Comparati v e e v aluation of methanol
pr oduction pr ocesses using natural gas: A
thermodynamic and economic assessment
v or gele gt v on
M.Sc.
T imo Dominik Blumber g
geb . in Berlin
v on der F akultät III - Prozesswissenschaften
der T echnischen Uni versität Berlin
zur Erlangung des akademischen Grades
Doktor der Ingenieurwissenschaften
-Dr .-Ing.-
genehmigte Dissertation
Promotionsausschuss
V orsitzender: Prof. Dr . F . Behrendt
Gutachter: Prof. Dr .-Ing. G. Tsatsaronis
Gutachterin: Prof. Dr . T . Morozyuk
Gutachter: Dr .-Ing. Y .D. Lee
T ag der wissenschaftlichen Aussprache: 27.07.2018
Berlin 2018

Ac kno wledg ement
This work w as generated during the time I ha v e work ed as a research associate at the
Department of Ener gy Engineering of the Zentralinstitut EL Gouna at the Berlin Institute
of T echnology (T echnische Uni v ersität Berlin). I enjoyed the teaching and the supervision
of students, the administrati v e w ork in my position as the coordinator of the Department,
and the pleasant atmosphere with my colleagues. The possibility of working in the tw o
countries German y and Egypt contrib uted greatly to my indi vidual de velopment and my
social an cultural understanding. In this conte xt, my special thanks go to Eng. Samih
Sa wiris, without whom I would not ha ve had this opportunity . He and his compan y
Orascom b uilt up a unique Campus, which enables researchers to conduct their work in a
perfect scientific en vironment.
I am v ery grateful to Prof. Geor ge Tsatsaronis and Prof.T etyana Morozyuk for the
supervision of this thesis. Both professors alw ays supported me in scientific and personal
terms. I also would lik e to e xpress my gratitude to Dr . Y oung Duk Lee for ha ving
re vie wed this thesis. Furthermore, I would like to thank Prof. Frank Behrendt for his
willingness to chair my defense.
Moreo ver I w ould lik e to thank my colleagues Stefan Bruche, Johannes W ellmann, Sarah
Hamdy , Stefanie T esch, Christoph Banhardt, Louay Elmorsy , Alexander Studniorz, Max
Sor genfrei, Saaed Sayadi, Bahaa Noaman, Jör g Rüdiger and Sarah Al Ahmad for the
fruitful discussions and their support. I really enjoyed the time we spent together in the
institutes, in the Egyptian campus and especially on the conferences.
Ev entually , I would like thank my w onderful family and my girlfriend who alw ays
supported and moti v ated me. Thanks to my mother Christine, my sister Nadine, my
brother -in-la w Robert, and their wonderful children Sophia and Isabel. I would like to
dedicate this work to my deceased f ather Armin.
Berlin, 2018
T imo Blumberg

Zusammenfassung
In der v orlie genden Arbeit werden v erschiedene erdgasbasierte Prozesse für die Erzeu-
gung v on Methanol aus thermodynamischer und ökonomischer Sicht analysiert. K om-
merziell v erfügbare T echnologien sowie vielv ersprechende Zukunftsalternati ven für
erhöhte Ef fizienz, eine v erbesserte W irtschaftlichkeit so wie reduzierte T reibhausg asemis-
sionen werden be wertet und miteinander v er glichen. Die untersuchten Prozesse basieren
auf der indirekten Synthese mit dem Zwischenschritt der Synthese gaserzeugung. Die
einzelnen Prozesse sind durch eine unterschiedliche Reformierungstechnologie so wie
einen isothermen Synthesereaktor gekennzeichnet. Die Auswahl geeigneter Reformer -
und Reaktortechnologien so wie deren Betrieb ist entscheidend für das thermodynamische
V erhalten einer Methanolanlage. P arameterstudien werden für die Basistechnologien
durchgeführt um v orteilhafte Betriebsparameter für hohe Umsätze und Ausbeuten zu bes-
timmen. Die kommerziell gängige Nutzung einf acher Reformertechnologien erzielt nur
geringe Produktausbeuten. Eine K ombination verschiedener Reformer - und Gaskondi-
tionierungstechnologien ermöglicht die Bereitstellung eines hochreakti v en Synthese gases
für höhere Produktausbeuten. Eine Untersuchung der beiden k ommerziell erfolgreich-
sten Reaktortechnologien weist wesentlich höhere Methanolausbeuten für die isotherme
Reaktortechnologie als für die adiabate Mehrbett-Quench-T echnologie aus. Eine En-
er gieanalyse zeigt, dass die Prozesse mit endothermer Reformertechnolgie zwangsläufig
eine K oproduktion von Methanol und Elektrizität für hohe Ef fizienz ermöglichen müssen.
Die
CO 2
Nutzung durch trockene Reformierung und Hydrierung im Synthesereaktor wird
erfolgreich für eine V erringerung der Emissionen und des Brennstof fbedarfs eingesetzt.
Im Rahmen einer Ex ergieanalyse werden die größten irre versibilitäten für die Prozesse
mit endothermer Reformierung ermittelt. Diese können der V erbrennung zur Prozess-
wärmebereitstellung zugeordnet werden. Mit Hilfe der Ertragsbedarfsmethode werden
K osten vorteile für die Prozesse mit autothermer T echnologie aufgedeckt. Die Ergebnisse
der e xer goök onomischen Analyse zeigen, dass der Zwei-Stufen-Reformierungsprozess
die Produkte zu den niedrigsten K osten bereitstellen kann, während das Methanol im
kon ventionellen Dampfreformierungsprozess am teuersten produziert wird. Auf der
K omponentenebene haben die V erbrennungseinheiten aufgrund v on Irre v ersibilitäten
und der Luftkompressor als kapitalintensi vste K omponente den größten K osteinfluss. Im
F alle einer Monetarisierung des integrierten
CO 2
steigen die Produktkosten an, w obei
der T rockenreformierungsprozess v on der größten K ostensteigerung betrof fen ist. Eine
erweiterte Ex er gieanalyse für den Zwei-Stufen-Reformierungsprozess zeigt, dass der
Großteil der Ex er gie v ernichtung endogen und für einige wichtige K omponenten vermei-
dbar ist.

Abstract
In the work, dif ferent processes for the production of methanol from natural gas are
analysed from a thermodynamic and an economic point of vie w . Commercially av ailable
processes and promising future alternati v es for increased thermodynamic ef ficiency ,
impro ved economics and reduced greenhouse gas emissions are assessed and compared.
The analysed processes are based on the indirect synthesis route via syngas production.
Each design features a dif ferent reforming technology in conjunction with an isothermal
synthesis reactor . The choice of suitable reformer and reactor technologies and their
operation are decisi v e for the thermodynamic performance of a methanol plant. Param-
eter studies are conducted for the basic technologies to determine fa v orable operation
conditions for high reactant con version and product yield. The commercial use of simple
reforming technologies only enables for a lo w methanol yield. A combination of dif ferent
reforming technologies with application of syngas conditioning units generates a highly
reacti v e syngas for impro v ed methanol yield. The two main reactor technologies in
commercial use are in vestig ated - the isothermal reactor technology allo ws much higher
con version rates than the adiabatic multi-bed quench reactor technology . An analysis of
the ener gy distrib ution sho ws that the processes with endothermic reforming ine vitably
feature a co-production of electricity for a high ef ficienc y .
CO 2
utilization by dry reform-
ing and direct hydrogenation in the synthesis unit is successfully used for an reduction
of the emissions and the fuel demand. A con ventional e x er getic analysis identifies the
highest inef ficiencies for processes with endothermic reforming technology , these being
primarily related to the comb ustion unit for process heat supply . The application of
the total re v enue requirement method identifies economic adv antages of the processes
with autothermal reforming caused by a lo w fuel consumption. Furthermore, the results
of an e xer goeconomic analysis sho w , that the two-step reforming process can generate
the products at the lo west cost, while methanol obtained from a con v entional steam
methane reforming process is associated with the highest cost. On a component le v el,
the comb ustion units and the air compressor ha v e the highest cost significance due to
the cost rate associated with irre v ersibilities and the in vestment, respecti v ely . In case
of a monetization of the inte grated
CO 2
, the product cost increase. In particular in the
dry reforming process, the product costs are af fected by the cost of
CO 2
. An adv anced
e xer getic analysis of the two-step reforming process re v eals that the major part of the
inef ficiencies is endogenous and a v oidable for some important components.
v

Contents
A c k n o w l e d g e m e n t ............................... i
Z u s a m m e n f a s s u n g ................................ i i i
A b s t r a c t ..................................... v
N o m e n c l a t u r e .................................. x i v
1 Intr oduction 1
2 State of Resear ch 9
2.1 Syngas Generation from Natural Gas . . . . . . . . . . . . . . . . . . . 10
2 . 1 . 1 P r e t r e a t m e n t ........................... 1 2
2 . 1 . 2 P r e - r e f o r m i n g ........................... 1 2
2.1.3 Hydrocarbon Reforming . . . . . . . . . . . . . . . . . . . . . 13
2 . 1 . 4 S t e a m R e f o r m i n g ......................... 1 3
2.1.5 Heat Exchange Reforming / Gas Heated Reforming . . . . . . . 15
2.1.6 P artial Oxidation . . . . . . . . . . . . . . . . . . . . . . . . . 16
2.1.7 Autothermal Reforming . . . . . . . . . . . . . . . . . . . . . 17
2.1.8 T wo-step Reforming . . . . . . . . . . . . . . . . . . . . . . . 19
2.1.9 Dry and Mixed Reforming . . . . . . . . . . . . . . . . . . . . 22
2.1.10 Process Selection Criteria for Methanol Generation . . . . . . . 23
2.2 Conditioning of Synthesis Gas . . . . . . . . . . . . . . . . . . . . . . 26
2.2.1 Carbon Utilization Measures in the Production of Methanol . . 26
2.2.2 (Re v erse) W ater-g as Shift Reaction . . . . . . . . . . . . . . . 28
2.3 Methanol Synthesis Process . . . . . . . . . . . . . . . . . . . . . . . 29
2.3.1 Chemistry and Thermodynamics . . . . . . . . . . . . . . . . . 29
2.3.2 Methanol Reactor T echnology and Synthesis Configuration . . . 30
2.3.2.1 Adiabatic Reactor / Multi-Bed Con verter . . . . . . . 33
2.3.2.2 Isothermal Reactor / Single-Bed Con verter . . . . . . 34
2.4 Crude Methanol Purification . . . . . . . . . . . . . . . . . . . . . . . 36
3 Methodology 38
3.1 Thermodynamic Analysis . . . . . . . . . . . . . . . . . . . . . . . . . 38
3.1.1 Ener getic Analysis . . . . . . . . . . . . . . . . . . . . . . . . 38
3.1.2 Con ventional Exer getic Analysis . . . . . . . . . . . . . . . . . 40
3 . 2 E c o n o m i c A n a l y s i s ............................ 4 2
3.3 Ex er goeconomic Analysis . . . . . . . . . . . . . . . . . . . . . . . . 45
3.4 Adv anced Exer getic Analysis . . . . . . . . . . . . . . . . . . . . . . . 47
vi

