooling, liquefaction and sub-cooling. For the pre-cooling and liquefaction

blocks, mixed refrigerants are used, whereas nitrogen is employed for the sub-cooling block.

The flow diagram of the MR-X process is shown in Fig. B.13. The C3MR process is a part of

the MR-X process. All three processes (C3MR, AP-X and MR-X) use the multicomponent

5 M: Malaysian, MR: mixed-refrigerant, X: acknowledgement to AP-X

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mixture of the refrigerants as the working fluid for their systems. It can contain

hydrocarbons such as methane, ethane, propane, butane, and other light hydrocarbons, and

also nitrogen [26,50]. Note that suggestions for improving AP-X process given in [50] are

also implemented for MR-X process.

Operationally as other LNG process plants, wet natural gas (NG) undergoes the pre-

treatment process. In the pre-treatment process, acid gas is firstly being removed from the

NG. It then proceeds to dehydration (water removal) and mercury removal process. After

that, natural gas liquids (NGL) are withdrawn from the NG. The NGL normally is sent for

fractioning various by-products such as Ethane and LPG among others. The pre-treatment

and fractionation processes for the MR-X process plant, however, are not analysed in detail

in this thesis.

Cleaned and treated natural gas Stream is cooled down to − C within two heat exchangers of the pre-cooling block. The main mixed-refrigerant is fed (Stream 111) in

parallel to the natural gas for generating a refrigeration effect. Within the pre-cooling block,

there is a two-stage compression process with interstage cooling done by the ejection of the

vapor of the refrigerant in the pipe between the compressors. The liquefaction process takes

place within the “cryogenic heat exchanger” (simulated as the combination of HEX0 and

HEX1). The natural gas in the form of LNG leaves the liquefaction block with a temperature of − C Stream . The two-stage compression process with interstage cooling (using

cooling water) is part of the liquefaction block. The LNG sub-cooling process down to − C (Stream 6) takes place within HEX2. The sub-cooling block consists of a three-stage

compression process with interstage cooling (using cooling water) as well as an expansion

process and HEX 3 used for both (liquefaction and sub-cooling) blocks. Finally, the liquefied

natural gas is throttled to a slightly above ambient pressure value for storage and shipping

purposes.

4.3.2. Simulation and Energy Analysis

The MR-X process has been simulated using the Aspen Plus software [66]. The Soave-

Redlich-Kwong equation of state was applied. The ambient temperature was assumed to be

equal to 43 C, and the pressure 1.013 bar. The liquefaction capacity of the MR-X system was

assumed to be equal to 7.8 MTA, which corresponds to a mass flow rate of the natural gas at

224.22 kg/s. The compositions of all working fluids are shown in Table B.25. The

thermodynamic data for the material streams are given in Table B.26.

During the simulation, in the nitrogen circuit, the expanded stream 1005 should

carry the gas fraction at around 85% to avoid crossover in the HEX3 multi-flow heat

exchanger. For software issue, to simulate multi-flow heat exchanger, only one-side input

should be entered and the other side should be left emptied. Otherwise, crossover problem

would occur due to the conflict between both entered values. This also applies to a three-

flow heat exchanger where the freedom of calculation is necessary.

4 . P r o c e s s e s o f L i q u e f a c t i o n o f N a t u r a l G a s i n M a l a y s i a

72

The following data have been obtained from the energetic analysis:

Power: =LPMRW 113.74 MW, =MPMRW 40.63 MW, =2LPNW 60.12 MW, =2MPNW 70.41

MW, =2HPNW 38.41 MW, =1PCW 13.81 MW, =2PCW 30.27 MW, =1PW 0.89 MW, and =1EXPW

28.69 MW (note that for the energetic and exergetic analyses it is assumed that the

expander EXP1 does not generate electricity but drives one of the compressors of the sub-

cooling block). Heat transfer rates: 1PHEQ = 148.1 MW, 2PHEQ = 99.47 MW, 0HEXQ = 557.02

MW, 1HEXQ = 40.35 MW, 2HEXQ = 41.25 MW, and 3HEXQ = 120.11 MW. For the coefficient of

performance (COP), it is calculated as

netW

HHCOP

61 −=

(4.14)

where

1212

221

PPCPCHPN

MPNLPNMPMRLPMREXPnet

WWWW

WWWWWW

++++

++++−=

(4.15)

obtaining the COP = 0.58.

The cumulative curves for the MR-X process (Fig. B.14) consist of the processes

within the three blocks: pre-cooling, liquefaction and sub-cooling. The curves for the

liquefaction and sub-cooling processes correlate very well with the results reported by

other authors (for example [26,74,95,96]). The distance between the curves for the pre-

cooling process demonstrates the advantage of replacing propane as a one-component

refrigerant by a mixed refrigerant.

4.3.3. Exergy Analysis

An energetic analysis is complemented and enhanced by an exergetic one, in which one

calculates the real thermodynamic inefficiencies (exergy destructions) within a system, the

true thermodynamic value of all energy carriers, and variables that unambiguously

characterize the performance of a system (or one of its components) from the

thermodynamic viewpoint [62]. Since all processes here occur either below the

environmental temperature or cross it, the physical exergy of each material stream is split

into its thermal and mechanical parts [97].

The exergy of fuel, k,FE and the exergy of product, k,PE of the MR-X process are

defined according to the principles described in [62,66,84,98] by considering the fact that

the processes occur below ambient temperature, as they are reported in [23,97]. All coolers

and mixers are dissipative components [62,99]. The results obtained from the exergetic

analysis are shown in Fig. 4.9 and Fig. 4.10.

4 . P r o c e s s e s o f L i q u e f a c t i o n o f N a t u r a l G a s i n M a l a y s i a

73

Fig. 4.9. Exergy destruction (MW) and exergy destruction ratio (%) for the components of

the MR-X process.

Fig. 4.10. Exergetic efficiency of selected components of the MR-X process.

EXP115.236%

HEX022.5410%

HEX15.352%

HEX24.462%

HEX310.544%

PHE13.141%

PHE29.984%

LPMR27.2012%

MPMR10.294%

LPN214.166% MPN2

20.359% HPN2

9.334%

PC13.702%

PC27.643%

Others9.984%

Dissipatives42.5218%

V03.451%

V15.672%

V20.240%

V38.023%

4 . P r o c e s s e s o f L i q u e f a c t i o n o f N a t u r a l G a s i n M a l a y s i a

74

The values of the exergy destruction ratio (Fig. 4.9) are calculated by tot,D

k,Dk

E

Ey

= . For

the overall MR-X process the following results obtained:

• The exergy of fuel for the overall system defined by ( )MM

nettotF EEWE 61, −+= is

equal to =tot,FE 342 MW.

• The exergy of the product for the overall system defined by TTtot,P EEE 16

−= is equal

to =tot,PE 109 MW.

• Since all coolers and mixers are dissipative components, =tot,LE 0.

• The total exergy destruction within all system components is =tot,DE 231 MW, and

• The overall exergetic efficiency is equal to 32%.

For the exergetic evaluation of the MR-X process, the concept of “exergy of

product/exergy of fuel” was applied. Therefore, a comparison between the results obtained

in this section for the MR-X process and the results obtained by other authors for the C3MR,

AP-X and DMR processes are not meaningful.

The results obtained from the exergetic analysis show that the components with the

highest values of the exergy destruction are the LPMR with y=12% and HEXO with y=10%

within the liquefaction block, as well as MPN2 (y=9%), LPN2 (y=6%), and EX1 (y=6%)

within the sub-cooling block. However, all components of the MR-X process have relatively

high exergetic efficiencies. The maximum potential for improving the MR-X process from the

thermodynamic viewpoint is within the sub-cooling block.

4.3.4. Economic Analysis

Reference to the previous economics of C3MR and AP-X process plants is essential to the

MR-X process. No economic data related to the AP-X are given in the open literature. Only a

comparison between C3MR and AP-X was reported in [84], and a general plant comparison

in [87] is available. The following assumptions were made for the analysis: (a) the total

plant capacity is equal to 8 MTA, the operation conditions, marine facilities and temporary

infrastructure remain the same for both LNG plants; (b) the C3MR plant is working with two

trains, while the AP-X plant uses one train. As a result, it is obtained that (a) the cost of plant

facilities for the C3MR and AP-X plants are estimated at 1.795 and 1.596 (2013 $ bn),

respectively; and (b) the capital unit cost is 0.224 and 0.119 (2013 $ bn/mtpa) for the C3MR

and AP-X plants, respectively. These values are 30% lower than the values reported in Table

C.18. This can be explained by the fact that the offsite costs (storage, loading, and whatnot) were left out of the economic analysis they usually amount to − % of the overall costs of an LNG plant) [84].

4 . P r o c e s s e s o f L i q u e f a c t i o n o f N a t u r a l G a s i n M a l a y s i a

75

Therefore, an economic analysis was carried out based on the TRR (Total Revenue

Requirement) method [62]. In order to conduct a detailed economic analysis, it is important

to “accurately” estimate the purchased equipment cost (PEC) while the remaining cost

contributors are calculated based on this value. The most critical assumptions and results

are given for each component group:

• Multi-flow, low-temperature heat exchangers (i.e., HEX0, HEX1, HEX2, HEX3, PHE1

and PHE2) are considered as quite challenging due to their particular design. Only a

few companies in the world produce such heat exchangers; therefore, the data

required for the economic analysis of the heat exchangers used for LNG processes

are confidential and not available in the open literature. Before a cost estimation,

one needs to calculate the heat transfer area (A) based on the operation conditions

and to estimate the value of the overall heat transfer coefficient (U) [82,100] for

each heat exchanger: UHEX0=2500 W/m2K; UHEX1=1700 W/m2K; UHEX2=1200 W/m2K;

UHEX3=1300 W/m2K; UPHE1=1200 W/m2K and UPHE2=1000 W/m2K.

• Turbomachinery (i.e., propane and mixed refrigerant compressors, as well as the

propane expander) is the second most important group of equipment within MR-X

process. Centrifugal compressors are selected. All compressors are driven by electric

motors. For the cost calculations of these equipment-items, the net required power (

kW ) is used as a sizing factor [100].

• Remaining equipment items – interstage coolers, separators, mixers and throttling

valves. Interstage coolers are considered to be shell-and-tube heat exchangers. The

UA values of the coolers were obtained from the AspenPlus simulation software, and

the overall heat transfer coefficients (U) were taken from the literature, i.e., U=500

W/m2K [101]. Separators, mixers and throttling valves were evaluated economically

based on data reported in [100].

Fig. 4.11 shows the estimated values of PEC for selected components and the distribution of

PEC among the components.

4 . P r o c e s s e s o f L i q u e f a c t i o n o f N a t u r a l G a s i n M a l a y s i a

76

Fig. 4.11. PEC ($ mil) for selected components of the MR-X process and distribution of the

PEC among the components.

4 . P r o c e s s e s o f L i q u e f a c t i o n o f N a t u r a l G a s i n M a l a y s i a

77

Table 4.3. Estimation of the fixed-capital investment.

2013 $

Direct costs

Onsite costs

Total purchased equipment cost (PEC) 700

PEC installation (45% of PEC) 315

Piping (35% of PEC) 245

Instrumentation and control (20 % of PEC) 140

Electrical equipment and materials (20% of PEC) 140

Total onsite costs 1540

Offsite costs 0

Total direct costs 1540

Indirect costs

Engineering and supervision (35% of PEC) 245

Construction costs (15% of DC) 231

Contingencies (10% of IC) 48

Total indirect costs 524

Fixed-capital investment 2064

The calculation of fixed-capital investment (including assumptions made) is given in Table

4.3. Since the location of the evaluated LNG plant is unknown, office costs were not taken

into consideration. Based on the estimated fixed-capital investment and the assumptions for

the economic, financial, operating, and market input variables, the total revenue

requirement is calculated on a year-by-year basis. Finally, the non-uniform annual monetary

values associated with the investment (carrying charges, CC), operating, maintenance

(OMC), and fuel costs (FC) of the system being analyzed) are levelized, that is they are

converted into an equivalent series of constant payments (annuities). The series of annual

costs associated with carrying charges (CCj) and expenses (FCj and OMCj) for the jth year of

plant operation are not uniform. The levelized carrying charges are calculated as

( ) CRFInvestmentCapitalTotalCCL ×= (4.16)

where the capital-recovery factor (CRF) is given by

( )( ) 11

1

−+

+=

neff

neffeff

i

iiCRF

(4.17)

Here, the ieff is the average annual effective discount rate (cost of money), and n

denotes the plant economic life expressed in years.

For the fuel, the constant escalation levelization factor (CELF) calculates

4 . P r o c e s s e s o f L i q u e f a c t i o n o f N a t u r a l G a s i n M a l a y s i a

78

( )( ) CRF

k

kkFCCELFFCFC

FC

nFCFC

L−

−=×=

1

100

(4.18)

with eff

FCFC

i

rk

++

=1

1 and rFC = const. The term rFC denotes the average annual nominal

escalation rate for fuel cost. The levelized annual operating and maintenance costs OMCL are

given by

( )( ) CRF

k

kkOMCCELFOMCOMC

FC

nFCFC

L−

−=×=

1

100

(4.19)

with eff

OMCOMC

i

rk

++

=1

1 and rOMC = const, where rOMC is the average annual nominal escalation

rate for the operating and maintenance costs. Finally, the levelized total revenue

requirement (TRR) is obtained from

TRR = CC + FC + OMC (4.20)

4 . P r o c e s s e s o f L i q u e f a c t i o n o f N a t u r a l G a s i n M a l a y s i a

79

Fig. 4.12. Levelized total revenue requirement for the MR-X process using different

assumptions for the economic analysis: OMC as a function of CC - between 1% and 10% and

cost of the electricity – between 0.05 and 0.20 $/kWh.

4 . P r o c e s s e s o f L i q u e f a c t i o n o f N a t u r a l G a s i n M a l a y s i a

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The following assumptions have been made for the economic analysis:

• The operating and maintenance costs are a function of the CC and are varied

between 1% and 10%.

• The cost of electricity is assumed to vary between 0.05 and 0.20 $/kWh.

• The LNG plant operates with a 100% capacity during 7446 hours per year

(capacity factor = 85%).

• The average cost of money is ieff= 10%.

• The plant economic life is n=20 years.

• The average general inflation rate is rn=2.5%.

It is fully considered that the reported economic data cannot be accepted as absolute

correct because no detailed cost information is given in the open literature. However, the

relative economic characteristics are also very helpful for evaluating novel (and non-

proven) processes.