3.5 Simulation and Software . . . . . . . . . . . . . . . . . . . . . . . . . 50
4 System Design and Modelling 52
4.1 P arameter Study for the Synthesis of Process Designs . . . . . . . . . . 52
4.1.1 Reforming of Methane . . . . . . . . . . . . . . . . . . . . . . 52
4.1.1.1 Steam Reforming . . . . . . . . . . . . . . . . . . . 53
4.1.1.2 Autothermal Reforming . . . . . . . . . . . . . . . . 54
4.1.1.3 Dry and Combined Reforming . . . . . . . . . . . . 55
4.1.1.4 Comparati v e Assessment of the Syngas Characteristics 57
4.1.2 Methanol Synthesis . . . . . . . . . . . . . . . . . . . . . . . . 59
4.2 Ov ervie w of Processes and Subsystems . . . . . . . . . . . . . . . . . 62
4 . 3 B a s i c A s s u m p t i o n s ............................ 6 3
4 . 4 S M R P r o c e s s ............................... 6 5
4 . 5 A T R P r o c e s s ................................ 6 9
4 . 6 D M R P r o c e s s ............................... 7 3
4 . 7 C M R P r o c e s s ............................... 7 7
4 . 8 S M R - A T R P r o c e s s ............................ 8 1
4 . 9 S M R - D M R P r o c e s s ............................ 8 6
5 Results and Discussion 91
5.1 Process Performance Analysis . . . . . . . . . . . . . . . . . . . . . . 91
5.2 Potential Analysis for Carbon Dioxide Utilization . . . . . . . . . . . . 97
5.3 Con ventional Exer getic Analysis . . . . . . . . . . . . . . . . . . . . . 98
5 . 4 E c o n o m i c A n a l y s i s ............................ 1 0 4
5.5 Ex er goeconomic Analysis . . . . . . . . . . . . . . . . . . . . . . . . 116
5.6 Adv anced Exer getic Analysis . . . . . . . . . . . . . . . . . . . . . . . 140
6 Conclusions and Outlook 146
6.1 Process Performance Analysis . . . . . . . . . . . . . . . . . . . . . . 147
6.2 Con ventional Exer getic Analysis . . . . . . . . . . . . . . . . . . . . . 148
6 . 3 E c o n o m i c A n a l y s i s ............................ 1 4 9
6.4 Ex er goeconomic Analysis . . . . . . . . . . . . . . . . . . . . . . . . 150
6.5 Adv anced Exer getic Analysis . . . . . . . . . . . . . . . . . . . . . . . 151
6 . 6 O u t l o o k .................................. 1 5 1
Bibliograph y 152
A ppendix 165
A Specification of Industrial Methanol and Composition of Natural Gas 165
B Reaction Kinetics 167
C Sensitivity Analyses of the Methanol Reactor 172
D Detailed Str eam Data f or the Pr ocesses 175
vii

E Exergetic Analysis 190
F Adv anced Exergetic Analysis 199
viii

List of Figures
1.1 Methanol demand forecast. . . . . . . . . . . . . . . . . . . . . . . . . . . 3
1.2 Methanol product distribution in 2016. . . . . . . . . . . . . . . . . . . . . 4
1.3 Ov ervie w of the methanol v alue chain. . . . . . . . . . . . . . . . . . . . . 7
2.1
Schematic of the synthesis routes for the production of methanol from natural
g a s . ...................................... 1 1
2.2 T ypes of heat exchange reformers. . . . . . . . . . . . . . . . . . . . . . . 15
2.3 Design of an autothermal reformer . . . . . . . . . . . . . . . . . . . . . . . 18
2.4 T wo-step reforming concepts. . . . . . . . . . . . . . . . . . . . . . . . . . 21
2.5 Capacity range of the commercial reforming technologies. . . . . . . . . . 25
2.6 Ov ervie w of commercial methanol reactor technologies. . . . . . . . . . . 31
2.7 Scheme of the isothermal ractor and the multi-bed quench reactor . . . . . . 34
2.8 Process operation lines of an adiabatic reactor and an isothermal reactor . . . 35
2.9 Flo wsheet of a three-column distillation system. . . . . . . . . . . . . . . . 37
3.1
Ov ervie w of splitting options of the e x er gy destruction within a component
in an adv anced ex er gy analsyis. . . . . . . . . . . . . . . . . . . . . . . . . 48
4.1 Sensiti vity analysis of the steam methane reforming technology . . . . . . . 53
4.2 Sensiti vity analysis of the autothermal reforming technology . . . . . . . . . 55
4.3 Sensiti vity analysis of the dry methane reforming technology . . . . . . . . . 56
4.4
Range of the stoichiometric measures for the common syngas production
t e c h n o l o g i e s . ................................. 5 7
4.5 Sensiti vity of the methanol yield to the inlet composition at equilibrium. . . 58
4.6
Sensiti vity of the methanol yield to the inlet composition and the operation
conditions for dif ferent kinetic models. . . . . . . . . . . . . . . . . . . . . 60
4.7
Methanol yield from the isothermal reactor (left) and the adibatic quench
reactor (right) for dif ferent processes. . . . . . . . . . . . . . . . . . . . . . 61
4.8 Process flo wsheet of the methanol plant with steam reforming. . . . . . . . 68
4.9 Process flo wsheet of the methanol plant with autothermal reforming. . . . . 72
4.10 Process flo wsheet of the methanol plant with dry reforming. . . . . . . . . 76
4.11 Process flo wsheet of the methanol plant with combined reforming. . . . . . 80
4.12 T emperature profiles of heat transfer within the CMR process. . . . . . . . 81
4.13 T emperature profiles of heat transfer within the SMR-A TR process. . . . . . 83
4.14 Process flo wsheet of the methanol plant with two-step reforming. . . . . . . 84
4.15 T emperature profiles of heat transfer within the SMR+DMR process. . . . . 88
ix

4.16 Process flo wsheet of the methanol plant with steam and dry reforming. . . . 89
5.1 Ener gy distribution within the analysed processes. . . . . . . . . . . . . . . 93
5.2
Results obtained from the con ventional e x er getic analysis for the aggre gated
s u b s y s t e m s . .................................. 1 0 1
5.3
Impact of the natural gas cost and the sales price of electicity on the sales
price of methanol for the SMR process. . . . . . . . . . . . . . . . . . . . 107
5.4
Impact of the natural gas price and the sales price of electicity on the sales
price of methanol for the A TR process. . . . . . . . . . . . . . . . . . . . . 109
5.5
Impact of the natural gas cost and the sales price of electicity on the sales
price of methanol for the DMR process. . . . . . . . . . . . . . . . . . . . 109
5.6
Impact of the natural gas cost and the sales price of electicity on the sales
price of methanol for the CMR process. . . . . . . . . . . . . . . . . . . . 110
5.7
Impact of the natural gas cost and the sales price of electicity on the sales
price of methanol for the SMR-A TR process. . . . . . . . . . . . . . . . . 111
5.8
Impact of the natural gas cost and the sales price of electicity on the sales
price of methanol for the SMR-DMR process. . . . . . . . . . . . . . . . . 112
5.9
Sensiti vity of the contrib ution mar gin depending on the minimum methanol
price and the selling price of elecrtcity for fuel cost of 3 US$/GJ. . . . . . . 113
5.10
Sensiti vity of the contrib ution mar gin depending on the minimum methanol
price and the selling price of elecrtcity for fuel cost of 7 US$/GJ. . . . . . . 114
5.11
Sensiti vity of the contrib ution mar gin depending on the minimum methanol
price and the selling price of elecrtcity for fuel cost of 11 US$/GJ. . . . . . 115
5.12
Relati v e cost rate associated with the in v estment and the inef ficiencies of the
o v e r a l l p r o c e s s e s . ............................... 1 1 6
5.13 Le velized cost of methanol as a function of the fuel cost. . . . . . . . . . . 136
5.14 Le velized cost of electricity as a function of the fuel cost. . . . . . . . . . . 137
5.15
Le v elized cost of methanol as a function of the fuel cost under uncertainty
of the cost for CO 2 . .............................. 1 3 8
5.16
Le v elized cost of electricty as a function of the fuel cost under uncertainty
of the cost for CO 2 . .............................. 1 3 9
5.17
Results of splitting the e xer gy destruction into its unav oidable and a v oidable
endogenous and exogenous parts. . . . . . . . . . . . . . . . . . . . . . . . 145
C.1 Contour plot of the component con versions within the synthesis. . . . . . . 173
D.1 Process flo wsheet of the methanol plant with steam reforming. . . . . . . . 176
D.2 Process flo wsheet of the methanol plant with autothermal reforming. . . . . 178
D.3 Process flo wsheet of the methanol plant with dry reforming. . . . . . . . . 180
D.4 Process flo wsheet of the methanol plant with combined reforming. . . . . . 182
D.5 Process flo wsheet of the methanol plant with tw o-step reforming. . . . . . . 184
D.6 Process flo wsheet of the methanol plant with steam and dry reforming. . . . 186
E.1
Ex er gy destruction rate within the aggre gated subsystems of the SMR process.
193
E.2
Ex er gy destruction rate within the aggreg ated subsystems of the A TR process.
194
E.3
Ex er gy destruction rate
˙
E D
for the aggre gated subsystems of the DMR process.
195
x

E.4
Ex er gy destruction rate
˙
E D
for the aggre gated subsystems of the CMR process.
196
E.5
Ex er gy destruction rate within the aggre gated subsystems of the SMR-A TR
p r o c e s s . .................................... 1 9 7
E.6
Ex er gy destruction rate within the aggreg ated subsystems of the SMR-DMR
p r o c e s s . .................................... 1 9 8
xi

List of T ab les
2.1
Features of the isothermal reactor technology and the adiabatic multi-bed
quench reactor technology . . . . . . . . . . . . . . . . . . . . . . . . . . . 32
3.1 Specifications of the economic analysis. . . . . . . . . . . . . . . . . . . . 45
4.1 Specifications of the analysed processes. . . . . . . . . . . . . . . . . . . . 62
4.2 Basic assumptions for all processes. . . . . . . . . . . . . . . . . . . . . . 64
4.3 Design specifications of the SMR process. . . . . . . . . . . . . . . . . . . 66
4.4 Simulation results for selected flo ws of the SMR process. . . . . . . . . . . 67
4.5 Design specifications of the A TR process. . . . . . . . . . . . . . . . . . . 70
4.6 Simulation results for selected flo ws of the A TR process. . . . . . . . . . . 71
4.7 Design specifications of the DMR process. . . . . . . . . . . . . . . . . . . 74
4.8 Simulation results for selected flo ws of the DMR process. . . . . . . . . . . 75
4.9 Design specifications of the CMR process. . . . . . . . . . . . . . . . . . . 78
4.10 Simulation results for selected flo ws of the CMR process. . . . . . . . . . . 79
4.11 Design specifications of the SMR-A TR process. . . . . . . . . . . . . . . . 82
4.12 Simulation results for selected flo ws of the SMR-A TR process. . . . . . . . 85
4.13 Design specifications of the SMR-DMR process. . . . . . . . . . . . . . . 87
4.14 Simulation results for selected flo ws of the SMR-DMR process. . . . . . . 90
5.1 Selected results of the o v erall process. . . . . . . . . . . . . . . . . . . . . 92
5.2 Selected results for the reforming unit and the synthesis reactor . . . . . . . . 95
5.3 K e y figures of the CO 2 -abatement potential within the analysed processes. . 97
5.4
Results obtained from the con ventional e x er getic analysis for the ov erall
s y s t e m . .................................... 9 9
5.5
Results obtained from the con ventional e x er getic analysis for the components
with the highest ex er gy destruction. . . . . . . . . . . . . . . . . . . . . . 103
5.6 Results obtained from the economic analysis for the o verall systems. . . . . 105
5.7
Results obtained from the ex er goeconomic analysis for a selection of com-
ponents of the reference SMR process. . . . . . . . . . . . . . . . . . . . . 120
5.8
Results obtained from the ex er goeconomic analysis for a selection of com-
ponents of the reference A TR process. . . . . . . . . . . . . . . . . . . . . 123
5.9
Results obtained from the ex er goeconomic analysis for a selection of com-
ponents of the reference DMR process. . . . . . . . . . . . . . . . . . . . . 126
5.10
Results obtained from the ex er goeconomic analysis for a selection of com-
ponents of the reference CMR process. . . . . . . . . . . . . . . . . . . . . 129
xii

5.11
Results obtained from the ex er goeconomic analysis for a selection of com-
ponents of the reference SMR-A TR process. . . . . . . . . . . . . . . . . . 132
5.12
Results obtained from the ex er goeconomic analysis for a selection of com-
ponents of the reference SMR-DMR process. . . . . . . . . . . . . . . . . 134
5.13 Assumptions for the determination of the unav oidable e x er gy destruction. . 141
A.1 U.S. Federal grade specification for Methanol. . . . . . . . . . . . . . . . . 165
A.2 Composition of natural gas assumed for the simulations. . . . . . . . . . . 166
D.1 Stream data obtained from the simulation of the SMR process. . . . . . . . 177
D.2 Stream data obtained from the simulation of the A TR process. . . . . . . . 179
D.3 Stream data obtained from the simulation of the DMR process. . . . . . . . 181
D.4 Stream data obtained from the simulation of the CMR process. . . . . . . . 183
D.5 Stream data obtained from the simulation of the SMR-A TR process. . . . . 185
D.6 Stream data obtained from the simulation of the SMR-DMR process. . . . . 188
E.1 Definitions of the e x er getic ef ficienc y for selected components. . . . . . . . 191
E.1
Results obtained from the con ventional e x er getic analysis for the aggre gated
subsystems of the SMR process. . . . . . . . . . . . . . . . . . . . . . . . 193
E.2
Results obtained from the con ventional e x er getic analysis for the aggre gated
subsystems of the A TR process. . . . . . . . . . . . . . . . . . . . . . . . . 194
E.3
Ex er gy destruction rate within the aggreg ated subsystems of the DMR process.
195
E.4
Ex er gy destruction rate within the aggre gated subsystems of the CMR process.
196
E.5
Results obtained from the con ventional e x er gy analysis for the aggre gated
subsystems of SMR-A TR process. . . . . . . . . . . . . . . . . . . . . . . 197
E.6
Results obtained from the con ventional e x er gy analysis for the aggre gated
subsystems of the SMR-DMR process. . . . . . . . . . . . . . . . . . . . . 198
F .1
Results obtained from the con ventional and the adv anced ex er getic analyses
for selected component groups. . . . . . . . . . . . . . . . . . . . . . . . . 201
xiii

Nomenc lature
E A acti v ation energy J
f e xer goeconomic f actor %
h specific enthalpy J/kg
˙
H enthalpy flo w rate W
i interest rate %
k heat transfer coef ficient W/m 2 K
ke kinetic energy J/kg
K E total kinetic ener gy J
˙ m mass flo w rate kg/s
˙ n mole flo w rate mol
p pressure bar
pe specific potential ener gy J/kg
PE potential ener gy J
˙
Q rate of heat transfer J
r relati v e cost dif ference %
R reaction rate
R ideal gas constant J/molK
T temperature ◦ C, K
U internal ener gy J
U
o verall heat transfer coef fi-
cient
W/(m 2 K)
V v olume m 3
˙
W po wer W
x mole fraction mol/mol
x steam quality kg steam /kg tot
y D , y ∗
D e xer gy destruction ratio %
Gr eek symbols
α scaling e xponent −
∆ dif ference −
ε e xer getic ef ficiency %
xi v

η s isentropic ef ficienc y %
η pol polytropic ef ficienc y %
τ
time/ residence time / annual
full load hours
s/h
ρ density kg/m 3
Superscripts
˙ time deri v ativ e
˜ mole related
ad s adsorption
A V a v oidable
CH chemical
EN endogenous
EX e xogenous
PH physical
UN una v oidable
0 refers to ambient conditions
C compressor
c h chemical
cv control v olume
D destruction
el electrical
F fuel
gen general
i , j running indicies
in inlet
k k -th plant component
l length of tube
L le v elized
Subscripts
L loss
n nominal, number of years
out outlet
xv

pol polytropic
PE purchased equipment
s isentropic
sys subsystem
t ot o v erall system
T turbine/ e xpander
X capacity of the amount X
Y capacity of the amount Y
Abbr e viations
AF U D C allo w ance for funds during construction time
A GR acid gas remo v al
ASU air separation unit
A TR autothermal reforming
C distillation column in flo wsheet
C C carrying char ges
CCS carbon capture and storage
CCU carbon capture and utilization
C E L F constant escalation le velization f actor
CEPCI chemical engineering plant cost inde x
CM compressor in flo wsheet
CMR combined methane reforming
CtL coal to liquid
D drum in flo wsheet
DISTL distillation subsystem
DMR dry methane reforming
DMT Dimethylterephtalat
ECO economizer
EV A e v aporator
F C fuel cost
F C I fixed capital in vestment
GtL gas to liquid
HEX heat exchanger in flo wsheet
HHV higher heating v alue
HP high pressure
xvi

Abbr e viations
HRSG heat reco very steam generator
ICI Imperial Chemical Industries
IP intermediate pressure
LHV lower heating v alue
LCOE le v elized cost of electricity
LCOM le v elized cost of methanol
LP lo w pressure
LPM liquid phase methanation
MEA Monoethanolamine
MC minimum cost
MTBE methyl-tert-b utylether
MT O Methanol-to-olefins process
MTP Methanol-to-propylene process
mtpd metric tons per day
NIST National Institute of Standard and T echnology
O/C oxygen to carbon ratio
OMC operation and maintenance cost
PC-SAFT perturbed chain statistical association fluid theory
PEC purchased equipment cost
R reactor in flo wsheet
RE subsystem reforming unit
RKS-BM
Redlich-Kwong-Soa v e with Boston-Matthias alpha
function
R TFR rec ycle to feed ratio
S/C steam to carbon ratio
SH superheater
SPECO specific e xer gy costing
SMR steam methane reforming
SYN subsystem synthesis unit
T turbine in flo wsheet
T C I total capital in vestment
WGS water -g as shift
xvii

Chapter 1
Intr oduction
The rising standard of li ving of an e vermore gro wing world population requires a steady
economic gro wth, which can only be ensured by co v ering the demand for ener gy and
other essential products. The use of fossil feedstock currently is the only mean to cope
with rapid increasing consumption. Coal, oil and natural gas are still considered as a
reliable and relati v ely ine xpensi v e primary resource of ener gy for most industrialized
countries. In 2017 fossil fuels accounted for 82.7% of the global energy supply [1, 2].
Besides being the major ener gy source, fossil substances also serv e as a feedstock for
a great v ariety of deri ved hydrocarbon materials and products. The range comprises
petrochemical and chemical products, including plastics, pharmaceuticals and synthetic
fuels, as well as diesel oil or gasoline [3]. The chemicals and petrochemicals sector is
one of the fastest gro wing industries and by far the lar gest industrial ener gy consumer
accounting for 10% of the global final ener gy consumption in 2014 [4]. The strong
demand gro wth for hydrocarbon deri ved products, in particular methanol, increases the
pressure on the diminishing fossil resources.
A problem closely linked to the intensi ve use of fossil fuels refers to the emissions of
greenhouse gases. Utilization of coal, oil and natural gas in the ener gy and industrial
sector significantly contrib utes to anthropogenic greenhouse gas emissions. F or e xample,
the chemicals and petrochemicals sector accounted for 13% of the direct industrial
CO 2
emissions in 2014 [4]. Carbon capture and storage (CCS) and carbon capture
and utilization (CCU) represent the two major strate gies to reduce these emissions [5].
While CCS technologies are de v eloped with a focus on capturing and storing
CO 2
in
lar ge quantities, no vel concepts for the chemical v alorization of
CO 2
are de v eloped
simultaneously . Although the former technology is currently considered as the most
promising a v oidance mechanism, the CCU appears to be more beneficial, due to the
abatement of
CO 2
emissions in conjunction with an ef fecti v e production of v aluable
fuels and chemicals [6].
1

Chapter 1 Introduction
In the last decade, ne w carbon utilization concepts were proposed for the inte gration
of
CO 2
as an auxiliary feedstock or commodity into Gas-to-Liquids (GtL) processes [7,
8]. The lar ge-scale production of methanol from its predominating feedstock natural gas
(NG) attracted particular attention, due to the increasing product demand and ab undant
gas resources [9–12].
A promising approach to reduce the dependence from fossil fuels is based on a
successi v e substitution by methanol and its v ariety of h ydrocarbonecous deri v ates. In
2005, this approach has been suggested by Geor ge Olah, USA, who entitled the concept
as the Methanol Economy [13–15]. Thus, methanol as an inte gral part of life not only
would constitute a partial replacement of fossil fuels, b ut also could act as a
CO 2
sink
when chemical rec ycling is integrated into future and retrofitting manuf acture concepts.
Methanol is rated among the most important feedstock for the chemical, petrochemical
and ener gy industries, with a worldwide demand of 83 million metric tons in 2016 [16,
17]. At present, the b ulk chemical is almost e xclusi v ely produced from synthesis gas
deri v ed by catalytic reforming and gasification of natural g as and coal. Commercial
synthesis routes are associated with technological complexity , energy intensity and
costly operation. Ho we ver , only fossil-based processes enable a lar ge-scale production
which can hold pace with the rising demand. The deri vation from biomass, biogas,
municipal waste and agricultural by-products is of minor significance [3]. Alternati v e
direct con version processes without the initial and cost intensi ve step of syngas production
may economically be more viable b ut up to no w industrially not feasible.
Fig. 1.1 sho ws the distrib ution of the methanol demand in millions of metric tons
and is based on data obtained from [18]. The global production capacity is represented
by a solid line, while the dashed line refers to the a v erage annual operation rate of the
facilities. The capacity is forecasted to grow annually by an a verage rate of 5% to reach
approximately 140 million metric tons in 2020 [16, 19]. W ithin the last decade, China
became by far the biggest consumer of methanol [19]. Furthermore, the highest growth
rates are predicted for the Chinese market [20]. The lo w operation rate of 63% in 2015 is
forecasted to increase to 80% in 2020. This improv ement is caused by the decommis-
sioning of small coal-based plants and their substitution by modern gas-based lar ge-scale
plants. The graphs sho w that the a v ailable production capacity at high operation rates is
suf ficient to co v er the worldwide methanol demand.
2

2000 2005 2010 2015 2020 2025
0
20
40
60
80
100
120
140
160
Production capacity [10 6 mt]
0
20
40
60
80
100
Operation rate [%]
Japan South America Middle East
Eastern & Central Europe W estern Europe North America
Asia-P acific China Operation Rate
Figur e 1.1: Methanol demand forecast (based on [18]).
Methanol is produced worldwide in about 90 plants which are mainly located in Asia,
the Middle East, North and South America, and Europe [21]. In the recent past, a v ariety
of production facilities mo v ed from the consuming re gions to the countries of the ra w
material e xtraction, which thus could incorporate a further step in the v alue chain [3]. At
present the majority of the methanol plants in commercial operation has a capacity belo w
3,000 metric tons per day (mtpd) [6]. The increase in demand and the economies of scale
dri v e the progress to wards higher single plant capacities of up to 10,000 mtpd [22–24].
Depending on the re gion, the predominant feedstock for methanol generation is natural
gas or coal. The Middle East and Asia ha v e become the lar gest methanol-producing
re gions, due to ab undant reserv es of ra w materials [6]. While large-scale production
facilities are established in re gions with lar ge reserv es of lo w-cost natural gas (the Middle
East, North America), small coal-based plants are installed in the Asian-pacific region.
The major part of the global methanol production capacity is based on natural gas, due to
lo wer comple xity of the synthesis route, higher operation rate, and reduced in vestment
3

Chapter 1 Introduction
and operation cost [25]. Most likely , gas will continue to be the main feedstock in the
near future, due to recent adv ances in shale gas production [26, 27].
The construction of a v ariety of ne w plants with large capacities increases the com-
petition and in general the pressure on the market price. During the past 20 years, the
methanol market price w as subject to lar ge temporal and re gional v olatilities [28]. Con-
sidering the time period since 1975, the a v erage methanol wholesale price amounts to
175 US$/mt [25]. Between 1998 and 2018, methanol was traded on dif ferent mark ets
for 100 - 850 US-$ per metric ton. Generally , due to a v ariety of impact factors an
estimation and prediction of the production cost and market price is dif ficult. A study
conducted by Masih et al. [29] suggests that the natural gas price is the dri ving force for
the methanol price le v el in Europe and North America, while the prices on the Asian
continent are determined by the gro wing mark et of deri v ati ve products. Lastly , the
increasing a v ailability of shale gas in the United States resulted in a significant drop of
natural gas prices, leading to a higher profitability of the domestic methanol production
[6]. Independent of market conditions, the production cost can acti vely be reduced by
the selection of suitable technology and the associated adv antages through upscaling.
Dimethyl Ether
8% Methylamines
3%
Chloromethanes
1%
Acetic Acid
8%
F ormaldehyde
25%
Methyl Metacrylate
1%
MT O/MTP
22%
MTBE / T AME
7%
DMT / Others
5%
Solv ents
4%
Gasoline / Fuel
16%
Figur e 1.2: Methanol product distrib ution in 2016 (based on [25]).
Methanol plays an inte gral role in the global economy . Since it is the starting material
for man y chemical industries, the profits of a v ariety of deri v ati ve markets, and in
general the a v ailability of products, directly depend on the price of the feedstock [6].
The industrial applications range from the further processing to b ulk chemicals, e.g.,
4