Fig. 4.13. Cost per unit of mass of the liquefaction process when different assumptions for

the economic analysis are used.

The results obtained from the economic analysis show that approximately 46% of the total

PEC are associated with multi-flow, low-temperature heat exchangers; % − with turbomachinery and only 3% with the remaining equipment. The preliminary estimation of

the fixed-capital investment of the novel MR-X LNG process is 2.064 $ bn or 0.265 (2013 $

bn/MTA). These values correlate well with the data reported in this section and the data

associated with the C3MR and AP-X process plants. The operation and maintenance costs

affect significantly the economic performance of the LNG plant. At low fuel cost (cost of

4 . P r o c e s s e s o f L i q u e f a c t i o n o f N a t u r a l G a s i n M a l a y s i a

81

electricity), the values of CC and OMC are dominant. However, with increasing fuel cost

(0.10 $/kWh) the fuel costs became comparable with the CC and OMC and even dominate.

The cost of the liquefaction process varies between 0.07 and 0.19 $/kgLNG (Fig. 4.13) using

different assumptions for the economic analysis. The results obtained from the economic

analysis show that the assumptions made play a significant role. These assumptions refer to

the cost of electricity and to OMC (% of CC), i.e. they depend on the country where the LNG

plant is located.

82

5. Conclusion and Future Works

In this thesis, several high energy-intensity LNG plants have been evaluated using energetic,

exergetic and economic analyses. A new LNG process concept has been designed, the MR-X

process, and has been analyzed to evaluate its thermodynamic and economic performances

and efficiency. All LNG processes possess particular characteristics where direct comparison

may not produce accurate perspectives. Nevertheless, in improving existing processes and

designing the new ones, scientists generally implement at least a general benchmark on

current technologies the world realistically has, and from where and how they can proceed

further.

Among important parameters to understand and analyze the LNG processes are

their principle of operations, through practical simulations and methodologies to analyze

them. While many ways are available for researchers to select in understanding and

bringing the technologies to the next level, practical solutions are significant especially to

the process owners and their stakeholders. The LNG circuits are complex enough, and in

analyzing them properly and practically without losing important characteristics are

delicate. Simulating the LNG process in Aspen was helpful, and many potentials for future

works are derived from such platform. However, like many other tool, the methodologies,

including energetic, exergetic, and economic analyses, have rooms for improvements. The

LNG processes itself, for example their multi-components compositions, compressor-power

arrangement, and heat exchangers’ duty are as well have various ways for further

advancement.

The newly introduced MR-X process combines the advantages and as much as

possible overcomes the limitations of the C3MR, AP-X and DMR processes. The MR-X system

coefficient of the performance and the exergetic efficiency are equal to 0.58 and 32%,

respectively. These relatively high values demonstrate that the novel process is efficient

from the thermodynamic viewpoint.

The preliminary estimation of the fixed-capital investment for the MR-X system

leads to 2.064 $ bn or 0.265 (2013 $ bn/MTA). The values reported for the AP-X plant are

slightly higher. This is an important point that should be clarified in the future, i.e. the effect

of the assumptions used for the economic analysis of an LNG plant needs to be studied.

Advanced exergy-based methods are recommended to be applied in the future, to obtain

more accurate and more specific information about the options for improving the

performance of the novel MR-X process from the economic and environmental viewpoints.

5 . C o n c l u s i o n a n d F u t u r e W o r k s

83

The AP-X process is the largest LNG power plant so far exists. The process is not only

complex due to its multi-component refrigeration in the MCHE, but it also uses the classical

pure propane refrigerant to precool a never-seen-in-history LNG capacity. The economics of

scale could be used for its justification, however the use of mixed-refrigerant as the

precooling agent is engaging, as per proven by the MR-X process. The AP-X system

coefficient of the performance and the exergetic efficiency are equal to 0.41 and 6%,

respectively. 24% of the total exergy destruction is associated with the multi-flow heat

exchanger HEX0 and HEX1. These are the MCHE which developers are strict on their data

confidentiality. The values may have better accuracies should more information about the

system and the system components publicly available, and advanced exergy is implemented,

where the interdependence between the components as well as the real potential for

improving the overall system will be discovered and discussed. The preliminary estimation

of the fixed-capital investment for the MR-X system leads to 5.8 $ bn or 0.95 (2013 $

bn/MTA).

Since all LNG processes are not only energy-intensive but also cost- and

environmental-impact-intensive, the design of the heat exchanger should be a central focus

of such LNG plants, especially the heat transfer characteristics. Despite the fact that the

investment cost of the heat exchanger is relatively high due to the large heat transfer surface

and to a complex and unique design, decreasing the inefficiencies within this component

and accepting higher investment cost will finally lead to a significant decrease in the overall

cost of the generated LNG.

In other perspective, for example the view of current global awareness and practical

applications in many governmental and societal levels to protect the earth against global

warming and ozone layer depletion, the selection and usage of refrigerants in the

liquefaction of natural gas should be considered essential. There are adverse effects

generated from the usage of the refrigerants, categorically could be in two perspectives: 1)

direct effect, and 2) indirect effects.

The direct effects are widely known for the past decades, as the development of non-

toxic and non-flammable refrigerants happening throughout the globe, and the application

of Montreal and Kyoto protocols at almost all countries and refrigerant manufacturers.

Notably while these developments and protocols are being actively referred to, the

flammable and toxic refrigerants may still exist because of research and military necessities.

Moreover, while direct effects are often clear for us to point out, the indirect effects of the

refrigerants are not so clear. LNG plants particularly in Malaysia have moved from steam-

based drivers towards gas-turbine drivers. The steam-based components, compared to the

current gas-turbine driven system, have more mechanical parts and, therefore, are less

efficient in electricity consumption as fuel. However, present gas-turbine drivers are still

dependent on fossil fuel-based starters.

In general, the indirect effects are due to the involvement of electricity or mechanical

components that needed to drive these refrigerants. The electrical and mechanical parts of

5 . C o n c l u s i o n a n d F u t u r e W o r k s

84

the LNG system need to be driven by fossil fuel that in turn indirectly contaminates earth's

environment. Subsequently, this contamination would cost LNG plant provider in monies

term where local government policies restricted the environment contamination. Indirectly

also, the reputation and future opportunities for these fossil-fuel-plant companies are in

risks.

The FLNG process is very near future work for LNG scientists. So far while there are

only Malaysian and Australian FLNG systems, the world is racing harder than before in this

new technology where different challenges are to be embraced. Deeper and stranded gas

reserves are becoming reachable than before as demand in higher quality of energy as well

as in environmental protection are increasing. As Malaysian LNG provider is the current

global pioneer on FLNG, studies on the sloshing effect of refrigerants and natural gas being

liquefied on ship, as well as all analyses done in this thesis and their proposals as future

works could be executed on this offshore process.

The analyzed MR-X process has shown the beneficial use of the mixed-refrigerant.

While not denying the use of Nitrogen as the sole refrigerant in the recently-launched AP-

NTM FLNG system, exergetic efficiency may come in handy in seeing these processes in terms

of exergetic efficiency performance and their exergy destructions. Furthermore,

exergoeconomics and exergoenvironmental analyses should be applied, that will provide

clearer reasonable and practical results on this latest LNG arena.

85

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Appendix A. Research

Contributions

1. Morosuk T, Tesch S, Hiemann a., Tsatsaronis G, Bin Omar N. 2015.

“Evaluation of the PRICO liquefaction process using exergy-based methods.”

Journal of Natural Gas Science and Engineering.

2. Bin Omar MN, Morosuk T, Tsatsaronis G. 2014.

“Thermodynamic And Economic Evaluation Of A Novel Mixed-Refrigerant Process

For The Liquefaction Of Natural Gas.”

Proceedings of the ASME2014 International Mechanical Engineering Congress &

Exposition (IMECE2014), Montreal, Canada.

3. Bin Omar N, Morosuk T, Tsatsaronis G. 2014.

“A Novel Mixed-Refrigerant Process for the Liquefaction of Natural Gas.”

Proceedings of ECOS 2014 - The 27th International Conference on Efficiency, Cost,

Optimization, Simulation and Environmental Impact of Energy Systems on June 15-

19, Turku, Finland.

4. T. Morosuk, A. Hiemann, N. Bin Omar, Tsatsaronis G. 2013.

“Exergy-Based Methods Applied To The Chain “Natural Gas – LNG – Natural Gas” − Liquefaction Using A Single Mixed-Refrigerant Process.”

3rd International Exergy, Life Cycle Assessment, and Sustainability Workshop &

Symposium (ELCAS3) 07-09 July, 2013, NISYROS – GREECE.

5. Omar MN Bin, Morosuk T, Tsatsaronis G. 2012.

“Exergy Analyses Applied to an AP-X Process for the Liquefaction Of Natural Gas.”

Proceedings of the ASME 2012 International Mechanical Engineering Congress &

Exposition IMECE2012, Houston, Texas.

6. Morosuk T, Nazri M, Omar B, Tsatsaronis G, Naw R. 2010.

“Advanced Exergetic Analysis of a Refrigeration System for Liquefaction of Natural

Gas.”

Proceedings of the 23-rd International Conference on Efficiency, Cost, Optimization

Simulation and Environmental Impact of Energy Systems (ECOS2010), Lausanne,

Switzerland.

92

Appendix B. Energy and Exergy

Analyses – Data, Flow and

Results

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

93

B.1 General Information on Liquefaction Processes

Fig. B.1. Classification of natural gas liquefaction processes [93].

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

94

Table B.1. Malaysian LNG Plants.

Bintulu MLNG

Satu

Bintulu MLNG

Dua

Bintulu MLNG

Tiga

PETRONAS

Floating LNG

Project 1

(PFLNG 1)

PETRONAS

Floating LNG

Project 2

(PFLNG 2)

Start-up 1983 1995 2003 Late 2015 2018

Operator MLNG Satu MLNG Dua MLNG Tiga PFLNG 1 PFLNG 2

Shareholders

PETRONAS

(60%), Shell

(17.5%),

Mitsubishi

(17.5%),

Sarawak state

government

(5%)

PETRONAS

(60%), Shell

(15%),

Mitsubishi

(15%),

Sarawak state

government

(10%)

PETRONAS

(60%), Shell

(10%), Nippon

Oil (10%),

Occidental

LNG Malaysia

(10%),

Sarawak state

government

(10%)

(TBC) (TBC)

Capacity

(MMTPA) 8.1 7.8 6.8 1.2 1.5

No. of trains 3 3 2 1 1

LNG process

type

Air Products’

C3MR

Air Products’

C3MR

Air Products’

C3MR

Air Products’

AP-NTM

Air Products’

AP-NTM

Gas fields Central

Luconia Basin

Central

Luconia Basin Jintan

Kanowit

(offshore

Sarawak)

Rotan

(offshore

Sabah)

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

95

Table B.2. Liquefaction Plants with specific LNG Technology, sorted by year of project start [102].

Country Project Name Start Year

Nameplate Capacity (mmtpa)

Owners Liquefaction Technology

US Kenai LNG* 1969 1.5 ConocoPhillips ConocoPhillips

Optimized Cascade®

Algeria Skikda - GL1K (T1-4) 1972 1 Sonatrach Teal (T1-3), PRICO (T4)

Brunei Brunei LNGT1-5 1972 7.2 Government of Brunei, Shell, Mitsubishi APC C3MR

Indonesia Bontang LNG T1-2 1977 5.4 Pertamina APC C3MR

United Arab Emirates

ADGAS LNGT1-2 1977 2.6 ADNOC, Mitsui, BP, TOTAL APC C3MR

Algeria Arzew- GL1Z (T1-6) 1978 6.6 Sonatrach APC C3MR

Indonesia Arun LNG T1 1978 1.65 Pertamina APC C3MR

Algeria Arzew-GL2Z (T1-6) 1981 8.2 Sonatrach APC C3MR

Algeria Skikda - GL2K (T5-6) 1981 2.2 Sonatrach PRICO

Indonesia Bontang LNG T3-4 1983 5.4 Pertamina APC C3MR

Malaysia MLNG Satu (T1-3) 1983 8.1 PETRONAS, Mitsubishi, Sarawak State government APC C3MR

Indonesia Arun LNG T6 1986 2.5 Pertamina APC C3MR

Australia North West Shelf T1 1989 2.5 BHP Billiton, BP, Chevron, Shell, Woodside, Mitsubishi,

Mitsui APC C3MR

Australia North West Shelf T2 1989 2.5 BHP Billiton, BP, Chevron, Shell, Woodside, Mitsubishi,

Mitsui APC C3MR

Indonesia Bontang LNG T5 1989 2.9 Pertamina APC C3MR

Australia North West Shelf T3 1992 2.5 BHP Billiton, BP, Chevron, Shell, Woodside, Mitsubishi,

Mitsui APC C3MR

Indonesia Bontang LNG T6 1994 2.9 Pertamina APC C3MR

United Arab Emirates

ADGAS LNG T3 1994 3.2 ADNOC, Mitsui, BP, TOTAL APC C3MR

Malaysia MLNG Dua (T1-3) 1995 7.8 PETRONAS, Shell, Mitsubishi, Sarawak State government APC C3MR

Qatar Qatargas I (T1) 1997 3.2 Qatar Petroleum, ExxonMobil, TOTAL, Marubeni, Mitsui APC C3MR

Qatar Qatargas I (T2) 1997 3.2 Qatar Petroleum, ExxonMobil, TOTAL, Marubeni, Mitsui APC C3MR

Indonesia Bontang LNG T7 1998 2.7 Pertamina APC C3MR

Qatar Qatargas I (T3) 1998 3.1 Qatar Petroleum, ExxonMobil, TOTAL, Mitsui, Marubeni APC C3MR

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

96

Indonesia Bontang LNG T8 1999 3 Pertamina APC C3MR

Nigeria NLNGT1 1999 3.3 NNPC, Shell, TOTAL, Eni APC C3MR

Qatar RasGas I (T1) 1999 3.3 Qatar Petroleum, ExxonMobil, KOGAS, Itochu, LNG Japan APC C3MR

Trinidad ALNG T1 1999 3.3 BP, BG, Repsol, CIC, NGC Trinidad ConocoPhillips

Optimized Cascade®

Nigeria NLNG T2 2000 3.3 NNPC, Shell, TOTAL, Eni APC C3MR

Oman Oman LNG T1 2000 3.55 Petroleum Development Oman (PDO), Shell, TOTAL,

Korea LNG, Partex, Mitsubishi, Mitsui, Itochu APC C3MR

Oman Oman LNG T2 2000 3.55 Petroleum Development Oman (PDO), Shell, TOTAL,

Korea LNG, Partex, Mitsubishi, Mitsui, Itochu APC C3MR

Qatar RasGas I (T2) 2000 3.3 Qatar Petroleum, ExxonMobil, KOGAS, Itochu, LNG Japan APC C3MR