formaldehyde and acetic acid, to the production of synthetic fuels [18]. The piechart in
Figure 1.2 sho ws the distrib ution of methanol applications in the year 2016 based on
data obtained from [25]. About 69% of methanol is produced for the chemical mark et,
wherein the focus in on the production olefins (alkenes) and formaldeh yde as do wnstream
intermediate chemicals. In particular , the methanol-to-olefins process (MT O), which
allo ws the production of consumer plastics lik e polypropylene and polyethylene, is
associated with high gro wth rates. As there are other competing processes for their
production (e.g. steam cracking), the methanol price is crucial for the spread of the MT O
technology in this se gment. Formaldeh yde as the lar gest single consumer of methanol is
used for the production of plastics and resins, pharmaceuticals, chemical fibres, paint
and pesticides [6, 30].
In addition to the significance as a basic chemical, methanol is required for the
production of a v ariety of synthetic fuels lik e dimethyl ether (DME), methyl-tert-
b utylether (MTBE), T AME (tert-amyl methyl ether) and gasoline [6, 16, 17]. Fur -
thermore, methanol/gasoline blends also ha v e a lar ge significance for the transport sector .
Approximately one third of the total methanol production is processed to synthetic fuels,
whereby the relati v e share of end-products in total methanol use stagnated or slightly
decreased o ver the past decade [31]. For an impro v ed o v ervie w , the flo wsheet of the
entire methanol v alue chain is presented in Figure 1.3.
Against the background of an ener getic use, the properties of easy storage and trans-
portability make methanol a f a v ourable and competiti v e ener gy carrier . In contrast to
electricity , methanol can be transported o v er long distances without significant ener gy
losses. At the consumer , methanol can be directly con verted into electricity by means
of a direct methanol fuel cell (DMFC). Moreo ver , despite of the recent advancements
of battery technology , there is still no lar ge-scale storage of electricity possible [3]. In
re gard to crude oil and other fossil fuels, methanol has the potential to displace these
materials, since the entire chain of hydrocarbon fuels can be deri v ed from methanol.
Since methanol can be produced from an y hydrocarbon source, price imbalancies and
shortages in con ventional fossil fuels can be used to gain opportunity re venues for the end
products. Furthermore, ethanol and hydrogen are often considered as a direct competitor
to methanol.
The main adv antage of methanol o v er ethanol is that the latter relies e xclusi v ely on
agricultural resources. Re garding a hydrogen economy , substantial disadv antages refer to
a high v olatility , explosion risk and limitations in terms of storage and transportation [3].
F or a more ef ficient, economically and ecologically more viable use of resources in
the methanol production industry , three main goals can be defined:
5

Chapter 1 Introduction
1.
Inno v ati v e and more ef ficient processes based on indirect synthesis routes with
a special focus on ne w reforming and alternati ve gasification technologies for a
reduction of capital in vestment and fuel cost.
2.
Further de v elopment of sustainable production concepts for reduced fuel consump-
tion through chemical v alorization of carbon dioxide.
3.
Maximization of single train capacities to increase the economic attraction through
scale ef fects.
6

Coal W aste Biomass Heavy Oil CO 2 Biogas Natural
Gas
Naphta
Gasification Reforming CO 2 -
Capture
W ater Gas Shift H 2 -addition CO 2 -addition H 2 -separation CO 2 -separation
Ammonia
Synthesis
DME Direct-
Synthesis
Methanol
Synthesis
Fischer -T ropsch
Synthesis Methanation
Crude Syngas
Syngas Conditioning
Synthesis
Chemical End-Pr oducts Energetic
End-Pr oducts
Methanol SNG
DME
Ammonia W ax, Diesel, etc.
Gasoline
MTBE
DME
Acetic Acid F ormaldehyde
Solv ents
Methylformate Olefins
Methyl
Metacrylate
Methylamines
Chrloromethanes
Olefins
Figur e 1.3: Overvie w of the methanol value chain.
7

Chapter 1 Introduction
Scope of w ork
The focus of this research is to approach the objecti v es of a more ef ficient, ecologically
friendly and economically more viable methanol production. An endless v ariety of
synthesis routes for the indirect con version from natural g as has to be considered in
order to reach a final plant design. Profound in vestigations are performed for the syng as
generation unit and the synthesis reactor technology , since the operation of these units is
decisi v e for the selection of a plant design, the ov erall performance and consequently
also for the economics of a process. By application of sensiti vity analyses, fa v ourable
operation parameters for high con version and lar ge yield are determined. Finally , the
analyses enable for a synthesis of adv anced designs with impro v ed ef ficiency . T wo reactor
technologies dominating the methanol market are in vestigated - the isothermal reactor
design and the adiabatic reactor technology with direct and indirect cooling. T aking
into account the findings, six final process designs are subject to further in vestig ations.
These include three basic processes which use a single reformer technology for medium
capacities of up to 3,000 mtpd. Furthermore, three adv anced processes with mixed
reforming technology and syngas conditioning for higher capacity are e xamined. First,
the o verall process performance is in vestigated by k e y operating parameters, such as the
methanol yield, the rec ycle-to-feed-ratio and the methane intensity . Four process designs
inte grate
CO 2
for utilization at dif ferent locations of the synthesis routes. The utilization
strate gies are quantitati v ely assessed in terms of
CO 2
inte gration potential, ov erall
CO 2
abatement and the sa ving of fuel. Subsequently , an ex er getic analysis is conducted
for all designs to identify the main sources and the magnitude of the irre v ersibilities.
Furthermore, an economic analysis serv es for an in vestigation of the influences of
e xternal market conditions on the re venues and the contrib ution mar gin and sho ws the
interactions of the wholesale prices of generated products. The combination of these
two analyses in an e x er goeconomic analysis allo ws an identification of the internal cost
distrib ution within the components of each process. The analysis results in a calculation
of the le v elized cost of the generated products methanol and electricity . Subsequently ,
the impact of economic uncertainties on the product cost is studied. Finally , an adv anced
e xer getic analysis is conducted for an outstanding two-step reforming process to analyse
e xemplary the interactions and the impro v ement potential of the components.
8

Chapter 2
State of Resear c h
In general, Gas-to-Liquids-processes (GtL) can be di vided into direct and indirect liq-
uefaction processes. Indirect processes in v olv e the intermediate step of producing a
synthesis gas (syngas), consisting mainly of
CO
and
H 2
, before this mixture is further
processed to liquid end-products by means of catalysis. Large-scale applications of
methanol generation are e xclusi v ely based on the indirect synthesis route [25]. These
processes (in particular the syngas generation) are ener gy intensi v e and costly in terms
of in vestment and operation. The disadv antages ha v e led to studies on processes that
allo w for a direct synthesis of methanol, a v oiding the undesired step of syngas generation.
Although direct con version has been thoroughly in vestigated o v er the past fe w decades,
there are still barriers to an industrial implementation. The main obstacle for their spread
refers to weak characteristics (lo w selecti vity to
CH 3 OH
, lo w con version of
CH 4
) of
the catalyst materials which make an upscaling too cost-intensi ve or impossible. The
processes for direct con version of
CH 4
to methanol are briefly gi v en in the follo wing list:
1. Direct oxidation of methane to methanol [32–35]
2. Catalytic gas-phase oxidation of methane [25, 36]
3. Liquid-phase oxidation of methane to methanol [25, 37]
4. Methane into methanol con version through monohalogenated methanes [25, 38]
5. Microbial or photochemical con version of methane into methanol [25]
An o vervie w of v arious indirect synthesis routes for the production of methanol is
gi v en in Fig. 2.1. The processes dif fer in terms of syngas generation and conditioning
applications, synthesis reactor technology and crude product purification. The processing
steps of an indirect synthesis route for methanol production from natural gas will be
introduced in the follo wing.
9

Chapter 2 State of Research
2.1 Syngas Generation fr om Natural Gas
A v ariety of technologies is a v ailable to produce a syngas from natural gas. The choice of
an adequate syngas production technology has to consider a number of f actors: feedstock
a v ailability and composition, cost of the feedstock and the fuel, plant location, comple xity
of inte gration with existing f acilities, required reliability and en vironmental constraints
and capital cost considerations. The common objecti v e to all syngas production tech-
nologies is the provision of a synthesis g as with a stoichiometric composition for the
respecti v e synthesis.
Methanol in general can be produced with high selecti vity (99.9%) via h ydrogenation
of
CO
or
CO 2
, as shown in reaction Eqs. 2.25 and 2.26. The reaction mechanism is
not definiti v ely understood since there is a disagreement whether methanol is mainly
produced via reaction 2.25 or 2.26 [39]. Three dif ferent parameters are used to describe
the stoichiometry and reacti vity of a syng as and thus also the ef ficienc y of the synthesis.
Se v eral studies characterize the composition of a syngas simply by using the
H 2
/
CO
ratio, which has ideally a v alue of 2 with respect to the synthesis of methanol [40].
This approach ne glects the hydrogenation of
CO 2
via Eq. 2.26 in the stoichiometric
consideration. Although the methanol synthesis mainly proceeds ov er
CO
, the reaction
is 100 times faster in the presence of
CO 2
[41]. T aking into consideration the role of
CO 2
, an extension of the
H 2
/
CO
ratio by a second parameter referred to as stoichiometric
module S is proposed to describe the composition for co vering the stoichiometry of
both reactions [3, 42–45]. W ith respect to the synthesis of methanol ov er commercially
copper -based catalysts, the S module ideally has a v alue of slightly abo v e 2. T ypically ,
the feed gas composition for the methanol synthesis contains 4-10 mole-%
CO 2
for high
con version and yield [6, 46]. Using an ener gy-intensi v e acid gas remo v al (A GR), the
CO 2
-content can be further decreased to 1 - 3 mole-% to achie v e maximum methanol
yield [39].
Stoichiometric module S
S = ˙ n H 2 − ˙ n CO 2
˙ n CO + ˙ n CO 2
(2.1)
Commercially used catalysts sho w a high acti vity in reaction equation 2.27, whereby
CO
is in v olved in the re verse w ater gas shift (R WGS). Dybkjaer et al. [26] oppose that
the stoichiometric module S is not adequate to describe the reacti vity of a syngas in
methanol synthesis, since the module is independent of the w ater gas shift reaction 2.27.
10

2.1 Syngas Generation from Natural Gas
Instead the y suggest to use the
CO
/
CO 2
ratio for this purpose. A high ratio increases the
reaction rate and the per pass con version in methanol synthesis, since the
CO
is preferred
ove r
CO 2
as a reactant by copper based catalysts. Moreov er the production of water is
decreased, which prolongs the lifetime of the catalyst and reduces the water content in
the crude product.
A less common measure to describe the stoichiometry of a syngas is gi v en by the
stoichiometric ratio R in Eq. 2.2. A stoichiometric syngas for methanol synthesis has a
R v alue close to unity . Accordingly a carbon rich syngas has a R v alue belo w one, while
a hydrogen rich gas has a ratio abo v e unity .
Stoichiometric ratio R
R = ˙ n H 2
2˙ n CO + 3˙ n CO 2
(2.2)
Depending on the adopted ra w material, the applied syngas generation technology and
the used reforming agents, the composition and therefore the v alues of the stoichiometric
modules can widely v ary [47]. The most widely used technology for the con version
of natural gas to a synthesis gas in industrial methanol applications is steam reforming
by
H 2 O
(SMR), including gas heated reforming (GHR) and heat e xchange reforming
(HER). Furthermore, autothermal reforming (A TR), two-step reforming, and to a minor
e xtent catalytic and thermal partial oxidation (CPO X and TPO X) also found commercial
application. Mix ed or combined reforming (CMR) by integration of v arious reforming
agents can be mentioned as a separate cate gory . The dry reforming (DRM) using
CO 2
as
a reactant is considered as a pre-commercial technology . The v ariety of dif ferent process
routes is sho wn schematically in Fig. 2.1.
Figur e 2.1: Schematic of the synthesis routes for the production of methanol from natural gas.
11

Chapter 2 State of Research
The process units of each of these synthesis routes can be assigned to six major
subsystems: a pretreatment unit for the remo v al of sulfur impurities, a reforming unit
serving for the syngas generation, a conditioning section for syngas adjustment, a
synthesis unit for the con version of syng as to methanol, a purification section for refining
the crude methanol, and a steam cycle for balancing the o v erall heat and electricity supply
and demand of the process units. In the follo wing detailed information will be pro vided
for the technologies of the subsystems.
2.1.1 Pretreatment
Natural gas typically contains se v eral or g anic sulfur species, which are potentially
poisonous to the catalysts used in methanol production processes [48]. The catalysts
used in the reforming unit and the synthesis become quickly deacti v ated in their presence,
which decreases the performance of the o v erall plant. An upper sulfur le v el of 0.002 ppmv
for the processing of natural gas in reformers is recommended in [49]. Most common
applications use a serial arrangement of two absorption beds to decrease the sulfur le vel to
an acceptable le v el. The or ganic sulfur components, particularly carbon yl sulfide (
COS
),
are first hydrolysed to hydrogen sulfide (
H 2 S
) in a temperature range of 250 - 400
◦
C,
and a pressure range of 5 - 40 bar [50] (reaction Eq. 2.3). T ypically , the hydrogenation
step is carried out o ver nick el or cobalt molybdenum catalysts. Subsequently , the
H 2 S
is
absorbed on a zinc oxide bed to produce zinc sulphide ZnS according to reaction 2.4.
COS + H 2 O  CO 2 + H 2 S (2.3)
ZnO + H 2 S  ZnS + H 2 O (2.4)
2.1.2 Pre-ref orming
Natural gas inherently contains long-chained hydrocarbons, which can be beneficial for
the production of methanol b ut detrimental to the operation of the reformer . Higher
hydrocarbons are characterised by a high ratio of carbon to hydrogen and allo w a
generation of a stoichiometric more suitable synthesis gas than obtained by steam
reforming of methane. An adiabatic reformer prior to the main reforming unit permits
the use of hea vier feedstock and reduces the load on the main reformer , thus allo wing
higher throughput and impro ved ef ficiency [51]. The pre-reformer is operated within
a temperature range of 350 - 550
◦
C, which requires a preheating of the two reactants
12

2.1 Syngas Generation from Natural Gas
steam and natural gas [52]. The mixture passes through a bed of catalyst pellets, which
are typically based on nickel that is dispersed on an alumina stabilizer . The reactions
are carried out adiabatically , whereby all higher hydrocarbons are con verted to methane.
A portion of the methane is further reformed to hydrogen and carbon monoxide. The
process gas cools as the endothermic reactions proceed, and the pre-reformer ef fluent is
then reheated before entering the main reforming section. The inclusion of a pre-reformer
reduces the size of main reforming unit, since all energy added to the pre-reformer does
not ha v e to be supplied to the follo wing reformers. For e xample, the reduction in a
tub ular steam methane reformer duty is 10 - 15% corresponding to about 75 reformer
tubes.
2.1.3 Hydr ocarbon Ref orming
The reforming unit in general has a high impact on the o v erall process performance. A
careful selection of a specific technology based on defined criteria therefore is essential
for the o verall thermodynamic performance and the economics of a plant. After the
reforming technologies will be introduced in this chapter , a v ariety of selection criteria
will be outlined which assist in finding a suitable reforming technology .
2.1.4 Steam Ref orming
Catalytic steam reforming of methane (SMR) is the industrially most common used
technology for synthesis gas production from natural gas. Large-scale applications not
only refer to the production of methanol, but also include the manuf acture of other
precursors like h ydrogen and ammonia. In SMR, natural gas is con verted in highly
endothermic and irre v ersible reactions (see Eqs. 2.5 and 2.6 ) with steam ov er catalyst
pellets, typically based on
Ni/Al 2 O 3
, at high temperatures and lo w pressures [53]. The
process conditions are in the range of 20 - 40 bar with inlet temperatures of 300 - 650
◦
C and outlet temperatures of 700 - 1000
◦
C [33, 47, 52]. A lar ge amount of steam
according to a steam-to-carbon mole ratio (S/C ratio) of 3-5 is supplied to achie v e a
high con version of natural g as. This typically results in lar ge flo w rates do wnstream the
reforming unit. The stoichiometry of the product gas is highly sensiti ve to a v ariety of
factors, such as the reaction temperature, pressure and the S/C ratio. The SMR reaction
in Eq. 2.5 is ine vitably accompanied by the w ater -gas-shift in Eq. 2.6. In particular for
a high operation temperature, the con v ersion via the WGS is reduced, decreasing the
amount of undesired CO 2 in the product gas [47].
13

Chapter 2 State of Research
CH 4 + H 2 O  CO + 3H 2 ∆ H 298K = 206 kJ
mol (2.5)
CO + H 2 O  CO 2 + H 2 ∆ H 298K = − 42 kJ
mol (2.6)
The endothermic character of the main reaction requires an intense heat supply to the
reaction zone. Industrial SMR is commonly performed in a fixed bed reactor , consisting
of a v ariety of catalyst-filled tubes, which are located in the radiant section of a firebox.
The process heat is supplied by a number of b urners, which are fuelled with natural
gas and a v ent gas rec ycled from the synthesis. T o ensure an ef fecti v e heat transport
in radial direction, the tubes are characterized by a low diameter -to-height ratio. The
tubes typically ha v e a length of 10-14 m and an outer diameter of 10-12 cm [26]. The
v ariety of a v ailable firebox designs dif fers in terms of b urner configuration - the y can be
located on the roof, on the bottom, on lev elled terraces at the w alls or as side b urners
on the walls. The design has an implication on the heat flux and the temperature profile
and consequently also on the operation and the material stresses. Further information on
dif ferent firebox designs is reported in [54, 55].
Since only about 50% of the fired ener gy is directly absorbed by the tubes, a heat
reco very section is located do wnstream of the refractory lined zone. The thermal ef fi-
cienc y of the entire unit is increased to approximately 90%, by using the waste heat for
preheating the reactants and for electricity generation in a steam c ycle. An ef fecti v e heat
reco very reduces the flue g as temperature to about 150 ◦ C, whereby v alues belo w
100 ◦ C are possible in case of lo w sulfur species content.
A typical composition of a syngas obtained by SMR has a stoichiometric module of S
equals 3 and yields a
H 2
/
CO
ratio of the same range. The composition is far a w ay from
the desired stoichiometry required by the methanol synthesis (S = 2.05 and
H 2
/
CO
ratio
= 2). Due to the hydrogen surplus, a lar ge amount of syng as is carried out unreacted from
the synthesis and therefore agglomerates as a ballast in the rec ycle loop [6]. Unlike other
reforming technologies, the product gas from SMR has a lo w content of carbon species,
in particular
CO 2
. Therefore, inte gration of
CO 2
as a conditioning measure to increase
the methanol yield by direct hydrogenation via reaction 2.6 is discussed as an option for
carbon utilization [3, 56]. T ypically , about 1 mole of
CO 2
has to be added to 4 moles
of natural gas feed to achie v e a nearly stoichiometric syngas as it enters the methanol
loop [42]. Ho we v er , according to Klier et al. [57] an optimal syngas composition for
methanol synthesis is fixed for a
CO 2
mole fraction of 2%. At much higher
CO 2
-le v els,
the methanol yield would be reduced.
14

2.1 Syngas Generation from Natural Gas
Furthermore, other syngas conditioning measures could be applied to achie v e an
H 2
/
CO
ratio
=
2. Since typical ratios for SMR are f ar abo v e this v alue, a composition adjustment
by re v erse w ater gas shift (R WGS) or a remov al of
H 2
by pressure swing absorption
(PSA) would be beneficial. A conditioning by blending the syngas with
CO 2
upstream
a R WGS unit increases the amount of
CO
and
H 2 O
to adjust the
H 2
/
CO
ratio to the
stoichiometric requirements. Ho we v er , this carbon utilization measure can only be
applied for a lo w content of CO 2 in the synthesis gas.
2.1.5 Heat Exchange Ref orming / Gas Heated Ref orming
The heat e xchange reforming technology was de veloped to reduce the in vestment cost of
syngas-based processes [6]. In a heat e xchange reformer , natural gas and other reactants
are e xposed to con vecti ve heat that is supplied by a flue gas or process gas. The design
is similar to that of a tub ular gas/gas heat e xchanger . In contrast to fired reformers, the
heat within heat e xchange reformers is mainly transferred by con vection. V arious types
of heat e xchange reformers ha v e been de v eloped by se v eral companies [52, 54] in the
80s and 90s. The technologies dif fer in terms of the heat source (process gas or flue gas),
the tube design (bayonet or straight through), the maximum heat flux, and in operation
parameters. A schematic of the three main heat e xchange reformer types is pro vided in
Figure 2.2. Further information on the dif ferent designs is reported in [54].
Figur e 2.2: T ypes of heat exchange reformer a) with "straight-through" tubes, b) with bayonet
tubes, c) with mixing of heating gas and product gas before heat e xchange.
15

Chapter 2 State of Research
The GHR technology can be combined with an autothermal reformer (A TR, see Section
2.1.8) or oxygen blo wn reformer in serial or parallel configurations. A GHR based on
SMR is often used in re v amp projects with a stand-alone A TR to increase the plant
capacity and the carbon ef ficienc y [52].
F or a combination of GHR and A TR, the endothermic reactions within the primary gas
heated steam reformer consume the excess heat released by the secondary e xothermic
A TR. The heat is transferred by recycling the hot ef fluent from the A TR to the shell-side
of the GHR. In two-step reforming concepts the majority of the methane feed is con verted
within the A TR (60 - 80%) while the remainder is completed by the GHR. The main
adv antages of a GHR (in tw o-step reforming) o v er con ventional tub ular SMR refer to a
small size, lo wer in vestment costs and reduced steam utility . The technological progress
dro ve the S/C ratio do wn to v alues of 0.6 - 1.5. The GHR types a) and b) can be used in a
serial or parallel arrangement, while type c) only can be used in a parallel configuration.
The gas heated reformer technology has been successfully demonstrated on a lar ge scale
in the Sasol Synfuel Complex in Secunda, South Africa, and in a GtL plant in Nigeria
[52, 58].
2.1.6 P artial Oxidation
The carbon content of a syngas produced by con v entional steam reforming technology is
belo w the stoichiometric requirement of the methanol synthesis. The surplus in hydrogen
requires an addition carbon components, which can either be supplied by an inte gration of
a rec ycled
CO 2
stream or alternati v ely by application of the partial oxidation technology
(PO X). PO X refers to a substoichiometric combustion, which can be carried out with
or without the use of a catalyst (catalytic partial oxidation CPO X and thermal partial
oxidation TPO X) [59, 60]. The operation can be conducted in a wide temperature range
of 800 - 1500
◦
C at ele v ated pressures [59, 61, 62]. Eqs. 2.7 - 2.9 sho w the main reactions
occuring in the partial oxidation of methane.
CH 4 + 0 . 5 O 2  CO + 2H 2 ∆ H 298K = − 36 kJ
mol (2.7)
2 CO  C + CO 2 ∆ H 298K = − 172 . 5 kJ
mol (2.8)
H 2 + 0 . 5 O 2  H 2 O ∆ H 298K = − 241 . 8 kJ
mol (2.9)
16

2.1 Syngas Generation from Natural Gas
The partial oxidation of methane is a highly e xothermic process, which yields a syngas
with a lo w
H 2
/
CO
ratio of around 1.6 - 2 (see reaction Eq. 2.7). The composition of the
process gas can be adjusted by controlling the mole ratio of oxygen to methane [59],
which is also termed as O/C ratio. Depending on the operation parameters, the PO X
reaction is accompanied by undesired carbon deposition according to the Boudouard
reaction in Eq. 2.8. The excess heat of the hot process gas may be reco v ered as steam for
electricity generation or for supply to other units [3]. In this study , only the non-catalytic
partial oxidation is considered as a part of the autothermal reforming unit.
2.1.7 A utothermal Ref orming
The use of autothermal reforming (A TR) for syngas generation in methanol and am-
monia production started in the middle of the
20 th
century . Since then, A TR has been
incorporated in v arious application fields for the manuf acture of a v ariety of products.
T oday , lar ge scale applications can be found in Europe (2002), South Africa (2004),
Qatar (2006) and other parts of the world [26, 63]. The A TR technology is preferentially
used for methanol production capacties abo v e 5,000 mtpd due to the economies of scale
ef fect. In 2012, an A TR based methanol plant with a single-train capacity of 10,000 mtpd
started operation in Nigeria [23].
A TR mer ges steam reforming and partial oxidation in one compact pressure vessel. As
sho wn in Figure 2.3, the reactor consists of a comb ustion chamber (comb ustion zone) and
a fix ed catalyst bed within a refractory lined pressure shell (catalytic zone) [56, 64]. The
b urner arrangement allo ws a soot-free comb ustion, a homogenous gas and temperature
distrib ution, and protects the v essel and nozzles from the hot flame core.
The hydrocarbon feedstock is blended with steam and reacts substoichiometrically
with highly concentrated oxygen in a v ariety of comb ustion reactions that proceed along
with steam reforming and the shift reaction [26, 64]. The set of reactions carried out
within the comb ustion and catalytic zone of the A TR is sho wn in Eqs. 2.10 - 2.12.
17