Nigeria NLNGT3 2002 3 NNPC, Shell, TOTAL, Eni APC C3MR

Trinidad ALNG T2 2002 3.5 BP, BG, Repsol ConocoPhillips

Optimized Cascade®

Malaysia MLNG Tiga (T1-2) 2003 6.8 PETRONAS, Shell, Nippon, Sarawak State government,

Mitsubishi APC C3MR

Trinidad ALNG T3 2003 3.5 BP, BG, Shell ConocoPhillips

Optimized Cascade®

Australia North West Shelf T4 2004 4.4 BHP Billiton, BP, Chevron, Shell, Woodside, Mitsubishi,

Mitsui APC C3MR

Qatar RasGas II (T1) 2004 4.7 Qatar Petroleum, ExxonMobil APC C3MRI Split MR™

Egypt ELNG T1 2005 3.6 BG, PETRONAS, EGAS, EGPC, GDF SUEZ ConocoPhillips

Optimized Cascade®

Egypt ELNG T2 2005 3.6 BG, PETRONAS, EGAS, EGPC ConocoPhillips

Optimized Cascade®

Egypt SEGAS T1 2005 5 Gas Natural Fenosa, Eni, EGPC, EGAS APC C3MRI Split MR™

Qatar RasGas II (T2) 2005 4.7 Qatar Petroleum, ExxonMobil APC C3MRI Split MR™

Australia Darwin LNG T1 2006 3.6 ConocoPhillips, Santos, INPEX, Eni, TEPCO, Tokyo Gas ConocoPhillips

Optimized Cascade®

Nigeria NLNG T4 2006 4.1 NNPC, Shell, TOTAL, Eni APC C3MR

Nigeria NLNG T5 2006 4.1 NNPC, Shell, TOTAL, Eni APC C3MR

Oman Qalhat LNG 2006 3.7 Omani Govt, Petroleum Development Oman (PDO), Shell,

Mitsubishi, Gas Natural Fenosa, Eni, Itochu, Osaka Gas, TOTAL, Korea LNG, Mitsui, Partex

APC C3MR

Trinidad ALNG T4 2006 5.2 BP, BG, Repsol, NGC Trinidad ConocoPhillips

Optimized Cascade®

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

97

Equatorial Guinea

EG LNGT1 2007 3.7 Marathon, Sonagas, Mitsui, Marubeni ConocoPhillips

Optimized Cascade®

Norway Snohvit LNG T1 2007 4.2 Statoil, Petoro, TOTAL, GDF SUEZ, RWE Linde MFC

Qatar RasGas II (T3) 2007 4.7 Qatar Petroleum, ExxonMobil APC C3MRI Split MR™

Australia North West Shelf T5 2008 4.4 BHP Billiton, BP, Chevron, Shell, Woodside, Mitsubishi,

Mitsui APC C3MR

Nigeria NLNG T6 2008 4.1 NNPC, Shell, TOTAL, Eni APC C3MR

Indonesia Tangguh LNG T1 2009 3.8 BP, CNOOC, Mitsubishi, INPEX, JOGMEC, JX Nippon Oil & Energy, LNG Japan, Talisman Energy, Kanematsu, Mitsui

APC C3MRI Split MR™

Indonesia Tangguh LNG T2 2009 3.8 BP, CNOOC, Mitsubishi, INPEX, JOGMEC, JX Nippon Oil & Energy, LNG Japan, Talisman Energy, Kanematsu, Mitsui

APC C3MRI Split MR™

Qatar Qatargas II (T1) 2009 7.8 Qatar Petroleum, ExxonMobil APC AP-X

Qatar Qatargas II (T2) 2009 7.8 Qatar Petroleum, ExxonMobil, TOTAL APC AP-X

Qatar RasGas III (T1) 2009 7.8 Qatar Petroleum, ExxonMobil APC AP-X

Russia Sakhalin 2 (T1) 2009 4.8 Gazprom, Shell, Mitsui, Mitsubishi Shell DMR

Russia Sakhalin 2 (T2) 2009 4.8 Gazprom, Shell, Mitsui, Mitsubishi Shell DMR

Yemen Yemen LNG T1 2009 3.35 TOTAL, Hunt Oil, Yemen Gas Co., SK Corp, KOGAS, GASSP,

Hyundai APC C3MRI Split MR™

Malaysia MLNG Dua Debottleneck 2010 1.2 PETRONAS, Shell, Mitsubishi, Sarawak State government APC C3MR

Norway Skangass LNG 2010 0.3 Skangass Linde LIMUM

Peru Peru LNG 2010 4.45 Hunt Oil, Repsol, SK Corp, Marubeni APC C3MRI Split MR™

Qatar Qatargas III 2010 7.8 Qatar Petroleum, ConocoPhillips, Mitsui APC AP-X

Qatar RasGas III (T2) 2010 7.8 Qatar Petroleum, ExxonMobil APC AP-X

Yemen Yemen LNG T2 2010 3.35 TOTAL, Hunt Oil, Yemen Gas Co., SK Corp, KOGAS, GASSP,

Hyundai APC C3MRI Split MR™

Qatar Qatargas IV 2011 7.8 Qatar Petroleum, Shell APC AP-X

Australia Pluto LNG T1 2012 4.3 Woodside, Kansai Electric, Tokyo Gas Shell propane precooled mixed refrigerant design

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

98

Fig. B.2. A typical coil-wound MCHE for a C3MR process-based LNG plant [103].

The picture on the LEFT side shows the coiling of the cooling tube and eventually will be

placed inside a bundled-shell shown in the RIGHT side. Apart from liquefaction purpose, the

shell directly serves as secondary protection in case of tube leakages thus adding more

value towards the selection of MCHE in LNG production.

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

99

B.2 System testing using PRICO® process [23,24,32]

Fig. B.3. Flow diagram of PRICO process: CM1 - Compressor 1; COL - Cooler; CM2 -

Compressor 2; CD - Condenser; HE - Heat exchanger; TV - Throttling Valve [24].

Table B.3. Composition and concentration of natural gas and refrigerants.

Component Formula Refrigerant [34] (% mol)

Natural gas [97] (% mol)

Methane CH4 0.30 0.88

Ethane C2H6 0.30 0.08

Propane C3H8 − 0.02

Butane C4H10 0.25 −

Nitrogen N2 0.15 0.02

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

100

Table B.4. Thermodynamic data for material streams at real operating conditions S

tre

am

Ma

teri

al

stre

am

m

(kg/s)

T

(ºC)

p

(bar)

x

(kg/kg)

h

(kJ/kg)

s

(kJ/kg K)

TE

(kW)

ME

(kW)

PHE

(kW)

1

Re

frig

era

nt

475

15 3 − -2443 -5.15 0.14 9.61 9.75

2 71 8 − -2345 -5.09 2.84 25.36 28.20

3 30 8 − -2320 -5.32 0.03 25.36 25.39

4 91 22 − -2313 -5.25 5.86 60.50 66.36

5 30 22 0.84 -2510 -5.87 0.08 61.36 61.44

6 -159 19 0 -3160 -9.22 94.29 133.25 227.53

7 -162 6 0.07 -3159 -9.19 108.33 114.07 222.39

8

Na

tura

l g

as

38 67 − -4162 -6.77 0.03 24.91 24.94

9 50 -159 64 0 -4999 -11.12 16.49 31.48 47.97

10 -163 1 0.05 -4999 -10.98 46.08 0.18 46.26

11 3 -163 1 1 -3403 -4.85 0.51 0.18 0.69

12 47 -163 1 0 -5088 -11.32 45.58 0.18 45.76

Table B.5. Reference values for the exergetic analysis (state 0) for material streams.

Stream Material stream

T

(ºC)

p

(bar)

h

(kJ/kg)

s

(kJ/kg K)

1 - 7 Refrigerant 25 1.013

-2426 -4.81

8-10 Natural gas -4141 -5.03

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

101

Table B.6. Detailed thermodynamic data of each chemical component in the streams within

mixed refrigerant.

Stream 1 2 3 4 5 6 7 8 9 10 11 12

Vapor Mass Fraction (kg/kg)

Methane 0.148 0.148 0.148 0.148 0.144 − 0.002 0.786 − 0.035 0.661 −

Ethane 0.277 0.277 0.277 0.277 0.249 − 0.000 0.134 − 0.000 0.000 −

Propane 0.000 0.000 0.000 0.000 0.000 − 0.000 0.049 − 0.000 0.000 −

Butane 0.446 0.446 0.446 0.446 0.224 − 0.000 0.000 − 0.000 0.000 −

Nitrogen 0.129 0.129 0.129 0.129 0.128 − 0.057 0.031 − 0.018 0.338 −

Liquid Mass Fraction (kg/kg)

Methane − − − − 0.004 0.148 0.145 − 0.786 0.751 − 0.793

Ethane − − − − 0.028 0.277 0.277 − 0.134 0.134 − 0.141

Propane − − − − 0.000 0.000 0.000 − 0.049 0.049 − 0.052

Butane − − − − 0.222 0.446 0.446 − 0.000 0.000 − 0.000

Nitrogen − − − − 0.001 0.129 0.072 − 0.031 0.013 − 0.014

Partial Pressure (bar)

Methane 0.90 2.40 2.40 6.60 7.69 − 0.43 58.96 − 0.80 0.80 −

Ethane 0.90 2.40 2.40 6.60 7.10 − 0.00 5.36 − 0.00 0.00 −

Propane − − − − − − − 1.34 − 0.00 0.00 −

Butane 0.75 2.00 2.00 5.50 3.31 − 0.00 − − − − −

Nitrogen 0.45 1.20 1.20 3.30 3.91 − 5.57 1.34 − 0.24 0.24 −

Table B.7. Definition of the exergy of fuel and the exergy of product for the components of

the PRICO® process.

Component k,FE k,PE

CM1 T

CMCM,F EWE 111 +=

TMM

CM,P E)EE(E 2221 +−=

COL Dissipative component, 32 EEE COL,D −=

CM2 22 CMCM,F WE = 342 EEE CM,P

−=

CD Dissipative component, 54 EEE CD,D −=

HE TTMM

TT

HE,F

EE)EE(

)EE(E

8517

17

++−+

+−=

TTMM

MM

HE,P

EE)EE(

)EE(E

9656

89

++−+

+−=

TV1 MM

TV,F EEE 761 −=

TT

TV,P EEE 671 −=

TV2 MM

TV,F EEE 1092 −=

TT

TV,P EEE 9102 −=

Overall

system 21 CMCMtot,F WWE += 810 EEE tot,P −=

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

102

Table B.8. Results obtained from the exergetic analysis.

Component k,FE

(MW)

k,PE

(MW)

k,DE

(MW)

(%)

CM1 44.84 37.81 7.03 84.3

COL − − 2.80 −

CM2 46.51 40.96 5.55 88.1

CD − − 4.91 −

HE 231.98 189.23 42.75 81.6

TV1 19.18 14.04 5.14 73.2

TV2 31.66 29.59 2.06 93.5

Overall

91.21 20.96 70.24 23.0

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

103

B.3 C3MR Process [73]

Table B.9. Composition for the C3MR process mixed-refrigerant in molar percentage.

Component Formula

Mole fraction (%)

Mixed-refrigerant

Natural gas

Methane CH4 41.8 87.5

Ethane C2H6 29.9 5.5

Propane C3H8 21.3 2.1

N-Butane C4H10 - 0.5

I-Butane C4H10 - 0.3

Nitrogen N2 7.0 4.1

Table B.10. LMR and HMR compositions.

Component Formula Mixed-refrigerant (% mol)

LMR HMR

Methane CH4 66.6 28.0

Ethane C2H6 14.9 38.2

Propane C3H8 3.4 31.3

Nitrogen N2 15.1 2.5

Table B.11. Boiling temperatures (in C) for refrigerants at different pressures [104].

Methane Ethane Propane Nitrogen

1 bar -161.7 -89.1 -42.2 -195.9

3 bar -146.5 -66 -13.9 -185.3

6 bar -134.5 -47.9 8.2 -176.8

19 bar -108.7 -9.3 54.8 -158.5

22 bar -104.7 -3.5 61.8 -155.8

Table B.12. Stream 7 molar fraction and its partial pressures

Component Mixed-

refrigerant (% mol)

Partial pressure

(bar)

Methane 41.8 17.82 Ethane 29.9 12.75

Propane 21.3 9.08

Nitrogen 7.0 2.98

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

104

Table B.13. Enthalpy and Entropy Values Required for Exergy Calculation. S

tre

am

Ma

teri

al

Str

ea

m

T

(K)