Chapter 2 State of Research
Figur e 2.3: Design of a autothermal reformer (o wn presentation based on [65]).
Comb ustion zone:
CH 4 + 1 . 5O
2  CO + 2H
2 O ∆ H 298K = − 520 kJ
mol (2.10)
Comb ustion and catalytic zone:
CH 4 + H 2 O  CO + 3H
2 ∆ H 298K = 206 kJ
mol (2.11)
CO + H 2 O  CO 2 + H 2 ∆ H 298K = − 42 kJ
mol (2.12)
The partial oxidation of methane (reaction Eq. 2.10) can cause an increased tempera-
ture up to 3000
◦
C within the flame core in the upper part of the comb ustion zone. The
v essel temperature is limited by the thermal stability of the catalyst and of the interior
refractory lining [6]. After leaving the comb ustion zone, the mixture is equilibrated
in the catalytic zone. The size and shape of the catalyst pellets are optimized for high
acti vity and a lo w pressure drop to enable a compact design. Small sintering ef fects on
the catalyst material are tolerated.
Industrial autothermal reformers are operated in a pressure range of 30 - 50 bar with
outlet stream temperatures between 950 - 1100
◦
C [6, 66]. Due to the higher suction
pressure, the electricity demand of the syngas compression is only about 50% compared
to the ener gy demand in a con ventional plant using steam reforming [6]. In oder to
18

2.1 Syngas Generation from Natural Gas
keep the temperature in a moderate range and to allo w a higher hydrogen content, steam
is injected to initiate the endothermic steam methane reforming according to reaction
Eq. 2.11. The A TR offers great fle xibility in terms of operation, which adjusted to the
respecti v e do wnstream application. A TR is operated with a lo w S/C ratio, thus reducing
the flo w rate through the plant and dimishing the in vestment cost. Manuf acturers of A TR
report a steam-to-carbon mole ratio (S/C ratio) of 0.5 - 1.5, while the oxygen-to-carbon
mole ratio (O/C ratio) is in the range of 0.6 - 1.0 [22, 66].
The process gas produced by A TR typically is of high temperature and therefore
requires intensi v e cooling before further processing. The thermal energy is used for
medium pressure steam generation. W ith respect to the composition, the syngas module
S is in the range of 1 - 2, and has a CO/CO 2 ratio of 2.5 - 5.5 [22, 56, 66].
The fraction of uncon verted methane in the ef fluent (methane slip) is belo w 3 mole-%
[65]. The high carbon monoxide content results in a high reacti vity of the syngas. Even
though the high
CO
content makes the syng as highly reacti v e, further syngas conditioning
steps might be necessary to increase hydrogen content while lo wering the carbon dioxide
fraction. Approaches to increase the hydrogen content are typically realized by hydrogen
reco very from the pur ge gas within the synthesis loop through application of a pressure
swing absorption unit [56, 66]. Other possibilities refer to a WGS unit which can be
used in conjunction with carbon dioxide remov al by application of industrially prov en
physical absorption processes (Sele xol R
 , Rectisol R
 ).
2.1.8 T w o-step Ref orming
T wo-step reforming refers to a pro v en process that has been applied commercially by
se v eral companies for the production of a v ariety of products [56]. The reforming concept
comprises a serial or parallel arrangement of a primary tub ular steam reformer and an
oxygen enhanced secondary reformer (PO X) or alternati v ely an autothermal reformer
[35, 42, 56]. The primary reformer can be designed as a fired reformer or a gas heated
reformer , using the ef fluent of the secondary reformer as a heat source (see Section 2.1.5).
The two basic configurations are presented in Fig. 2.4.
In the serial arrangement (2.4 a)), the entire fuel stream is fed to the inlet of the
primary reformer and passes through both reactors. This concept is often used for
retrofitting con ventional plants based on SMR to debottleneck the reforming unit. In
man y configurations, the hot ef fluent from the secondary reformer is recycled to the
primary reformer for heat supply to the SMR (see Section 2.1.5). The primary reformer
can be operated with a significant leakage of methane, thus requiring a lo w S/C ratio and
19

Chapter 2 State of Research
a lo w operation temperature. Compared with stand-alone SMR, a reduction of 60% in
the transferred duty and of 75 - 80% in the reformer tube weight is obtained. The sa vings
thus obtained more than compensate the in vestment e xpenditures for the air separation
unit, which is required to supply oxygen to the A TR [67]. Compared to stand-alone A TR,
the demand for oxygen and therefore the size of the ASU is reduced due to lo wer O/C
ratio and partial con version of methane in the primary reformer . The main disadvantage
of the serial configuration refers to the infle xibility of the operation pressure due to
direct coupling of the reactors. The syngas composition can be controlled by the outlet
temperature of the primary refomer , which is typically operated in a low temperature
range of 700 - 800 ◦ C and pressures of 35 - 40 bar [66].
In the parallel arrangement (Fig. 2.4 b)), the fuel is di vided into two fractions whereby
the first fraction is reformed at high pressure and moderate temperature in the SMR unit.
Subsequently , the ef fluent from the primary reformer is mix ed with the second fraction
of the feedstock before supplying it to the secondary reactor . A small component size
and an independent setting of the operation parameters are the major benefits of the
parallel arrangement. The main disadv antage refers to the need of additional components
compared to the serial arrangement.
In general, the combination of endothermic SMR and e xothermic A TR can ensure a
thermoneutral generation of a syngas with f a v orable (nearly stoichiometric) composi-
tion for high methanol yield [3, 42]. Furthermore, two-step reforming concepts allo w
increased fuel fle xibility with reg ard to the production of a stoichiometric synthesis gas.
The methane slip of two-step reforming concepts is relati vely lo w , thus reducing the
ballast of uncon verted
CH 4
in the synthesis unit. Ex emplary plants based on two-step
reforming started up in Norway in 1997 (2,400 MTPD) and in Saudi Arabia in 2008
(5,000 MTPD)[56].
20

2.1 Syngas Generation from Natural Gas
Figur e 2.4: Serial a) and parallel b) configuration of the two-step reforming process (o wn
presentation based on [54]).
21

Chapter 2 State of Research
2.1.9 Dr y and Mix ed Ref orming
Dry reforming of methane (DMR), also termed as
CO 2
-reforming, is a well kno wn
con version process which attracts man y research ef forts due to v arious en vironmental
and economic incenti v es [6, 12, 68–70]. The technology is considered a promising
carbon utilization concept which shall be used primarily for the production of methanol
and urea. Incenti v es for the implementation not only refer to the mitigation of
CO 2
-
emissions b ut also aim at reducing the amount of used feedstock to decrease the operation
e xpenses. Despite these attracti v e features, there are no commercial applications for pure
DMR. The major disadv antages are related to the poor catalyst performance and the
ener gy intensity of the reforming process [69]. F or instance, the import of
CO 2
may need
compression and purification of an e xtra feedstock, which requires further unit operations
and a higher in vestment. According to Holm-Larsen et al. [71]
CO 2
-reforming is only
economically feasible at locations where a lar ge and relati v ely pure amount of methanol
is accessable.
The con version of methane by carbon dioxide is ine vitably associated with catalyst
deacti v ation by carbon deposition. DMR is a highly endothermic reaction, which requires
an operation temperature in the range of 800 - 1000
◦
C to attain satisfying equilibrium
con versions and to minimize the dri ving force for undesired carbon production [68,
69]. Catalysts used in DMR are typically based on
Ni
(e.g.
Ni/MgO
,
Ni/MgAl 2 O 4
),
whereby other nobel metals are preferable with respect to carbon desposition, b ut still
are uneconomically in lar ge quantities as used in industrial applications. Reaction Eq.
2.13 sho ws that DMR generates a syngas with an equimolar amount of
CO
and
H 2
[68].
CH 4 + CO 2  2CO + 2H 2 ∆ H 298K = 247 kJ
mol (2.13)
Dry reforming can also be combined with steam reforming (also named combined
reforming) and partial oxidation to adjust the syngas composition to w ards the stoichio-
metric requirement of the synthesis. A combined use of the reforming agents
O 2
,
CO 2
,
and
H 2 O
is termed mix ed reforming or bi-reforming and tri-reforming, respecti v ely . An
addition of
O 2
is typically used to compensate the heat deficit of the DMR in reaction
Eq. 2.14 by the e xothermicity of the oxidation reaction in Eq. 2.15.
CH 4 + 0 . 5O 2  CO + 2H 2 ∆ H 298K = − 36 kJ
mol (2.14)
CH 4 + 2O 2  CO 2 + 2H 2 O ∆ H 298K = − 802 kJ
mol (2.15)
22

2.1 Syngas Generation from Natural Gas
The injection of
H 2 O
is a tool to increase and control the hydrogen content of the syn-
thesis gas from mix ed reforming. The process gas obtained from dry reforming typically
is lean in hydrogen, being unsuitable as a metgas (syng as for methanol synthesis). The
supply of steam initiates the SMR, which can compensate for the deficit of hydrogen.
The supply of both reforming agents therefore most likely will result in a syngas with a
S module of equal to two. The catalysts applied to combined reforming hav e the same
characteristics as in SMR.
CH 4 + H 2 O  CO + 3H 2 ∆ H 298K = 206 kJ
mol (2.16)
3 CH 4 + CO 2 + 2 H 2 O  4 CO + 8 H 2 ∆ H 696K = 206 (2.17)
DMR inherently is accompanied by the decomposition of methane as well as the
Boudouard reaction to produce harmful carbon on the catalyst surface as sho wn by Eq.
2.18 and Eq. 2.19. In addition to a careful selection of operation parameters, the supply
of
H 2 O
in bi- and tri-reforming is a successful measure to a v oid carbon formation and
therefore catalyst deacti v ation [12].
CH 4  C ( s ) + 2 H 2 ∆ H 298K = 75 k J
mol (2.18)
2 CO  C ( s ) + CO 2 ∆ H 298K = − 172 . 5 kJ
mol (2.19)
In case steam is used as a reforming agent, the R WGS reaction simultaneously is
initiated, decreasing the
H 2
/
CO
ratio to a v alue belo w one [68]. The R WGS is only
present at temperatures belo w 820
◦
C while DMR proceeds in the fo w ard direction abo ve
640
◦
C [72]. Carbon formation according to reaction 2.18 and 2.19 occurs abov e 560
◦
C
and belo w 700
◦
C, respecti v ely . W ithin this temperature range, carbon formation reaches
a maximum due to simultaneous occurence of both mechanisms. Thus, bi-reforming
should also be carried out abo ve temperatures of 800
◦
C to obtain a syngas with an
impro ved composition for the methanol synthesis.
2.1.10 Pr ocess Selection Criteria f or Methanol Generation
A careful selection of a suitable reforming technology is essential for the thermodynamic
ef ficienc y and the profitability of a methanol plant. The syngas generation unit is the
most e xpensi v e unit, accounting for approximately 60% of the o v erall in vestment cost.
23

Chapter 2 State of Research
Moreo ver , the performance of the reforming unit determines the demand for fuel, which
constitutes the major cost contrib utor to the v ariable cost. According to [6], a cost
breakdo wn for a natural gas based methanol plant sho ws that the fuel cost account for 50
- 75% of the o verall cost. The follo wing selection criteria can be defined:
1. Single-train capacity
2. Economies of scale
3. Composition of the syngas product
4. Heat inte gration and co-generation of electricity
5. Cost of the feedstock
6. In vestment cost
7. A v ailability of utilities
The selection of a specific reforming technology primarily depends on the single-train
capacity and cost reduction ef fects associated with an upscaling of the unit. Sev eral
studies presented a capacity range demarcation for the use of reforming technologies
in methanol plants [6, 56]. The operation of SMR-based methanol production is eco-
nomically justifiable only for a single-train capacity of up to 2,500 mtpd [55]. Abo v e
this capacity , lar ge steam reformers become progressi v ely more e xpensi v e and thus
nearly sho w no economies of scale [6]. T wo-step reforming concepts are economically
beneficial in the intermediate capacity range of 2,500 to 5,000 mtpd. For lar ge-scale
capacities abo ve 5,000 mtpd, A TR is the preferred technology , as higher cost reductions
are associated with the upscaling of an air separation unit (ASU) than with tub ular re-
forming. Ho we v er , there are no absolute limits for the application of dif ferent reforming
technologies b ut certain ranges of preference [6]. The typical capacity range for the
reforming technologies is presented in Fig. 2.5. DMR and mix ed-reforming cannot be
sho wn as no applications e xist on an industrial scale. Generally , single PO X is not used
for syngas generation in the production of methanol, due to high cost associated with the
lar ge oxygen supply .
24

2.1 Syngas Generation from Natural Gas
SMR SMR+A TR A TR
024 6 8 10
0
10
20
30
40
50
60
70
80
90
100
Cost Index [-]
SMR
A TR
SMR+A TR
024 6 8 10 0
10
20
30
40
50
60
70
80
90
100
Capacity [10 3 mtpd]
Figur e 2.5: Capacity range of the commercial reforming technologies.
The reforming technology also has an impact on the size and the in vestment cost
of do wnstream components through the composition and the flo w rate of the provided
synthesis gas. Depending on the composition, the size of do wnstream components in the
synthesis and purification unit are determined and further syngas conditioning steps may
be required.
Although the methane slip is generally in a lo w range, small dif ferences in inert content
can ha v e a significant impact on the size of the synthesis unit as the inert components
agglomerate in the synthesis loop. Furthermore, the
CO 2
content of the process gas
(typically 2 - 10 %) on one hand acts as a reaction promoter and on the other hand causes
a production of undesired water . The high water content ine vitably leads to an increase
of the number and size of distillation columns and the heat duty required for product
purification. Since the syngas composition is considered as a ke y parameter for the
design layout and the operation of a methanol plant, special attention will be gi v en to
this selection criterion in Section 4.1.
Furthermore, the endothermic or exothermic character of the reaction mechanism
is decisi v e for the fuel supply , the heat integration management and e ventually a co-
25

Chapter 2 State of Research
generation of electricity . The operation of a SMR and DMR unit requires the supply
of high temperature process heat which is typically generated by comb ustion of natural
gas. A use of lo w temperature waste heat by recov ery as steam for co-generation of
electricity is essential for a high ef ficienc y of those processes. In contrast, exothermic
A TR can dispense with the supply of process heat and additional fuel. Ho we v er , process
inte gration of the A TR also requires and heat management system.
Significant dif ferences also e xist in terms of operating pressure. An energy-intensi ve
syngas compression do wnstream of the reforming unit is required to compress the syngas
to the pressure le v el of the synthesis unit. Depending on the reforming technology , the
ener gy demand of the multi-compressor unit significantly v aries. In particular for DMR,
the compressor unit constitutes a lar ge ener gy consumer , while the ener gy demand of a
compressor after a pressurized A TR is comparati v ely lo w . The cost of supplied feedstock
is the ke y to an economic success of a natural gas based methanol plant. In addition to
the cost for natural gas, the cost of the reforming agents oxygen and carbon dioxide also
ha v e a significant impact on the product cost. Economic adv antages for A TR and DMR
particularly occur , when oxygen is or carbon dioxide is a v ailable at lo w cost.
2.2 Conditioning of Synthesis Gas
2.2.1 Carbon Utilization Measures in the Pr oduction of
Methanol
T wo major concepts for the utilization of carbon dioxide in a methanol production
process via syngas production are of scientific and industrial interest [11, 44]. In general,
the utilization of
CO 2
is a promising approach to v alorize a w aste product for the
generation of v aluable fuels while reducing the
CO 2
emissions and the demand for fossil
feedstock. The
CO 2
may either be added as a reforming agent for syngas production
(see Section 2.1.9) or within the loop of the synthesis unit for direct hydrogenation in the
synthesis reactor [44, 68, 72–76]. From a mass balance point of vie w , the location of the
injection does not make a dif ference. Ho we v er , from the perspectiv e of the plant units
the dif ference is significant. For instance, the reaction rate in the methanol synthesis is
highly dependent on the
CO
/
CO 2
ratio and increases with increasing ratio. Therefore,
the synthesis unit significantly benefits when
CO 2
is added before the reformer , where
most of it is con verted to CO by the re verse water -gas-shift reaction. Other drawbacks
associated with dry reforming refer to the high cok e production and the lack of e xperience
26

2.2 Conditioning of Synthesis Gas
re garding the upsclaing of the technology [74].
Alternati v ely ,
CO 2
can be directly injected into the synthesis unit to produce methanol by
the catalysed hydrogenation according to the reaction in Eq. 2.26. The main adv antage
of this carbon utilization option can be seen in the reduction of the reformer load. Since
a highly concentrated
CO 2
stream is directly fed to the loop or the inlet of the syngas
compressor , less syngas needs to be produced in the reforming section. An injection
do wnstream of the syngas generation unit results in high
CO 2
mole fractions of the
reactor feed gas, which may decrease the product yield and the ef ficienc y of the o v erall
plant. This is caused by the lo w acti vity and selectivity of commercial three component
catalysts in the hydrogenation of
CO 2
. Thus, the potential for
CO 2
inte gration is limited
by catalytic constraints. Efforts ha ve been made to de v elop catalysts that are better
adapted to high
CO 2
feed. The proposed catalysts are based on
Cu
-
Zn
oxides (with
additi v es such as
ZrO 2
,
GaO 3
and
SiO 2
) and sho w a lo w selecti vity in re v erse water -
gas shift reaction while ha ving a high selecti vity for the h ydrogenation [74, 77]. This
direct hydrogenation constitutes a solution for the production of green methanol, when
the required hydrogen is produced by electrolysis, which is dri v en by electricity from
rene w able ener gy sources. In 2010, Carbon Rec ycling International [78] installed such a
system with an annual capacity of 3,000 t/a in Iceland.
CO 2
is supplied by an aluminium
production facility and a geothermal plant, while
H 2
is generated from an electrolysis.
Approximately 10 tons of methanol are produced from 18 tons of
CO 2
. Furthermore,
Mitsui Chemicals Inc. [79] installed a unit in Japan, having an annual methanol capacity
100 t [80]. Here, the
CO 2
is supplied as a waste product from an eth ylene production
plant.
F or an assessment of the carbon utilization measures, the analysed processes are
benchmarked ag ainst a con ventional process based on steam methane reforming. Some
ke y figures are introduced to e valuate the characteristics and the impact of the
CO 2
inte gration. These measures include the
CO 2
-emission intensity , the
CO 2
-feed intensity ,
the relati v e
CH 4
-reduction and the
CO 2
-abatement. In Eqs. 2.22 and 2.23 the subscript
0 refers to the benchmarking process. The o verall e x er gy of product
˙
E P ,tot
(see Section
3.1.2) is taken as a basis for comparison, since dif ferent products of different ener getic
quality are generated by the processes.
CO 2 − emission intensity = CO 2 , emis
˙
E P ,tot (2.20)
CO 2 − feed intensity = CO 2 , feed
˙
E P ,tot (2.21)
27

Chapter 2 State of Research
CH 4 − reduction = CH 4 , tot , 0
˙
E P ,tot,0 · ˙
E P ,tot − CH 4 , tot (2.22)
CO 2 − abatement = CO 2 , emis , 0
˙
E P ,tot,0 −
CO 2 , emis − CO 2 , feed
˙
E P ,tot (2.23)
2.2.2 (Rever se) W ater -gas Shift Reaction
Almost without e xception, the synthesis gas from the reforming unit does not meet
the stoichiometric requirements of the synthesis. SMR typically provides a syng as
with e xcess hydrogen, while DMR and A TR technology most likely generate a syngas
with a hydrogen deficienc y . An ef ficient and economic measure to adjust the syngas
composition to the corresponding stoichiometry of the synthesis refers to the use of a
water -g as shift unit. A distinction can be made between the clean gas con version and
the sour gas con v ersion. In clean gas con version, the unit is preceded by a purification
section to remo ve the major part of the sulfur components and higher h ydrocarbons that
may be present in the gas, while the sour shift accepts an untreated feed g as. Depending
on the temperature le v el, a second distinction is made between the high temperature shift
con version (300 - 500
◦
C) and the lo w temperature shift con v ersion (180 - 280
◦
C) [39].
The lo w temperature shift is only applied to obtain a syngas with v ery lo w
CO
content,
which is not of interest for the synthesis of methanol.
The water -g as shift equilibrium defines a firm relationship between
CO
,
CO 2
,
H 2
and
H 2 O
which is independent of the pressure and determined by the temperature. The
reaction sho wn in 2.24 is typically catalysed via iron oxide with heat release.
CO + H 2 O  H 2 + CO 2 ∆ H 298K = − 41 . 47 kJ
mol (2.24)
The temperature range of the shift con version is determined by the thermal resistance of
the catalyst material and the acti v ation energy of the reaction. In case of a
CO
deficienc y ,
as gi v en for SMR and possibly for A TR, the reaction can be re versed (re v erse w ater -gas
shift R WGS). According to the endothermic nature, the con version of the re v erse reaction
is fa v oured at high temperatures. Furthermore, a separation of water from the feed gas
shifts the equilibrium to w ards high CO yield [81].
28

2.3 Methanol Synthesis Process
2.3 Methanol Synthesis Pr ocess
2.3.1 Chemistr y and Thermod ynamics
In methanol synthesis, a syngas consisting mainly of
CO
,
H 2
and
CO 2
is catalytically
con verted to produce ra w methanol. The synthesis comprises a set of exothermic reactions
that proceed selecti v ely to gi ve a single product.
CO
and
CO 2
react with
H 2
according to
the reaction Eqs. 2.25 and 2.26. Additionally the reactants are in v olved in the w ater -gas
shift reaction 2.27, which strongly influences the synthesis by of fering a mechanism for
the intercon version between CO and CO 2 [82].
CO + 2H 2  CH 3 OH ∆ H 298K = − 90 . 64 kJ
mol (2.25)
CO 2 + 3H 2  CH 3 OH + H 2 O ∆ H 298K = − 49 . 67 kJ
mol (2.26)
CO + H 2 O → CO 2 + H 2 ∆ H 298K = − 41 . 47 kJ
mol (2.27)
The con version rate of the reactants is restricted by chemical equilibrium and the kinetic
characteristics of the synthesis. Due to the e xothermic and non-equimolar character of
the reactions with fe wer molecules on the product side than on the reactant side, high
con version to methanol is f a v ored at lo w temperature and high pressure. The synthetic
reactions cause a remarkable v olume reduction. In order to achiev e moderate reaction
rates, the synthesis typically is conducted in the presence of a copper -zinc based catalyst
(CuO/ZnO/
Al 2 O 3
) at temperatures between 200 - 300
◦
C and pressures in the range of
50 - 100 bar , resulting in a yield of 5 - 15 mole-% methanol at the outlet [3, 23, 30, 46,
56]. The commercial catalysts ha v e a high selecti vity to methanol of abov e 99.5% [6].
Additional information on commercially used catalysts is reported in [25, 30].
Dif ferent measures e xist to e v aluate the performance of the synthesis. The carbon
ef ficienc y
C E
measures the incorporation of carbon atoms of the reactant into the final
product. The component con version
X
is the ratio of the dif ference of the amounts of a
reactant at the inlet and outlet to the amount of the reactant at the inlet. The selecti vity
S
is calculated as the ratio of the desired product and the con verted feedstock that is used
to generate the product.
29