P

(bar)

hj (kJ/kg) sj (kJ/kg.K) ���� ���� ���� ���� ���� ����

1

Mix

ed

-re

frig

era

nt

238.57 2.82 -3009.96 -2906.56 -2903.83 -5.62 -5.23 -4.91

2A 300.95 7.50 -2908.57 -2913.78 -2903.83 -5.53 -5.55 -4.91

2B 361.72 17.50 -2800.92 -2930.20 -2903.83 -5.46 -5.85 -4.91

2C 305.00 17.33 -2916.40 -2929.91 -2903.83 -5.80 -5.85 -4.91

2D 382.27 48.60 -2788.57 -2993.79 -2903.83 -5.72 -6.32 -4.91

3 305.00 48.20 -2974.43 -2992.80 -2903.83 -6.26 -6.32 -4.91

4 291.00 46.74 -3032.58 -2989.24 -2903.83 -6.45 -6.30 -4.91

5 279.00 45.33 -3108.24 -2985.87 -2903.83 -6.71 -6.28 -4.91

6 257.00 43.96 -3228.68 -2982.65 -2903.83 -7.15 -6.26 -4.91

7 240.00 42.63 -3307.31 -2979.59 -2903.83 -7.46 -6.25 -4.91

8 240.00 42.63 -3319.20 -3187.16 -3143.89 -5.96 -5.47 -3.87

9 146.00 41.34 -3822.78 -3185.78 -3143.89 -8.57 -5.45 -3.87

10 134.00 40.09 -3860.07 -3184.45 -3143.89 -8.83 -5.44 -3.87

11 118.04 3.00 -3860.07 -3145.89 -3143.89 -8.72 -4.31 -3.87

12 125.98 2.91 -3743.54 -3145.80 -3143.89 -7.77 -4.29 -3.87

13 135.13 2.91 -3605.33 -2906.69 -2903.83 -8.85 -5.24 -4.91

14 240.00 42.63 -3302.76 -3000.20 -2812.01 -8.04 -6.92 -5.37

15 146.00 41.34 -3552.42 -2992.40 -2812.01 -9.35 -6.89 -5.37

16 141.34 2.91 -3552.42 -2815.42 -2812.01 -9.29 -5.67 -5.37

17

Pro

pa

ne

, C3H

8

331.52 14.30 -2351.54 -2742.93 -2376.42 -6.50 -7.74 -6.11

18 305.00 14.30 -2722.83 -2742.93 -2376.42 -7.68 -7.74 -6.11

19 287.34 7.20 -2722.83 -2393.78 -2376.42 -7.67 -6.52 -6.11

20 287.34 7.20 -2675.58 -2393.78 -2376.42 -7.50 -6.52 -6.11

21 287.34 7.20 -2773.09 -2393.78 -2376.42 -7.84 -6.52 -6.11

22 275.45 5.10 -2773.09 -2387.45 -2376.42 -7.84 -6.44 -6.11

23 275.45 5.10 -2687.94 -2387.45 -2376.42 -7.53 -6.44 -6.11

24 275.45 5.10 -2805.09 -2387.45 -2376.42 -7.95 -6.44 -6.11

25 253.91 2.50 -2805.09 -2380.27 -2376.42 -7.94 -6.29 -6.11

26 253.91 2.50 -2600.62 -2380.27 -2376.42 -7.14 -6.29 -6.11

27 253.91 2.50 -2859.76 -2380.27 -2376.42 -8.16 -6.29 -6.11

28 237.14 1.30 -2859.76 -2377.15 -2376.42 -8.15 -6.16 -6.11

29 238.50 1.30 -2471.73 -2377.15 -2376.42 -6.51 -6.16 -6.11

30 265.42 2.50 -2434.22 -2380.27 -2376.42 -6.48 -6.29 -6.11

31 258.13 2.50 -2445.80 -2380.27 -2376.42 -6.53 -6.29 -6.11

32 289.34 5.10 -2402.76 -2387.45 -2376.42 -6.49 -6.44 -6.11

33 285.07 5.10 -2410.09 -2387.45 -2376.42 -6.52 -6.44 -6.11

34 301.12 7.20 -2388.42 -2393.78 -2376.42 -6.50 -6.52 -6.11

35 297.41 7.20 -2395.10 -2393.78 -2376.42 -6.53 -6.52 -6.11

36 287.34 7.20 -2413.10 -2393.78 -2376.42 -6.59 -6.52 -6.11

37 275.45 5.10 -2426.45 -2387.45 -2376.42 -6.58 -6.44 -6.11

38 253.91 2.50 -2452.42 -2380.27 -2376.42 -6.55 -6.29 -6.11

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

105

Fig. B.4. C3MR flowsheet using Aspen Plus [73].

NG

LNG

Liquefaction and

Sub-cooling Block

Pre-cooling

Block

Legend:

Natural Gas LNG refrigerant (Propane) mixed-refrigerant mixed-refrigerant (gas) mixed-refrigerant (liquid)

Final Product

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

106

Table B.14. Thermodynamic data for the material streams at real operating conditions for

C3MR process [73].

Str

ea

m

Ma

teri

al

stre

am

m T p x TE

ME

PHE

(kg/s) (°K) (bar) (kg/kg) (kW) (kW) (kW)

1 M

ixe

d R

efr

ige

ran

t 301.84 238.57 2.82 1.00 3556.46 28059.29 31615.75

2A 301.84 300.95 7.50 1.00 7.34 54453.44 54460.77

2B 301.84 361.72 17.50 1.00 3691.12 76447.38 80138.51

2C 301.84 305.00 17.33 1.00 46.17 76202.22 76248.39

2D 301.84 382.27 48.60 1.00 7234.19 99939.58 107173.77

3 301.84 305.00 48.20 1.00 62.28 99775.23 99837.51

4 301.84 291.00 46.74 0.91 185.95 99158.24 99344.20

5 301.84 279.00 45.33 0.73 1315.82 98535.69 99851.51

6 301.84 257.00 43.96 0.50 5553.83 97904.16 103457.98

7 301.84 240.00 42.63 0.36 10572.31 97264.45 107836.77

8 83.58 240.00 42.63 1.00 1273.96 36167.88 37441.84

9 83.58 146.00 41.34 0.00 24425.44 35890.78 60316.22

10 83.58 134.00 40.09 0.00 28179.01 35613.12 63792.13

11 83.58 118.04 3.00 0.18 50361.68 10719.10 61080.78

12 83.58 125.98 2.91 0.45 36602.91 10419.07 47021.97

13 301.84 135.13 2.91 0.22 113993.78 28915.64 142909.43

14 218.27 240.00 42.63 0.00 6465.87 59991.16 66457.03

15 218.27 146.00 41.34 0.00 37406.26 59706.93 97113.19

16 218.27 141.34 2.91 0.08 75098.15 18480.75 93578.90

17

Pro

pa

ne

, C3H

8

442.70 331.52 14.30 1.00 8842.63 52910.10 61752.73

18 442.70 305.00 14.30 0.00 101.36 52910.10 53011.46

19 442.70 287.34 7.20 0.14 5317.29 46310.07 51627.36

20 442.70 287.34 7.20 0.27 4530.45 46310.07 50840.52

21 322.79 287.34 7.20 0.00 4487.48 33766.25 38253.74

22 322.79 275.45 5.10 0.08 9723.92 28144.07 37867.99

23 322.79 275.45 5.10 0.31 7458.59 28144.07 35602.66

24 222.92 275.45 5.10 0.00 7303.29 19436.79 26740.08

25 222.92 253.91 2.50 0.13 15003.73 11026.38 26030.11

26 222.92 253.91 2.50 0.64 7062.21 11026.38 18088.59

27 81.10 253.91 2.50 0.00 6231.08 4011.54 10242.61

28 81.10 237.14 1.30 0.09 8964.84 1115.87 10080.71

29 81.10 238.50 1.30 1.00 869.21 1115.87 1985.07

30 81.10 265.42 2.50 1.00 256.38 4011.54 4267.92

31 222.92 258.13 2.50 1.00 1063.50 11026.38 12089.87

32 222.92 289.34 5.10 1.00 51.34 19436.79 19488.13

33 322.79 285.07 5.10 1.00 164.53 28144.07 28308.60

34 322.79 301.12 7.20 1.00 8.59 33766.25 33774.84

35 442.70 297.41 7.20 1.00 0.77 46310.07 46310.85

36 119.91 287.34 7.20 1.00 42.97 12543.82 12586.78

37 99.87 275.45 5.10 1.00 155.30 8707.28 8862.58

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

107

38 141.82 253.91 2.50 1.00 831.14 7014.84 7845.99

39

Na

tura

l G

as

158.42 300.00 65.00 1.00 2.36 86893.82 86896.18

40 158.42 291.00 63.05 1.00 35.98 86316.80 86352.77

41 158.42 279.00 61.16 1.00 267.99 85738.21 86006.20

42 158.42 257.00 59.33 1.00 1360.82 85158.73 86519.55

43 158.42 240.00 57.55 1.00 3002.05 84575.73 87577.78

44 158.42 146.00 55.82 0.00 48032.84 83989.62 132022.46

45 158.42 134.00 54.15 0.00 55971.06 83404.49 139375.55

46 158.42 113.15 1.22 0.18 127682.63 3918.38 131601.01

47 28.71 113.15 1.22 1.00 5427.65 711.13 6138.77

LNG LNG 129.71 113.15 1.22 0.00 121075.46 3207.19 124282.65

Table B.15. Standard Molar Chemical Exergy Values for Selected Substances at Tref =

298.15K. Model II is referred.

Substance Formula �� �

(kJ/kmol)

Propane C3H8 (g) 2154000

Nitrogen N2 (g) 720

Methane CH4 (g) 831650

Ethane C2H6 (g) 1495840

Table B.16. Chemical exergy result for affected streams for C3MR

Stream 7 8 12 13 14 16 � � (GW) 13.96 3.39 3.39 13.96 10.57 10.57

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

108

Table B.17. Definition of the exergy of fuel and the exergy of product for the components of the C3MR process.

Component � , ��,

HEX1** � = �� + �� + �� − � �+ �� − � �+ �� − � �+ ��� − �� � �� = �� + ��

HEX2 � = ��� − �� � + �� − � � + �� − � �+ �� − � � �� = ��� − ���+ ��� − �� �

HEX3 � = ��� − �� � + �� − � �+ �� − � �+ �� − � � �� = ��� − ��� + ��� − �� �

HEX4 � = ��� − �� � + �� − � � + �� − � � + �� − � � �� = ��� − ���+ ��� − �� �

MLHE1 � = ��� − ���+ �� − � � + �� − � �+ �� − � �

+ �� − � �

�� = ��� − �� �+ ��� − ��� + ��� − �� �

MLHE2 � = ��� − �� �+ �� − � � + �� − � � + �� − � � �� = ��� − ��� + ��� − �� �

COMP1 � = ��� − �� �+ � � � �� = �� − � �

COMP2 � = ��� − �� �+ � � � �� = �� − � �

COMP3** � = �� + � � � �� = �� − � � + ��

COMP4 � = �� + � � � �� = �� − � � + ��

COMP5 � = �� + � � � �� = �� − � �+ ��

COMP6 � = � � � �� = ��� − �� �+ �� − � �

COMP7 � = � � � �� = ��� − �� �+ �� − � �

** Components crossing the environmental temperature

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

109

MIX1 � = � ∗ � + � ∗ ( � − � ) �� = � ∗ �

MIX2 � = � ∗ ( � − � ) �� = � ∗ ( � − � )

MIX3 � = � ∗ ( � − � ) �� = � ∗ ( � − � )

MIX4 � = � ∗ ( � − � ) + � ∗ ( � − �) + � ∗ ( − ) �� = � ∗ ( � − � ) + � ∗ ( � − �) + � ∗ ( − )

VALVE1 � = �� − � � �� = ��� − �� �

VALVE2 � = �� − � � �� = ��� − �� �

VALVE3 � = �� − � � �� = ��� − �� �

VALVE4** � = �� − � �+ �� �� = ��� �

VALVE5 � = �� − � � �� = ��� − �� �

VALVE6 � = �� − � � �� = ��� − �� �

Overall system

� ,������ = ������������� ��,������ = � − �

** Components crossing the environmental temperature

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

110

Table B.18. Exergy rate of product and fuel for the selected components of the C3MR process

Component

,

(MW)

�,

(MW)

,

(MW)

� (%)

� ,

(%)

HEX1 2.05 0.22 1.82 10.85 1.17

HEX2 3.47 1.36 2.1 39.29 1.35

HEX3 9.15 5.33 3.82 58.24 2.45

HEX4 9.32 6.66 2.66 71.47 1.7

MLHE1 112.44 99.12 13.32 88.16 8.53

MLHE2 14.92 11.69 3.23 78.35 2.07

COMP1 3.99 2.90 1.1 72.52 0.7

COMP2 11.67 8.41 3.26 72.05 2.09

COMP3 7.94 5.63 2.31 70.92 1.48

COMP4 21.43 15.44 5.99 72.07 3.83

COMP5 37.56 26.40 11.16 70.29 7.14

COMP6 36.11 25.68 10.43 71.12 6.68

COMP7 42.87 30.93 11.95 72.13 7.65

Total System 156.22 52.48 103.74*) 33.59 66.41

*) Throttling valves and mixers as well as coolers (dissipative components) are included

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

111

B.4 AP-XTM Process [51]

Fig. B.5. A general schematic on the AP-X process.

Table B.19. Composition for the AP-X process in molar percentage.

Component Formula Mole fraction (%)

Natural

Mixed Refrigerant Nitrogen N2 0.1 1.4

Methane CH4 86.0 34.3

Ethane C2H6 7.5 39.5

Propane C3H8 3.5 0.6

I-Butane C4H10 1.0 9.1

Butane C4H10 1.0 15.1

I-Pentane C5H12 0.3 -

Pentane C5H12 0.2 -

Hexane C6H14 0.4 -

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

112

Fig. B.6. Flowsheet for AP-XTM process.

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

113

Fig. B.7. �- T diagram for HEX0 (∆ ����ℎ = 0.3�)

Fig. B.8. �- T diagram for HEX1 (∆ ����ℎ = 6.4�)

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

114

Fig. B.9. �- T diagram for HEX2 (∆ ����ℎ = 2.1�)

Fig. B.10. �- T diagram for HEX3 (∆ ����ℎ = 17.7�)

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

115

Table B.20. Thermodynamic data for the material streams at real operating conditions.

Stream

Ma

teri

al

Str

ea

m

(kg/s) T

(°C) P

(bar) x

ET (MW)

EM (MW)

EPH (MW)

1 N

atu

ral

Ga

s 224.23 43.00 65.00 0.00 0.00 121.85 121.85

2 224.23 17.85 65.00 0.00 0.64 121.85 122.49

3 224.23 7.00 65.00 0.01 1.53 121.85 123.37

4 224.23 -16.15 65.00 0.06 4.97 121.85 126.82

5 224.23 -33.15 65.00 0.12 9.48 121.85 131.32

6 224.23 -85.00 65.00 1.00 43.05 121.85 164.90

7 224.23 -111.00 65.00 1.00 60.43 121.85 182.27

8 224.23 -166.00 4.51 1.00 188.16 44.95 233.12

9 224.23 -165.84 1.05 1.00 231.58 1.06 232.64

10 224.23 -165.84 1.05 1.00 231.58 1.06 232.64

11

97

Mix

ed

-re

frig

era

nt

7000.00 65.00 38.00 0.00 26.82 1913.78 1940.60

98 7000.00 60.00 38.00 0.02 19.76 1913.78 1933.54

99 7000.00 55.00 38.00 0.08 9.76 1913.78 1923.54

100 7000.00 45.00 38.00 0.20 0.47 1913.78 1914.25

101 5062.25 45.00 38.00 0.00 0.33 1575.49 1575.82

102 1937.75 45.00 38.00 1.00 0.37 292.91 293.29

103 1937.75 -55.00 38.00 1.00 90.45 293.05 383.50

104 1937.75 -54.27 10.00 0.99 128.59 240.94 369.53

105 1937.75 28.57 10.00 0.52 5.30 240.94 246.24

106 6792.45 15.69 10.00 0.06 36.51 1235.37 1271.87

107 7000.00 14.57 10.00 0.07 42.20 1278.02 1320.22

107A 7000.00 42.18 21.00 0.01 0.60 1659.77 1660.36

107B 7000.00 40.00 19.45 0.01 0.57 1622.41 1622.98

107C 6897.81 40.00 19.45 0.00 0.17 1608.24 1608.42

108 6897.81 84.46 38.00 0.00 54.91 1899.65 1954.56

109 102.19 40.00 19.45 1.00 0.00 10.75 10.75

110 102.19 41.74 38.00 1.00 0.00 11.16 11.16

111 7000.00 82.41 38.00 0.00 53.42 1913.78 1967.19

112 5062.25 -100.00 38.00 1.00 881.43 1575.49 2456.92

113 5062.25 -111.00 34.50 1.00 1041.92 1540.12 2582.04

114 4854.70 -110.95 33.46 1.00 1008.19 1466.11 2474.30

115 4854.70 -111.37 10.00 0.98 1441.86 989.74 2431.60

116 4854.70 -105.00 10.00 0.90 1286.21 989.74 2275.94

117 4854.70 28.57 10.00 0.00 3.31 989.74 993.05

118 207.55 -110.95 33.46 1.00 43.10 62.68 105.78

119 207.55 -111.37 10.00 0.98 61.64 42.31 103.96

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

116

120 207.55 -34.05 10.00 0.28 11.42 42.31 53.73

500

Pro

pa

ne

4000.00 17.58 7.90 0.89 117.69 467.50 585.19

501 4000.00 17.58 7.90 0.66 90.24 467.50 557.74

502 2658.31 17.58 7.90 1.00 87.65 310.69 398.33

503 2658.31 6.61 5.80 0.92 128.88 266.42 395.30

504 2658.31 6.61 5.80 0.78 110.63 266.42 377.05

505 2069.74 6.61 5.80 1.00 108.31 207.43 315.74

506 2069.74 -17.09 2.70 0.85 189.04 118.33 307.37

507 2069.74 -17.09 2.70 0.55 130.90 118.33 249.23

508 1145.35 -17.09 2.70 1.00 120.97 65.48 186.45

509 1145.35 -36.01 1.30 0.89 166.57 16.83 183.41

510 1145.35 -36.01 1.30 0.00 21.53 16.83 38.36

511 1145.35 20.49 5.00 0.00 1.68 105.44 107.12

512 2069.74 1.82 2.70 0.00 10.04 118.33 128.37

513 2069.74 61.49 10.00 0.00 2.05 267.27 269.32

514 2658.31 45.45 5.80 0.00 0.09 266.42 266.51

515 2658.31 91.57 15.00 0.00 20.12 394.49 414.61

516 4000.00 62.51 7.90 0.00 4.47 467.50 471.96

517 4000.00 109.12 20.00 0.00 100.52 598.40 698.93

518 4000.00 31.85 20.00 1.00 2.29 598.40 600.70

519 1341.69 17.58 7.90 0.00 2.62 156.81 159.43

520 588.57 6.61 5.80 0.00 2.34 58.99 61.33

521 924.39 -17.09 2.70 0.00 9.82 52.85 62.67

1000

Nit

rog

en

1510.00 -149.70 11.00 0.00 179.57 338.02 517.59

1001 1510.00 -34.05 11.00 0.00 18.16 338.02 356.17

1001A 1510.00 30.60 21.00 0.00 0.41 429.80 430.20

1001B 1510.00 40.00 21.00 0.00 0.02 429.80 429.82

1001C 1510.00 137.99 44.00 0.00 19.61 535.07 554.68

1001D 1510.00 40.00 44.00 0.00 0.02 535.07 535.10

1002 1510.00 93.10 67.00 0.00 6.09 595.27 601.36

1003 1510.00 43.00 67.00 0.00 0.00 595.27 595.27

1004 1510.00 -100.00 67.00 0.00 0.00 621.61 621.61

1005 1510.00 -168.04 11.00 0.02 255.74 338.02 593.76

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

117

Table B.21. Mole flow rate of the mixed refrigerant.

Component Nitrogen Methane Ethane Propane iButane Butane

Total stream (kmol/s)

100 3.05 74.79 86.13 1.31 19.84 32.93

101 2.96 70.69 72.80 0.92 11.13 16.67

102 0.09 4.10 13.33 0.39 8.71 16.25

103 0.09 4.10 13.33 0.39 8.71 16.25

104 0.09 4.10 13.33 0.39 8.71 16.25

105 0.09 4.10 13.33 0.39 8.71 16.25

106 2.93 71.89 83.14 1.27 19.39 32.24

107 3.05 74.79 86.13 1.31 19.84 32.93

107A 3.05 74.79 86.13 1.31 19.84 32.93

107B 3.05 74.79 86.13 1.31 19.84 32.93

107C 3.05 74.71 85.76 1.29 19.36 31.88

108 3.05 74.71 85.76 1.29 19.36 31.88

109 0.00 0.08 0.37 0.02 0.49 1.05

110 0.00 0.08 0.37 0.02 0.49 1.05

111 3.05 74.79 86.13 1.31 19.84 32.93

112 2.96 70.69 72.80 0.92 11.13 16.67

113 2.96 70.69 72.80 0.92 11.13 16.67

114 2.84 67.80 69.81 0.88 10.67 15.99

115 2.84 67.80 69.81 0.88 10.67 15.99

116 2.84 67.80 69.81 0.88 10.67 15.99

117 2.84 67.80 69.81 0.88 10.67 15.99

118 0.12 2.90 2.98 0.04 0.46 0.68

119 0.12 2.90 2.98 0.04 0.46 0.68

120 0.12 2.90 2.98 0.04 0.46 0.68

Vapor phase (kmol/s)

100 2.96 70.69 72.80 0.92 11.13 16.67

101 2.96 70.69 72.80 0.92 11.13 16.67

102 0.00 0.00 0.00 0.00 0.00 0.00

103 0.00 0.00 0.00 0.00 0.00 0.00

104 0.02 0.29 0.06 0.00 0.00 0.00

105 0.09 3.84 10.05 0.19 2.52 3.76

106 2.93 71.60 81.24 1.17 16.02 24.92

107 3.05 74.45 83.86 1.19 15.93 24.49

107A 3.05 74.68 85.62 1.29 19.20 31.56

107B 3.05 74.71 85.76 1.29 19.36 31.88

107C 3.05 74.71 85.76 1.29 19.36 31.88

108 3.05 74.71 85.76 1.29 19.36 31.88

109 0.00 0.00 0.00 0.00 0.00 0.00

110 0.00 0.00 0.00 0.00 0.00 0.00

111 3.05 74.79 86.13 1.31 19.84 32.93

112 0.00 0.00 0.00 0.00 0.00 0.00

113 0.00 0.00 0.00 0.00 0.00 0.00

114 0.00 0.00 0.00 0.00 0.00 0.00

115 0.60 2.42 0.04 0.00 0.00 0.00

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

118

116 1.85 14.74 0.40 0.00 0.00 0.00

117 2.84 67.80 69.81 0.88 10.67 15.99

118 0.00 0.00 0.00 0.00 0.00 0.00

119 0.03 0.10 0.00 0.00 0.00 0.00

120 0.12 2.80 2.14 0.01 0.05 0.05

Liquid phase (kmol/s)

100 0.09 4.10 13.33 0.39 8.71 16.25

101 0.00 0.00 0.00 0.00 0.00 0.00

102 0.09 4.10 13.33 0.39 8.71 16.25

103 0.09 4.10 13.33 0.39 8.71 16.25

104 0.07 3.80 13.27 0.39 8.71 16.25

105 0.00 0.25 3.28 0.20 6.19 12.50

106 0.00 0.29 1.90 0.10 3.36 7.32

107 0.01 0.34 2.27 0.12 3.91 8.44

107A 0.00 0.11 0.51 0.02 0.64 1.37

107B 0.00 0.08 0.37 0.02 0.49 1.05

107C 0.00 0.00 0.00 0.00 0.00 0.00

108 0.00 0.00 0.00 0.00 0.00 0.00

109 0.00 0.08 0.37 0.02 0.49 1.05

110 0.00 0.08 0.37 0.02 0.49 1.05

111 0.00 0.00 0.00 0.00 0.00 0.00

112 2.96 70.69 72.80 0.92 11.13 16.67

113 2.96 70.69 72.80 0.92 11.13 16.67

114 2.84 67.80 69.81 0.88 10.67 15.99

115 2.24 65.38 69.77 0.88 10.67 15.99

116 0.99 53.06 69.42 0.88 10.67 15.99

117 0.00 0.00 0.00 0.00 0.00 0.00

118 0.12 2.90 2.98 0.04 0.46 0.68

119 0.10 2.80 2.98 0.04 0.46 0.68

120 0.00 0.10 0.85 0.03 0.40 0.63

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

119

Table B.22. Definition of the exergy of fuel and the exergy of product for the selected components of the AP-X process.

Component kFE , kPE ,

COMP1 15115101, )( COMP

TT

COMPF WEEE +−= )( 5105111,

MM

COMPP EEE −=

COMP2 25135122, )( COMP

TT

COMPF WEEE +−= )( 5125132,

MM

COMPP EEE −=

COMP3 33, COMPCOMPF WE = )()( 5145155145153,

MMTT

COMPP EEEEE −+−=

COMP4 44, COMPCOMPF WE = )()( 5165175165174,

MMTT

COMPP EEEEE −+−=

EXP1 1100510041, )( EXP

MM

EXPF WEEE −−= )( 100410051,

TT

EXPP EEE −=

HEX0 )()( 1171161051040,

TTTT

HEXF EEEEE −+−= )()()( 101112102103560,

TTTTTT

HEXP EEEEEEE −+−+−=

HEX1 )()( 1131121161151,

MMTT

HEXF EEEEE −+−= )()( 112113671,

TTTT

HEXP EEEEE −+−=

HEX2 )()()( 8710001005100010052,

MMMMTT

HEXF EEEEEEE −+−+−= )( 782,

TT

HEXP EEE −=

HEX3 )()( 100110001201193,

TTTT

HEXF EEEEE −+−= )( 100310043,

TT

HEXP EEE −=

HPN2 22, HPNHPNF WE = )()( 10011002100110022,

M

D

MT

D

T

HPNP EEEEE −+−=

LPMR )( 107107,

T

A

T

LPMRLPMRF EEWE −+= )( 107107,

MM

ALPMRP EEE −=

LPN2 )( 1001100122,

T

A

T

LPNLPNF EEWE −+= )( 100110012,

MM

ALPNP EEE −=

MPMR MPMRMPMRF WE =, )()( 107108107108,

M

C

MT

C

T

MPMRP EEEEE −+−=

MPN2 22, MPNMPNF WE = )()( 10011001100110012,

M

B

M

C

T

B

T

CMPNP EEEEE −+−=

P1 )( 11010911,

TT

PPF EEWE −+= )( 1091101,

MM

PP EEE −=

PHX1 )( 5015001,

TT

PHXF EEE −= )()( 97111121,

TTTT

PHXP EEEEE −+−=

PHX2 )( 5045032,

TT

PHXF EEE −= )()( 9897231,

TTTT

PHXP EEEEE −+−=

PHX3 )( 5075063,

TT

PHXF EEE −= )()( 9998341,

TTTT

PHXP EEEEE −+−=

PHX4 )( 5105094,

TT

PHXF EEE −= )()( 10099451,

TTTT

PHXP EEEEE −+−=

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

120

PSEP1 )( 5015011,

MT

PSEPF EEE +=

)()( 5195025195021,

MMTT

PSEPP EEEEE +++=

PSEP2 )( 5045042,

MT

PSEPF EEE +=

)()( 5205055205052,

MMTT

PSEPP EEEEE +++=

PSEP3 )( 5075073,

MT

PSEPF EEE +=

)()( 5215085215082,

MMTT

PSEPP EEEEE +++=

PV1 )( 5035021,

MM

PVF EEE −=

)( 5025031,

TT

PVP EEE −=

PV2 )( 5065052,

MM

PVF EEE −=

)( 5055062,

TT

PVP EEE −=

PV3 )( 5095083,

MM

PVF EEE −=

)( 5085093,

TT

PVP EEE −=

PV4 )( 5005184,

MM

PVF EEE −=

)( 5185004,

TT

PVP EEE −=

SEP2 )( 1001002,

MT

SEPF EEE +=

)()( 1021011021012,

MMTT

SEPP EEEEE +++=

SEP3 )( 1071073,

M

B

T

BSEPF EEE +=

)()( 1091071091073,

MM

C

TT

CSEPP EEEEE +++=

V0 )( 1041030,

MM

VF EEE −=

)( 1031040,

TT

VP EEE −=

V1 )( 1151141,

MM

VF EEE −=

)( 1141151,

TT

VP EEE −=

V2 )( 1191182,

MM

VF EEE −=

)( 1181192,

TT

VP EEE −=

V3 )( 983,

MM

VF EEE −=

)( 893,

TT

VP EEE −=

Total System

� ,������ = ������������� ��,������ = � − �

Note: Dissipative components are mixers, coolers, after- and intercoolers. Separator SEP1 has no flash product thus not considered.

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

121

Table B.23. Exergy rate of product and fuel for the AP-X process.

Component

,

(MW)

�,

(MW)

,

(MW)

� (%)

COMP1 120.5 88.6 31.9 74

COMP2 206.5 148.9 57.5 72

COMP3 223.8 128.1 95.8 57

COMP4 404.9 130.9 274.0 32

EXP1 283.6 188.8 94.8 67

HEX0 1406.2 1004.7 401.4 71

HEX1 191.0 177.9 13.2 93

HEX2 153.1 127.7 25.3 83

HEX3 211.6 0.0 211.6 0

HPN2 100.4 60.2 40.2 60

LPMR 546.0 381.8 164.2 70

LPN2 130.0 91.8 38.2 71

MPMR 550.1 291.4 258.7 53

MPN2 192.6 105.3 87.3 55

P1 0.5 0.4 0.0 89

PHX1 27.5 27.2 0.2 99

PHX2 18.2 7.9 10.3 44

PHX3 58.1 13.4 44.7 23

PHX4 145.0 13.8 131.2 10

Overall System 2024.7 111.27 2013.41*) 6

*) Throttling valves and mixers as well as coolers (dissipative components) are included

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

122

Table B.24. Power net required by AP-X process components.

Components Power (MW)

COMP1 100.70

COMP2 198.48

COMP3 203.82

COMP4 308.88

EXP1 -66.93

HPN2 94.35

LPMR 504.36

LPN2 112.23

MPMR 495.35

MPN2 173.00

P1 0.45

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

123

As of 2013, the AP-XTM process has been utilized in the liquefaction plant at 12% in capacity

worldwide, as per Fig. 1.4. This coverage does not include Malaysia. However, PETRONAS’

new floating LNG (PFLNG 1) process namely AP-NTM uses AP-XTM as its reference design

[105].

Fig. B.11. PETRONAS FLNG to be commissioned in 2015 [15].

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

124

B.5 MR-X Process [56,94]

Table B.25. Composition of NG and refrigerants for MR-X process.

Component Formula

Mole Fraction (%)

Natural Gas

Pre-cooling

Liquefaction Sub-

cooling

Nitrogen N2 0.1

1.4 100.0

Methane CH4 86.0

34.3

Ethane C2H6 7.5 45.5 39.5

Propane C3H8 3.5 4.9 0.6

I-Butane C4H10-2 1.0

9.1

Butane C4H10-1 1.0 49.6 15.1

I-Pentane C5H12-2 0.3

Pentane C5H12-1 0.2

Hexane C6H14-1 0.4

Legend:

natural gas L N G mixed-refrigerant nitrogen for pre-cooling block for sub-cooling block mixed refrigerant gas phase liquid phase of mixed refrigerant of mixed refrigerant

Fig. B.12. A general schematic on the MR-X process.

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

125

V1

HEX2

EXP1

V3

V2

HEX3

AC

SEP2

HEX0

V0

HEX1

SEP1

P1

MIX3

MIX2

MIX1

S1

HPN2

LPMR

ICMR

SEP3

MPMR

IC2

LPN2 MPN2

PHE2

PHE1

PS1

PV1

PM1

PV2

PC1

PAC

IC1

PC2

114

115

5

10056

1000 1004

7

118

119

1003

1001

120

1002

100

101

102

3

104

116

105117

103

1124

113

9

8

109

110

108

106

107

1001D

107A

107B

107C

1001C1001A

204

2

207

206

208

205

1

201

202

210

203209 211

1001B

111

212

99

NG

Flash gas

COLDBOX

LNG

Legend:

NG-LNG

Mixed-refrigerant for main

liquefaction section

Mixed-refrigerant for pre-

cooling section

Nitrogen for sub-cooling

section

Fig. B.13 Process flow diagram for MR-X process

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

126

Fig. B.14. Cumulative cooling curves for the MR-X process.

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

127

Table B.26. Thermodynamic data for the material streams (at real operating conditions).

Stream Material

stream

m p T TE

ME

PHE

(kg/s) (bar) (ºC) (MW) (MW) (MW)

1

NG

-LN

G

224.22

65.00 43.00 0.00 121.84 121.84

2 65.00 0.00 2.30 121.84 124.14

3 65.00 -33.15 9.46 121.84 131.31

4 63.05 -85.00 43.46 121.02 164.49

5 61.16 -111.00 61.37 120.22 181.59

6 59.32 -168.00 119.45 119.40 238.85

7 1.05 -165.21 229.77 1.06 230.83

8 1.05 -165.21 229.77 1.06 230.83

99

Ma

in M

ixe

d R

efr

ige

ran

t

958.33 50.00 60.00 2.31 275.11 277.42

100 958.33 50.00 50.00 0.36 275.11 275.47

101 638.36 50.00 50.00 0.29 212.79 213.08

102 319.98 50.00 50.00 0.06 56.91 56.97

103 319.98 48.50 -55.00 15.05 56.79 71.85

104 319.98 10.00 -57.17 25.17 43.23 68.40

105 319.98 10.00 41.84 0.02 43.23 43.25

106 932.16 10.00 33.18 0.32 169.62 169.94

107 958.33 10.00 32.99 0.30 174.97 175.27

107A 958.33 30.00 104.04 12.52 249.29 261.82

107B 958.33 30.00 43.00 0.02 249.29 249.31

107C 769.96 30.00 43.00 0.02 219.20 219.22

108 769.96 50.00 80.24 4.74 244.79 249.53

109 188.37 30.00 43.00 0.00 24.79 24.79

110 188.37 50.00 45.14 0.02 25.63 25.64

111 958.33 50.00 64.89 3.97 275.11 279.08

112 638.36 48.50 -100.00 102.03 211.54 313.57

113 638.36 34.50 -111.00 130.66 196.40 327.06

114 612.18 34.50 -111.00 125.31 188.34 313.65

115 612.18 10.00 -111.75 181.88 126.10 307.98

116 612.18 10.00 -100.02 145.94 126.10 272.04

117 612.18 10.00 41.84 0.01 126.10 126.11

118 26.17 34.50 -111.00 5.36 8.05 13.41

119 26.17 10.00 -111.75 7.78 5.39 13.17

120 26.17 10.00 26.55 0.02 5.39 5.41

1000

Nit

rog

en

649.17

11.00 -118.48 46.10 145.32 191.42

1001 10.67 26.55 0.30 143.46 143.77

1001A 21.00 115.16 4.96 184.77 189.73

1001B 21.00 43.00 0.00 184.77 184.77

A p p e n d i x B . E n e r g y a n d E x e r g y A n a l y s e s – D a t a , F l o w a n d R e s u l t s

128

1001C 44.00 145.93 9.76 230.03 239.79

1001D 44.00 43.00 0.00 230.03 230.03

1002 67.00 98.75 3.20 255.91 259.12

1003 67.00 43.00 0.00 255.91 255.91

1004 64.99 -100.00 43.01 254.03 297.04

1005 11.00 -168.04 107.81 145.32 253.13

201

Pre

-co

oli

ng

Mix

ed

Re

frig

era

nt

500.00 19.00 43.00 0.00 71.90 71.90

202 500.00 19.00 0.00 5.10 71.90 77.00

203 305.00 19.00 0.00 3.11 43.86 46.97

204 195.00 19.00 0.00 1.99 28.04 30.03

205 305.00 8.83 -5.33 9.36 36.82 46.18

206 195.00 19.00 -33.15 5.66 28.04 33.70

207 195.00 3.51 -38.82 19.03 13.87 32.90

208 195.00 3.51 26.35 0.16 13.87 14.03

209 195.00 9.00 67.67 0.43 23.71 24.14

210 305.00 8.83 63.77 0.49 36.82 37.30

211 500.00 8.83 65.22 0.85 60.36 61.20

212 500.00 19.00 102.14 11.94 71.90 83.84

129

Appendix C. Economic Analysis

Data and Flow – A Study Case on

the C3MR Process

A p p e n d i x C . E c o n o m i c A n a l y s i s D a t a a n d F l o w – A S t u d y C a s e o n t h e C 3 M R P r o c e s s

130

C.1 Purchased Equipment Costs (PEC) Estimates

Levelized values of total capital investment (TCI) comprising purchase equipment costs, fuel

costs, and operating and maintenance expenditures for the life span of the liquefaction plant

is assumed to be 20 years (2016-2036). Annually, the average plant capacity factor is

assumed to be 85 %, which means it operates 7446 hrs out of total 8760 hrs available. Table

C.1 presents all considered parameters for the analysis.

Table C.1. Parameters and assumptions used in TRR calculations [62]

Parameter (units) Value

1a. Average general inflation rate (1994-2017) (%) 5.0

b. Average nominal escalation rate of all costs (except fuel) (1994-2017) (%) 5.0

c. Average nominal escalation rate of electricity (1994-2017) (%) 6.0

2a. Beginning of the design and construction period Jan.1, 2014

b. Date of commercial operation Jan.1, 2016

3a. Plant economic life (years) 20

b. Plant life for tax purposes (years) 15

4. Plant financing fractions and required returns on capital:

Common Preferred

Type of financing Equity Stock Debt

Financing fraction (%) 35.0 15.0 50.0

Required annual return (%) 15.0 11.7 10.0

Resulting average cost of money (%) 12.0

5a. Average combined income tax rate (1994-2017) (%) 38

b. Average property tax rate (1994-2017) [% of PFI (in end-2015 dollars)] 1.5

c. Average insurance rate (1994-2017) [% of PFI (in end-2015 dollars)] 0.5

6. Average capacity factor (%) 85

7. Labor positions for operating and maintenance 30

8. Average labor rate ($/h) 2.24

9. Annual fixed operating and maintenance costs at full capacity (106 $) 0.336

10. Annual variable operating and maintenance costs at full capacity (106 $) 0.031

11. Unit cost of fuel (cent/kWh) 25

12. Allocation of plant facilities investment to the individual years

of design and construction (%)

Jan. 1-Dec.31, 2014 40

Jan. 1-Dec.31, 2015 60

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C.1.1 Heat Exchangers

It is important to size the C3MR heat exchanger before proceeding to its cost estimation. The

heat transfer area of the heat exchanger is used as the sizing parameter. The heat

exchangers’ UA values are obtained from the AspenPlus simulation results (U - the overall

heat transfer coefficient, and A - heat transfer area). Due to the calculation complexity

necessary for estimating the U, these values are assumed from literature sources

considering the heat transfer areas for each heat exchanger is calculated, and the type of the

fluids in contact with each other, as shown in Table C.2 [82,101]. Applying the sizing

parameter, the PEC for heat exchangers is estimated based on equation (3.20), as shown in

Table C.3.

Table C.2. U and A values of the liquefaction heat exchangers [101].

Component HEX1 HEX2 HEX3 HEX4 MLHE1 MLHE2

U [W/m2K] 1000 1200 1300 1500 2500 1700

A [m2] 2810 2785 3262 2379 8389 329

Table C.3. The purchased equipment cost of liquefaction heat exchangers (106$).

Component HEX1 HEX2 HEX3 HEX4 MLHE1 MLHE2

PEC [2012] 42.36 42.14 46.19 38.46 79.90 12.20

C.1.2 Dissipative coolers

Unlike the primary heat exchangers, the sizing of these dissipative coolers is based on the

assumption that they are standard equipment. The type as traditional fixed head shell-and-

tube heat exchanger is opted. The AspenPlus simulation provides coolers’ UA values, and the

overall heat transfer (U) values are assumed from literature considering the condition of

heat transfer (Table C.4). For the PEC calculation, equation (C.1) derived from the estimating

charts is applied before proceeding with equation (C.2). The use of correlation instead of

chart avoids reading inaccuracy. The cost calculated from this equation is in 2000 with

CEPCI of 394, and then the cost in 2012 is calculated with the aid of the equation (3.20) [70].

FP, FM and FL represent the shell-side pressure factor, material factor and correction of tube-

length, respectively. Giving to these definitions, the material of construction is selected as

carbon steel/stainless steel. Calculation from equation (C.3) is continued by using the

coefficients a and b for selected materials [70]. The pressure factor is calculated from

equation (C.4) by considering the shell-side pressure. FL value is assumed as 1. By

considering all the stated effects, the total PECs of the coolers are as per Table C.5.

A p p e n d i x C . E c o n o m i c A n a l y s i s D a t a a n d F l o w – A S t u d y C a s e o n t h e C 3 M R P r o c e s s

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Table C.4. U and A values of dissipative coolers [106].

Component COOLER COOLER2 COOLER3

U [W/m2K] 1200 500 500

A [m2] 10175 3069 3972

[ ] [ ]{ }2)ln(09861.0)ln(9228.00545.11exp AACB +−= (C.1)

BLMP CFFFPEC = (C.2)

b

M

AaF

+=100

(C.3)

2

1000017.0

100018.09803.0

+

+=

PPFP

(C.4)

Table C.5. Purchase equipment cost of dissipative coolers (106$).

Component COOLER COOLER2 COOLER3

PEC [2012] 3.69 1.00 1.49

C.1.3 Propane and mixed refrigerant compressors

The compressor is the second most important equipment for example in the C3MR process.

The type is centrifugal turbo, and electric motors drive it. For cost calculations, net required

power (PC) is used as size factor for each compressor (Table C.6). The limitation of the range

of the size factor situation mentioned in coolers for the cost calculations is also valid for the

compressors. The purchase equipment cost correlation is expressed by the equation (C.5).

FD and FM represent the effect of driving type and material used for the construction,

respectively. FD =1 for the electric motor drives and FM = 2.5 for stainless steel. The base

compressor purchase cost for a CEPCI of 394 in 2009 is given by the equation (C.6) [70]. The

final values for compressors and drivers purchased equipment costs are calculated by

taking all the mentioned effects and the CEPCI of 584.6 in 2012 into consideration. These

are shown in Table C.7 [83].

BMD CFFPEC = (C.5)

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Table C.6. The process work input (indicated and net required).

Component Indicated Power, PC

(MW) Net Required Power

(MW)

COMP1 3.04 3.38

COMP2 9.59 10.66

COMP3 6.99 7.77

COMP4 19.28 21.42

COMP5 30.60 34.00

COMP6 32.49 36.10

COMP7 38.58 42.87

TOTAL 140.57 156.20

( )[ ]{ }CB PlnexpC 80.02223.7 += (C.6)

Table C.7. The purchased equipment cost of the compressors (106 $).

Compressor COMP1 COMP2 COMP3 COMP4 COMP5 COM6 COMP7

PEC [2012] 4.28 10.72 8.32 18.73 33.77 35.43 40.64

C.1.4 Separators

Process phase separators are designed and sized as vertically oriented columns containing

little or no internals. Such design finds many different application areas such as flash drums,

reflux drums, storage drums, mixing vessels and chemical reactors. The separators

purchased equipment cost at CEPCI =394 is based on the weight of the shell and two 2:1

elliptical heads (W). The heads include an allowance for ladders, platforms, manholes and a

nominal number of nozzles. Its PEC is given by equation (C.7). FM, CV, CPL account for the

material factor (1.20 for low alloy steel), cost of empty vessel, and cost of extra equipment

(such as nozzles, manholes, ladders and supports), respectively. The cost of empty vessel

and heads for vertical orientation is given by the correlation equation (C.8). On sizing this

equipment, the ratio between length (L) and diameter of the vessel (D), L/D is assumed as 4.

Assuming the holdup time is at 2 minutes, the volumetric flow rate determines the volume

of the vessel.

PLVM CCFPEC += (C.7)

( )[ ] ( )[ ]{ }2

V Wln02297.0Wln18255.0775.6expC ++= (C.8)

A p p e n d i x C . E c o n o m i c A n a l y s i s D a t a a n d F l o w – A S t u d y C a s e o n t h e C 3 M R P r o c e s s

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Table C.8 Sizing parameters of the separators.

Parameter Unit SEPA1 SEPA2 SEPA3 SEPA4 SEPA5

Flow rate m3/s 8.42 9.51 25.59 1.95 12.07

Holdup time s 120 120 120 120 120

Volume m3 1011 1141 3070 233 1448

Diameter (D) m 6.85 7.13 9.92 4.20 7.72

Length (L) m 27 28 39 16 31

Table C.8 illustrates the sizing results. The added cost CPL, (platforms and ladders for the

vessels) depends on the diameter (D) and the length (L) which expressed by equation (C.9).

The weight (W) of the vessel or tower depends on the wall thickness of the shell and the two

heads. For cost estimation purposes, it is sufficient to assume shell thickness equal to the

head thickness. Then, the weight of the shell and two heads is estimated per equation (C.10).

( ) ( ) 80161.063316.01.237 LDCPL = (C.9)

( )( ) ρtDLtDπW SS 0.8++= (C.10)

where the term L accounts for the cylinder part, the term 0.8 D accounts for the two heads, ρ

represents the density of the carbon steel and tS the shell thickness. The effects of the

corrosion, wind and earthquake are excluded from the calculation and, in this case, shell

thickness is calculated as

d

dS

PSE

DPt

1.22 −= (C.11)

Table C.9 Purchased equipment cost of separators (106$).

Component SEPA1 SEPA2 SEPA3 SEPA4 SEPA5

PEC [2012] 0.911 0.798 1.06 0.987 0.419

where D = shell diameter, Pd = internal design pressure, S = maximum allowable stress of

the shell material at the design temperature and E = fractional weld efficiency. All the

required parameters are selected from the literature, and the weight of the vessel is

calculated. The final results of the purchased equipment cost of the separators are

illustrated in Table C.9 [100].

C.1.5 Valves and mixers

The purchased equipment costs for the mixers are assumed to be zero. Although there are

mixers in the simulation flow sheet of the process, in the real plant they are designed as the

A p p e n d i x C . E c o n o m i c A n a l y s i s D a t a a n d F l o w – A S t u d y C a s e o n t h e C 3 M R P r o c e s s

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junction of two different pipelines. Additionally, the PEC of the valves, based on the flow

rates of the propane and mixed refrigerant, are obtained through a personal interview [83]

since they are considered significant capacity components. After obtaining the base cost and

the base sizing factor as flow rate, the final valves PECs are calculated using equation (3.20).

The final PEC values for the valves are shown in Table C.10.

Table C.10 Purchased equipment cost of throttling valves (106$).

Component VALVE1 VALVE2 VALVE3 VALVE4 VALVE5 VALVE6

PEC [2012] 0.603 0.471 0.239 0.745 0.464 0.243

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C.2 Estimation of Total Capital Investment

The total PEC is 426.26*106 $ for the year 2012 inclusive the appropriate cost indexes.

Referring the percentages of the total PEC (Table 5.1), the fixed capital investment (FCI) cost

for the plant is estimated at 1.364*106 $. The total capital investment (TCI)6 is calculated at

1.459*106 $. The total capital investment breakdown is given in Table 3.1. The evaluation of

other outlays is given in detail in this section.

C.2.1 Calculation of startup costs (SUC) and working capital (WC)

The cost required for the plant startup is a part of other outlays. After calculating the fixed

capital investment, startup costs and the working capital are to be estimated. The startup

costs are the sum of the unescalated expenses, such as one month of fixed O&M costs, one

month of variable O&M costs at full load, one week of full load fuel and 2% of the plant

facilities investment:

)10*021.1364)(02.0(52

)10*163.98(

12

)10*031.0(

12

)10*336.0(SUC 6

666

2012 1,Jan +++=

After the escalation of the cost to the end of December 31, 2015

6

2015 31, Dec 10*49.35SUC = $

The working capital is the sum of the unescalated expenses of 2 months of fuel cost and

variable O&M costs at full load and 3 months of labor cost, plus a contingency at 25 % of

those said costs:

)25.1(4

10*154.0

6

10*163.98WC

66

2012 1,Jan

+

=

After the escalation of the cost to the end of December 31, 2015

64

2012 1,Jan 2015 31, Dec 10*916.24)1(*WCWC =+= ni $

6 defined as the sum of the fixed capital investment and other outlays

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Table C.11. Estimation of the total capital investment.

I. Fixed Capital Investment (FCI) 106 $

A. Direct costs (DC)

1. Onsite costs (ONSC)

Total purchased equipment cost (PEC) 426.26

PEC installation (45% of PEC) 191.81

Piping (35% of PEC) 149.19

Instrumentation and control (20 % of PEC) 85.25

Electrical equipment and materials (20% of PEC) 85.25

Total onsite costs 937.76

2. Offsite costs (OFSC)

Land (10% of PEC) 0

Civil, structural and architectural work (50% of PEC) 0

Service facilities (65% of PEC) 0

Total offsite costs 0

Total direct costs 937.76

B. Indirect costs (IC)

1. Engineering and supervision (35% of PEC) 149.19

2. Construction costs (15% of DC) 140.66

3. Contingencies (10% of FCI) 136.40

Total indirect costs 426.26

Fixed-capital investment (FCI) 1364.02

II. Other Outlays (OO)

A. Startup costs 29.20

B. Working capital 20.50

C. Cost of licensing, R&D 0

D. Allowance for funds used during construction (AFUD) 45.89

Total other outlays 95.59

Total capital investment (TCI) 1459.61

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C.2.2 Estimation of allowance for funds used during construction

(AFUDC)

The plant facilities investment (PFI) is provided in the equation (C.12). The land cost is

excluded from the economic analysis thus the PFI is equal to the plant fixed capital

investment. With an annual rate of 5 % to December 31, 2014, 40% of the plant facilities

investment should be escalated, and the rest of it (which accounts 60%) should be escalated

to December 31, 2015. This is according to the parameters assumed in the economic

analysis.

LandJan.1,2012Jan.1,2012 CostFCIPFI −= (C.12)

The allowance for funds used during construction for the liquefaction plant is calculated

separately and listed in Table C.12. It is based on the parameters assumed for the plant

financing fractions and required returns on capital. The total AFUDC calculated is

72.214*106 $ at the end of the year 2015. Consequently, the evaluations of the total non-

depreciable (TNI)7 in the equation (C.13) and depreciable capital investments (TDI) in

equation (C.14) for the liquefaction plant are carried out.

AFUDCEquityCommon++= WCCostTNI Land (C.13)

6

5Dec.31.201 10*497.56TNI =

Assessing the total depreciable capital investment as

TNITCITDI −= (C.14)

6

5Dec.31.201 10*115.1403TDI = $.

After the total depreciable investment (TDI) calculation, the modified accelerated cost

recovery system (MACRS) factors can be estimated for the tax life of the system which is 15

years. It is significant to note that, according to MACRS the depreciation is calculated for one

more extra years. Therefore, for a system that has the tax life of 15 years, the depreciation

should be calculated for 16 years (Table C.13) [62].

7 defined as the sum of the land cost, working capital and common equity of AFUDC.

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Table C.12. The calculated values for the allowance for funds used during construction (106 $).

Plant-Facilities Investment Common Equity Preferred Equity Debt

Construction

Year

Calendar

Year

In Jan.1,2012

Dollars

Amount of

Escalation

Escalated

Investment

Escalated

Investment

AFUDC

Escalated

Investment

AFUDC

Escalated

Investment

AFUDC

1 2014 545.609 55.925 601.533 210.537 31.581 90.230 10.557 300.767 30.077

2 2015 818.413 129.002 947.415 331.595 0.000 142.112 0.000 473.708 0.000

Subtotals 1364.021 184.927 1548.948 542.132 31.581 232.342 10.557 774.474 30.077

Total AFUDC 72.214

Total AFUDC in 2012 45.893

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Table C.13. Statutory percentages for use in the MACRS for a life period of 15 years, annual

tax depreciation and tax book at the end of each year for the LNG plant.

Year of Commercial Operation

Calendar

Year

MACRS Depreciation

Factor (%)

Annual Tax Depreciation

(106$)

End-Year Tax Book Value

(106$)

0 2015 - - 1403.115

1 2016 5.00 70.156 1332.960

2 2017 9.50 133.296 1199.664

3 2018 8.55 119.966 1079.697

4 2019 7.70 108.040 971.657

5 2020 6.93 97.236 874.421

6 2021 6.23 87.414 787.007

7 2022 5.90 82.784 704.224

8 2023 5.90 82.784 621.440

9 2024 5.91 82.924 538.516

10 2025 5.90 82.784 455.732

11 2026 5.91 82.924 372.808

12 2027 5.90 82.784 290.024

13 2028 5.91 82.924 207.100

14 2029 5.90 82.784 124.316

15 2030 5.91 82.924 41.392

16 2031 2.95 41.392 0.000

Totals 100.00 1403.115

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C.3 Estimation of Operating and Maintenance (O&M)

Costs

The operating and maintenance costs are defined as the total of the fixed and variable costs.

The compounds of the fixed O&M costs are operating labor, maintenance labor, maintenance

materials, overhead, administration and support, distribution and marketing and so on. For

the maintenance and operation of the plant, 30 labor positions are expected with an average

labor rate of 2.24 $ per hour assumed for Malaysia. The labor cost value is achieved for the

year 2009 and escalated to 2011. The average working hours for a labor position is

considered as 2080 hrs per annum, and the yearly direct labor cost is evaluated as

0.154*106 $. Based on this evaluation, the yearly fixed O&M costs and the yearly variable

O&M costs at full capacity are valued at 0.336*106 $ and 0.031*106 $, respectively.

Considering the capacity factor of the plant, these values are calculated as 0.286*106 $ and

0.026*106 $, respectively [107].

To calculate the corresponding costs for the first year of the operation, these fixed and

variable O&M costs are escalated at a nominal escalation rate of 5 % per year to the Jan.1,

2016.

$10*364.0)1(* 65 =+= nfixedJan.1,2011fixedJan.1,2016 iM&OM&O

$10*033.0)1(* 65 =+= nvariableJan.1,2011variableJan.1,2016 iM&OM&O

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C.4 Estimation of the Fuel Costs (FC)

Fuel costs are typically a part of operating and maintenance costs. Nevertheless, they are

determined independently in this analysis because of their importance for the thermal

systems. In LNG plants, the fuel of the system is normally the electricity supplied to the

propane and mixed refrigerant compressors. The unit electricity price in Malaysia is 25

Malaysian cents per kWh for the year 2011. The fuel price is then calculated as 98.162*106 $

in view of the working hours 7446 hrs in a year with 85 % of plant availability and the total

power consumption of the compressors as 156.22 MW. The calculated value is escalated to

the Jan.1, 2016 with an escalation rate of 6 % as [108].

65 10*364.131)1(* =+= nJan.1,2011Jan.1,2016 iFCFC $

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C.5 Estimation of Revenue Requirements

The annual total revenue requirement (TRR) of a thermal plant is defined as the income that

must be gained in a specified year. It should come from the trade of all the yields of the plant

to pay the system operating business for the expenses acquired in the same specified year.

This is also to ensure plant operation is economically sound. TRR is assessed as the total of

the eight yearly amounts: total capital recovery (TCR); minimum return on investment (ROI)

for common equity (ce), preferred stock (ps) and debt (d); income taxes (ITX); other taxes

and insurance (OTXI); fuel costs (FC); and operating and maintenance costs (O&M). Equation

(C.15) states such as follows [62]:

jjjjdj,psj,cej,jj M&OFCOTXIITXROIROIROITCRTRR +++++++= (C.15)

(9) (1) (2) (3) (4) (5) (6) (7) (8)

The year-by-year evaluation is executed to evaluate the levelized revenue requirement. It is

shown in Table C.14.

Meanwhile, the annuities and the cost levelization involved in the revenue requirements

estimates will refer to the following correlations ((C.16) to (C.17)) and ((C.18) to (C.19))

[62,83], respectively.

( )( )neffeff

n

eff

ii

iAP

+

−+=

1

11

(C.16)

( )( ) 11

1

−+

+==

n

eff

n

effeff

i

ii

P

ACRF

(C.17)

( )( ) CRF

k

kkFCCELFFCFC

FC

FCn

FC00L −

−×=×=

1

1

(C.18)

( )( ) CRF

k

kkOMCCELFOMCOMC

OMC

OMCn

OMC00L −

−×=×=

1

1

(C.19)

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Table C.14. Year-by-year revenue requirement breakdown of the LNG liquefaction plant (106 $)

Year

Calendar Year

(1)

Capital Recovery

(2) Return on Common

Equity

(3) Preferred

Stock Dividends

(4)

Interest on Dept

(5)

Income Taxes

(6) Other

Taxes and Insurance

(7)

Fuel Cost

(8)

O&M Cost

(9) Total

Revenue Requirement

1 2016 73.766 76.630 25.616 72.981 64.880 33.160 131.364 0.398 478.794

2 2017 97.760 72.680 24.335 69.332 37.681 33.160 139.246 0.418 474.611

3 2018 92.694 67.471 22.633 64.483 38.510 33.160 147.600 0.439 466.991

4 2019 88.162 62.528 21.020 59.887 39.024 33.160 156.457 0.461 460.700

5 2020 84.057 57.823 19.487 55.519 39.306 33.160 165.844 0.484 455.679

6 2021 80.325 53.333 18.026 51.355 39.391 33.160 175.795 0.508 451.892

7 2022 78.565 49.040 16.630 47.378 37.663 33.160 186.342 0.533 449.311

8 2023 78.565 44.838 15.265 43.489 34.252 33.160 197.523 0.560 447.651

9 2024 78.618 40.637 13.900 39.600 30.787 33.160 209.374 0.588 446.664

10 2025 78.565 36.433 12.534 35.709 27.426 33.160 221.937 0.617 446.380

11 2026 78.618 32.232 11.169 31.820 23.961 33.160 235.253 0.648 446.860

12 2027 78.565 28.028 9.803 27.928 20.601 33.160 249.368 0.681 448.132

13 2028 78.618 23.826 8.438 24.039 17.136 33.160 264.330 0.715 450.261

14 2029 78.565 19.622 7.072 20.147 13.775 33.160 280.190 0.750 453.281

15 2030 78.618 15.421 5.707 16.258 10.310 33.160 297.001 0.788 457.263

16 2031 62.836 11.217 4.341 12.367 22.679 33.160 314.821 0.827 462.247

17 2032 47.107 7.841 3.252 9.264 35.671 33.160 333.711 0.869 470.875

18 2033 47.107 5.292 2.439 6.948 33.610 33.160 353.733 0.912 483.201

29 2034 47.107 2.742 1.626 4.632 31.549 33.160 374.957 0.958 496.731

20 2035 47.107 0.192 0.813 2.316 29.488 33.160 397.455 1.005 511.536

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C.5.1 Total capital recovery

The total net investment of the system must be recovered during the assumed economic life

of the plant. Calculations of the total capital recovery of the natural gas liquefaction plant are

conducted by year by year analysis and illustrated in Table C.15. According to this table the

column D (total capital recovery) is previously calculated and illustrated in total revenue

requirement in Table C.14. The book life (BL) and the total depreciable investment (TDI) are

used to calculate the annual book depreciation (BD), shown in the column A of Table C.15 by

applying the straight-line method described as [62].

BL

TDIBD j = , j =1, ..., BL (C.20)

where the index j represents the j-th year of the book life. The deferred income taxes (DITX)

for the j-th year of tax life (TL) is illustrated in column B of Table C.15 which are coming

from the difference between the annual tax depreciation (TXD), which is expressed by the

equation (C.21) and using the MACRS method. The annual book depreciation (BD) is

calculated by the equation (C.22), where the term t denotes the average combined income

tax rate in Table C.1 [62].

jMACRS,j f*TDITDX = , j = 1, …, TL+1 (C.21)

( ) t*BDTXDDITX jjj −= , j = 1, …, TL+1 (C.22)

BL

CEAFRCEAF j = , j =1, ..., BL

(C.23)

The common equity allowance for funds used during the construction (CEAF) that is not

considered in the net depreciable investment is recovered by applying a constant annual

amount (RCEAF). It is illustrated in the column C of the Table C.15 and calculated by the

equation (C.23), where CEAF represents the common equity AFUDC at the end of the design

and construction period [62].

The total capital recovery (TCR) for the j-th year of book life is the sum of the book

depreciation, deferred income taxes and recovery of the common equity AFUDC as shown as

jjjj RCEAFDITXBDTCR ++= , j = 1, …, BL (C.24)

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Table C.15. Year by year capital recovery schedule for the LNG plant. (106 $)

Year of

Commercial Operation

Calendar

Year

(A) Annual Book Depreciation

(B) Deferred

Income Taxes

(C) Recovery of

Common Equity AFUDC

(D) Total

Capital Recovery

1 2016 70.156 0.000 3.611 73.766

2 2017 70.156 23.993 3.611 97.760

3 2018 70.156 18.928 3.611 92.694 4 2019 70.156 14.396 3.611 88.162 5 2020 70.156 10.290 3.611 84.057 6 2021 70.156 6.558 3.611 80.325

7 2022 70.156 4.799 3.611 78.565

8 2023 70.156 4.799 3.611 78.565

9 2024 70.156 4.852 3.611 78.618

10 2025 70.156 4.799 3.611 78.565 11 2026 70.156 4.852 3.611 78.618 12 2027 70.156 4.799 3.611 78.565

13 2028 70.156 4.852 3.611 78.618

14 2029 70.156 4.799 3.611 78.565

15 2030 70.156 4.852 3.611 78.618

16 2031 70.156 -10.930 3.611 62.836 17 2032 70.156 -26.659 3.611 47.107 18 2033 70.156 -26.659 3.611 47.107

19 2034 70.156 -26.659 3.611 47.107

20 2035 70.156 -26.659 3.611 47.107

Subtotal 1403.115 0 72.214 1475.329

Total Investment 1459.612

C.5.2 Returns on equity and debt

The distribution of capital recovery among debt, preferred stock and common equity year-

by-year for the LNG plant is shown in Table C.16. This table is applied to calculate the

balance representing the basis for determining the returns on equity and the debt interest.

It is listed in Table C.14, columns 2-4, at the beginning of each year for each type of

financing. The total investment is distributed considering the financing fractions (see Table

C.1). Determination of the balances at the beginning of the first year (BBY1,x) for the xth type

of financing is expressed as

x1,x f*TNIBBY = , x = d, ps, ce (C.25)

The adjustment columns in the Table C.16 for different fractions of financing are analyzed

using the following equations. Additionally, the adjustment column for common equity also

includes the recovery of common equity allowance for funds used during construction.

A p p e n d i x C . E c o n o m i c A n a l y s i s D a t a a n d F l o w – A S t u d y C a s e o n t h e C 3 M R P r o c e s s

147

djdj, fDITXADJ = , j = 1, …, BL (C.26)

psjpsj, fDITXADJ = , j = 1, …, BL (C.27)

jcejcej, RCEAFfDITXADJ += , j = 1, …, BL (C.28)

The columns representing the book depreciation in Table C.16 are evaluated for each type of

financing by applying the straight-line depreciation

BL

ADJBBY

BD

BL

1kk,x,x

j,x

∑−= =

1

, j =1, …, BL, x = d, ps, ce

(C.29)

The balance at the beginning of the j th year for each type of financing is assessed as

( )j,xj,x1,xjj,x ADJBDBBYBBY +−= − , j = 2, …, BL, x = d, ps, ce (C.30)

The return on investment (ROI) for the year j is shown as

xj,xj,x iBBYROI = , j = 1, …, BL , x = d, ps, ce (C.31)

where ix accounts for the annual rate of return for the xth investment.

C.5.3 Taxes and insurance

Taxes and insurance are valued in the columns 5 and 6 in the Table C.14. The income taxes

(ITXj) which are to be paid during the j th year are analyzed through the equation (C.32)

( ) jjpsj,cej,j DITXRCEAFROIROIt

tITX −++

−=

1

(C.32)

where the term t is the total income tax rate, listed in Table C.1. The yearly sum of the other

taxes (property taxes) and insurance costs (OTXI) may be evaluated as a constant

percentage of the escalated plant facilities investment (PFI) as

OTXIj f*PFIOTXI = , j = 1, …, BL (C.33)

where the value of the factor fOTXI is assumed to be 2 %, listed in Table C.1 (entries 5b and

5c).

C.5.4 Fuel, operating and maintenance costs

The total fuel costs and O&M costs for the LNG plant for the first year of the operation are

considered in the subchapters C.3 and C.4 as 0.397*106 $ /year and 131.364*106 $/year,

respectively.

A p p e n d i x C . E c o n o m i c A n a l y s i s D a t a a n d F l o w – A S t u d y C a s e o n t h e C 3 M R P r o c e s s

148

C.5.5 Total revenue requirement (TRR)

Total revenue requirement for the LNG plant is evaluated for each year separately with the

aid of the equation (C.15) and is provided in the last column of the Table C.14.

A p p e n d i x C . E c o n o m i c A n a l y s i s D a t a a n d F l o w – A S t u d y C a s e o n t h e C 3 M R P r o c e s s

149

Table C.16. Distribution of capital recovery for the LNG plant (106 $).

Debt

Preferred Stock

Common Equity

Year

Calendar

Year

Balance Beginning

of Year

Book

Depreciation

Adjustment

Balance

Beginning of Year

Book

Depreciation

Adjustment

Balance

Beginning of Year

Book

Depreciation

Adjustment

Total Capital

Recovery

1 2016 729.806 36.490 0.000 218.942 10.947 0.000 510.864 22.718 3.611 73.766

2 2017 693.316 36.490 11.997 207.995 10.947 3.599 484.535 22.718 12.008 97.760

3 2018 644.829 36.490 9.464 193.449 10.947 2.839 449.808 22.718 10.236 92.694

4 2019 598.874 36.490 7.198 179.662 10.947 2.159 416.855 22.718 8.649 88.162

5 2020 555.186 36.490 5.145 166.556 10.947 1.544 385.487 22.718 7.212 84.057

6 2021 513.551 36.490 3.279 154.065 10.947 0.984 355.556 22.718 5.906 80.325

7 2022 473.781 36.490 2.399 142.134 10.947 0.720 326.932 22.718 5.290 78.565

8 2023 434.892 36.490 2.399 130.468 10.947 0.720 298.923 22.718 5.290 78.565

9 2024 396.002 36.490 2.426 118.801 10.947 0.728 270.915 22.718 5.309 78.618

10 2025 357.086 36.490 2.399 107.126 10.947 0.720 242.887 22.718 5.290 78.565

11 2026 318.196 36.490 2.426 95.459 10.947 0.728 214.879 22.718 5.309 78.618

12 2027 279.280 36.490 2.399 83.784 10.947 0.720 186.851 22.718 5.290 78.565

13 2028 240.390 36.490 2.426 72.117 10.947 0.728 158.843 22.718 5.309 78.618

14 2029 201.474 36.490 2.399 60.442 10.947 0.720 130.816 22.718 5.290 78.565

15 2030 162.584 36.490 2.426 48.775 10.947 0.728 102.807 22.718 5.309 78.618

16 2031 123.668 36.490 -5.465 37.100 10.947 -1.640 74.780 22.718 -0.215 62.836

17 2032 92.643 36.490 -13.330 27.793 10.947 -3.999 52.276 22.718 -5.720 47.107

18 2033 69.482 36.490 -13.330 20.845 10.947 -3.999 35.278 22.718 -5.720 47.107

19 2034 46.321 36.490 -13.330 13.896 10.947 -3.999 18.279 22.718 -5.720 47.107

20 2035 23.161 36.490 -13.330 6.948 10.947 -3.999 1.281 22.718 -5.720 47.107

21 2035 0 36.490 0 0.000 10.947 0 -15.717 22.718 0 70.156

Totals 729.806 0 218.942 0 454.367 72.214 1475.329

A p p e n d i x C . E c o n o m i c A n a l y s i s D a t a a n d F l o w – A S t u d y C a s e o n t h e C 3 M R P r o c e s s

150

C.5.6 Levelized Costs and the Cost of the Main Product

The evaluation of the annual values of carrying charges (CC), fuel costs (FC) and O&M

expenses of a thermal system is necessary for the evaluation of the design modifications and

cost effectiveness. These cost components are expected to vary considerably over the entire

economic life of the production facilities. Carrying costs show a decreasing trend where the

fuel and O&M charges tend to increase with increasing years of operation. Therefore, it is

recommended to use levelized annual values for cost components in terms of design

modifications. The levelized value of the annual total revenue requirement is assessed by

taking the summation of the present values of the 20 annual values in Table C.14. These

values are then converted to equivalent annuities with the help of the equation (C.16). The

average cost of money (ieff) of 12 % is assumed as the discount rate during the evaluations.

The levelized fuel costs (FCL) are considered with the aid of the equations (C.17) and

(C.18) as

946.012.01

06.01=

++

=Fk

134.01)12.01(

)12.01(12.020

20

=−+

+=CRF

6206

10*145.190)946.01(

)134.0)(946.01(

06.1

10*364.131=

−−

×=LFC $.

Equally, the levelized annual operating and maintenance costs (OMCL) are projected using

the equations (C.17) and (C.19):

9375.012.01

05.01=

++

=OMk

6206

10*538.0)9375.01(

)134.0)(9375.01(9375.0

05.1

10*398.0=

−−

×=LOMC .

The levelized annual total revenue requirement is calculated as the sum of the

present values of 20 annual values given in year by year analysis in Table C.14. It is the

converted to an equivalent annuity with the aid of the equation (C.17).

∑==

−−

20

1nyearbyyearP TRRTRR

$10*595.46110*861.3447*134.0 66 === PL TRR*CRFTRR .

A p p e n d i x C . E c o n o m i c A n a l y s i s D a t a a n d F l o w – A S t u d y C a s e o n t h e C 3 M R P r o c e s s

151

The levelized annual carrying charges (CCL) can be estimated then as

$10*912.270 6=−−= LLLL OMCFCTRRCC

The ratio between the annual total revenue requirement (TRR) and the product

quantity (MPQ) calculates the unit cost of the product (MPUC):

MPQ

TRRMPUC L=

(C.34)

The primary product of the liquefaction plant is the liquefied natural gas with a mass

flow rate of 158.4 kg/s. The yearly quantity of the product is calculated seeing the plant

capacity of 85% as follow

610*246.41744636004.158 ==kg

Ton

yr

h

h

s

s

kgMPQ Ton/yr.

The levelized unit cost of the liquefied natural gas is then considered using the equation

(C.34):

kgTon /$11.0/$71.108Ton10*246.4

$10*595.461

MPQ

TRRMPUC

6

6L ====

A p p e n d i x C . E c o n o m i c A n a l y s i s D a t a a n d F l o w – A S t u d y C a s e o n t h e C 3 M R P r o c e s s

152

Table C.17. LNG plant data set [98].

LNG Plant

Country Startup

Year Technology

Capacity (mtpa)

Capital Unit Costs

($ bn/mtpa)

Marsa El Brega Libya 1970 APCI 2.6 0.163

Lumut Brunei 1972 APCI 5.3 0.253

Arun, Phase 1 Indonesia 1978 APCI 4.2 0.532

Adgas, Das Island Abu Dhabi 1977 APCI 3.0 0.414

Bonny Island, T1+2 Nigeria 1999 APCI 5.9 0.421

Bonny Island T3 Nigeria 2002 APCI 2.9 0.289

MLNG I, Bintulu Malaysia 1984 APCI 6.0 0.495

MLNG II (Dua) Malaysia 1995 APCI 5.2 0.403

MLNG III (TIGA) Malaysia 2003 APCI 7.6 0.211

Damietta Egypt 2005 APCI 5.0 0.208

Bonny Island, T4+5 Nigeria 2005 APCI 8.0 0.217

Qalhat, T3 Oman 2006 APCI 3.30 0.189

A p p e n d i x C . E c o n o m i c A n a l y s i s D a t a a n d F l o w – A S t u d y C a s e o n t h e C 3 M R P r o c e s s

153

Table C.18. Economic data for the selected LNG plants [98].

LNG Plant Country Initial

Start-up Capacity/ train

(MTA) Capital Unit Costs (2013 $ bn/MTA)

One train

Adgas, Das Island Abu Dhabi 1994 2.60 0.635

Bontang Indonesia 1998 3.00 0.605

Damietta Egypt 2005 5.00 0.293

Qalhat Oman 2006 3.30 0.267

Two trains

Rasgas Qatar 1999 2.60 0.443

Qalhat Oman 1999 3.00 0.388

MLNG Malaysia 2003 3.80 0.298

Bonny Island Nigeria 2005 4.00 0.306

Tangguh Indonesia 2007 3.50 0.282

Three trains

Burrup Northwest

Australia 1989 2.00 0.638

MLNG Malaysia 1995 2.60 0.569

Qatargas Qatar 1998 2.50 0.427

Bonny Island Nigeria 1999 3.20 0.594

Bonny Island Nigeria 2002 3.70 0.408

Seven trains

Bontang Indonesia 1999 2.60 0.368


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