Chapter 2 State of Research
Carbon efficiency
C E = ˙ n CH 3 OH
˙ n CO + ˙ n CO 2
(2.28)
Component con version
X m = ˙ n m , in − ˙ n m , out
˙ n m , in for m = CO , CO 2 and H 2 (2.29)
Pr oduct selectivity
S = ˙ n CH 3 OH
˙ n CO + CO 2 , in − ˙ n CO + CO 2 , out
(2.30)
2.3.2 Methanol Reactor T ec hnology and Synthesis
Configuration
The high pressure methanol synthesis de v eloped in the 1920s used a ZnO/
Cr 2 O 3
/
Al 2 O 3
catalyst which fa vored con version at temperatures between 320 - 450
◦
C, and a pressure in
the range of 250 - 350 bar [40, 83]. Adv ances in catalyst technology and the switch from
coal to natural gas as a predominant feedstock led to the de velopment of the lo w-pressure
methanol synthesis during the 1960s. T oday , the worldwide methanol production is based
on the lo w-pressure methanol process, which allo ws a con v ersion of syng as at 25 - 100
bar [84]. Consideration of the chemistry and thermodynamics of the methanol synthesis
led to the de v elopment of a v ariety of processes with different reactor technologies. In
addition to the con ventional g as phase reactor technology , liquid phase reactors ha ve
attracted particular attention due to superior temperature control and high con version
ef ficienc y . Ho we v er , the LPMEOH
TM
technology by the compan y air products only has
been used on an industrial scale in the Eastman Chemical’ s coal gasification comple x [40,
85]. Thus, only gas phase reactors found commercial applications in w orldwide methanol
plants. The gas phase technology can be subdi vided into tw o main cate gories, including
isothermal reactors and adiabatic quench reactors with direct and indirect cooling. Fig.
2.6 lists the processes according to the technology licensor .
The majority of the methanol production processes is based on the ICI-Synetix
∗
and
the Lur gi † reactor technology [30]. ICI-Synetix has a outstanding market share of 60%,
while 27% of the licensed processes are based on Lurgi technology . T o complete the
polypolistic market, Mitsubishi Gas Chemcial Co., Inc. has a share of 8%, KBR, Inc.
∗ T oday Johnson Matthey plc, Great Britain
† T oday Air Liquide Global E&C Solutions, Germany
30

2.3 Methanol Synthesis Process
of 3% and all others 2%. The indirect pseudo-isothermal reactor de veloped by Lur gi
[22, 52, 86] and the direct-cooled adiabatic quench reactor licensed by ICI Synetix /
Johnson Matthe y [87] are in the focus of this work. These corresponding synthesis
configurations mainly dif fer in heat inte gration, make-up g as introduction and reactor
design. An o vervie w of their features is provided in T able 2.1.
• Linde con verter [40, 47, 88]
• Lur gi isothermal reactor [89]
One-step reactor [22, 47, 52, 86]
T wo-step reactor [47]
• Mitsubishi Hea vy Industries
MGC/MHI Supercon verter [40, 90]
• ICI Quench Con verter [87, 91–93]
• Casale Adv anced Reactor [47, 94]
• Haldor T opsoe Collect
Mix ed Distrib ute Reactor [47]
• K ellogg Bro wn and Johnson
Spherical Reactor [40, 47]
• T oyo Engineering Corporation
MRF-Z [40, 95]
Gas Phase Reactor T echnologies
Isothermal r eactors Adiabatic r eactors
Multi-bed quench configuration
with dir ect cooling
Multi-bed quench configuration
with indir ect cooling
Figur e 2.6: Overvie w of commercial methanol reactor technologies.
Common industrial solutions for methanol production are e xclusi v ely based on the
lo w-pressure synthesis. These processes tend to lo w con version rates of the synthesis
gas per pass. Therefore the synthesis reactor in general is used in conjunction with a
rec ycle loop to achie v e adequate yields [40]. T ypically , the loop is equipped with a
pur ge to remo ve inert and hazardous components that w ould otherwise b uild up in the
process and reduce the partial pressure of the acti v e components [71]. For instance, a
methane slip of 2.5% in the reforming unit may result in a loop concentration of 25%.
31

Chapter 2 State of Research
The pur ge gas is either used for comb ustion in a furnace to generate process heat or the
contained hydrogen can be reco v ered by an PSA for syngas conditioning. The purge
ratio is typically in the range of 5-10%, depending on the stoichiometric ratio of the
reactants and the internal heat or hydrogen demand [40]. The pressure lev el results
from a trade of f between ener gy demand for compression of the syngas feed and the
amount of produced methanol. Generally , a high operation pressure fa v ors the production
of methanol due to an increase of the reactants partial pressures. In consequence, the
required amount of the catalyst and the reactor size are decreased which reduces the
capital in vestment. On the other hand, the operation and in vestment cost of the synthesis
gas compression progressi v ely increase. The compression of the make-up gas and the
uncon verted rec ycled syngas can be accomplished in one compressor or independent
compressors [96]. Catalysts are de veloped to allo w for adequate con version rates at a
pressure of e v en belo w 50 bar to further reduce the ener gy intensity and the operation
costs of the syngas compression [40].
T able 2.1: Features of the isothermal reactor technology and the adiabatic multi-bed quench
reactor technology .
Reactor T ype adiabatic quench isothermal
Compan y
ICI / Johnson Matthe y plc [97]
Lur gi / Air Liquide S.A. [98]
T emperature [ ◦ C] 250-300 [91] 230 - 265 [65, 86]
Pressure [bar] 50-100 [87, 91] 50-60 [86]
Catalyst Cu/ZnO/Al 2 O 3 Cu/ZnO/Al 2 O 3
Capacity [mtpd] 1,350 - 3,000 [92, 93] 1,200-5,000 [47]
The commercial reactors are designed for an economically and ef ficient heat remo v al.
[56]. A sensiti v e temperature control is required to allo w an operation within a defined
range. The temperature limits are determined by sufficient acti vity and thermal stability of
the catalyst material. Detailed information on the characteristics of the considered reactor
concepts will be gi v en in the follo wing section. An ov ervie w of their characteristics and
operation parameters is pro vided in T able 2.1. Furthermore, the reactor technologies are
depicted in Fig. 2.7, while the process operation lines are schematically sho wn in Fig.
2.8.
32

2.3 Methanol Synthesis Process
2.3.2.1 Adiabatic Reactor / Multi-Bed Con verter
The adiabatic synthesis concept comprises a number of catalyst beds installed in a series
of reactors or within a common pressure shell. T ypically , the temperature increases
linearly with the con version for an adiabatic operation of an e xothermic reaction [30].
The reaction pathway sho wn in Fig. 2.8 therefore is represented by a typical sawtooth
profile. In general, the reactor and bed sizes are designed to allo w the synthesis reaction
to reach equilibrium. The temperature rise requires a control, which is pro vided by
cooling de vices inbetween the separated catalyst beds[40]. Intermediate or inte grated
cooling concepts are used for damage pre v ention and displacement of the gas temperature
in direction of higher equilibrium con version [99]. T wo adiabatic synthesis concepts
can be distinguished. The first comprises a series of adiabatic reactors with indirect
cooling by application of intermediate heat e xchangers. The second concept refers to a
single shell containing multiple beds with direct cooling through injection of cold syngas
(quench gas). The quench configuration has a maximum single-train capacity of 3,000
mtpd [56, 93]. Thus, in lar ge-scale plants a parallel structure of se v eral quench reactors
is required.
In this work, only the quench technology is taken into consideration for process
inte gration (see Fig. 2.7). In particular , the quench reactor de veloped by ICI Synetix
in the 1960s is in vestig ated. W orldwide, this quench con verter is the most widespread
technology among the lo w-pressure methanol synthesis processes. Compared with other
commercially a v ailable technology , the reactor design is very simple and consequently
features a lo w financial in vestment. The reactor feed is split into se v eral fractions, which
are supplied to the synthesis reactor between four indi vidual catalyst beds by means of
ef fecti ve distrib ution de vices [47, 100]. A successi v e addition of cold quench gas allo ws
suf ficient temperature control. Compared with other designs, the catalyst utilization is
poor since not the entire amount of reactants passes through the total catalyst v olume.
Another disadv antage of the design refers to the fact that heat is not reco v ered at all or
inef fecti vely [43]. The synthesis reactor is typically operated in a temperature range
of 220 - 270
◦
C and a pressure range of 50 - 100 bar using a classic catalyst based on
Cu
/
Zn
/
Al 2 O 3
[93]. Due to irregular flo w distribution and v ariable void fraction along
the bed, the catalyst pellets do not recei v e the same gas flo w . The unev en distrib ution
may cause cold and hot reaction zones, which result in lo w reaction rates and catalyst
deacti v ation [47]. T oday , se veral further de v elopments of the ICI technology e xist for
adv anced quench gas distrib ution [92, 93].
33

Chapter 2 State of Research
Figur e 2.7:
Scheme of the Lur gi isothermal reactor [65] (left) and ICI Synetix multi-bed quench
reactor (right).
2.3.2.2 Isothermal Reactor / Single-Bed Con verter
An isothermal reactor is designed as a shell and tube heat e xchanger with a catalyst on
the tube side and a cooling medium, mostly water , on the shell side. The heat of reaction
is continuously remo ved through indirect heat e xchange and reco v ered as steam on the
shell side. An outer steam drum ensures a stable operation of the system, by controlling
the temperature through an adjustment of the pressure and the flo w rate of the boiling
water . The cooling medium typically has a pressure in the range of 30 - 50 bar to ensure
an operation temperature between 240 - 260
◦
C for high reaction rates. The produced
medium pressure steam is further inte grated into the process and may be used to reboil
the bottom product in the distillation columns. Due to a complex mechanical design, the
isothermal reactor has comparati v ely high in vestment costs and a limited capacity . Some
reactor designs are equipped with an adiabatic top layer to a v oid a preheating of the
synthesis gas at the reactor inlet. The catalytic reaction in the adiabatic layer preheats the
feed to a temperature abo ve the cooling medium temperature to ensure an optimal heat
34

2.3 Methanol Synthesis Process
remo v al. This reduces the length of the tubes as they are used for heat remo v al rather
than for preheating the feed gas. The application allo ws a decrease in size of 10 - 15%,
resulting in an in vestment reduction by 15 - 25% [56].
Fig. 2.7 sho ws the pseudo-isothermal boiling w ater reactor de v eloped by Lur gi GmbH,
which has found wide distrib ution in industry [30, 47]. The reactor consists of a vertical
shell and tube heat exchanger with fix ed tube sheets. The design corresponds to the
general description gi v en abo v e, whereby an adiabatic layer for preheating is missing.
The typical operation conditions are in the range of 50 - 100 bar and 230 - 265
◦
C.
Steam is produced at 40 - 50 bar and can be reused in the compression section or in
the distillation unit. Air Liquide also of fers a Dual-Reactor system (two-stage con verter
system) that is used to process syngas more ef fecti vely for high production capacities.
The unit is a combination of two methanol con v erters. The isothermal water -cooled
reactor is combined in a series with a gas-cooled reactor . In this work, the one-stage
isothermal reactor system is applied on a lar ge-scale, since the flow rates within the
rec ycle loop are in an adequate range.
Figur e 2.8: Process operation lines of an adiabatic reactor and an isothermal reactor .
35

Chapter 2 State of Research
2.4 Crude Methanol Purification
Although the industrial catalysts ha v e a high selecti vity of S > 99.5%, the reactor ef fluent
may contain se v eral impurities [83, 101]. The crude product not only contains water
produced as a by-product in the hydrogenation reaction of
CO 2
, b ut also may include
other impurities, such as dimethyl ether , methyl formiate, acetone, higher chained
hydrocarbons (
C 5
-
C 10
), ethanol and higher alcohols as well as dissolv ed gases (
CH 4
,
CO
,
CO 2
), termed as light ends. The amount of formed by-products depends on the type
and the age of the catalyst and the operation conditions of the synthesis [6, 39]. The
reactions o ver current copper catalysts hardly produce aldeh ydes and ketones, unlike
the catalysts based on zinc-oxide which were used in high pressure processes in the past
[39]. Furthermore, the lo w-pressure synthesis only generates minor quantities of formic
and acetic acid. These chemical components ha ve particularly led to corrosion problems
in the high pressure synthesis.
The objecti v e of methanol purification is to remo v e these impurities to generate a
marketable product. The quality of methanol and the quantities of impurities allo wed
to be present are defined and classified by the ASTM D1152 specification [102]. Three
dif ferent qualities of methanol can be distinguished according T able A.1 in Appendix A
- fuel grade, grade A and grade AA Methanol. Methanol with fuel grade specification
is used as b urner or motor fuel, while grade A methanol in general serv es as a solv ent.
Grade AA methanol is observ ed to become a standard for the methanol industry and is
e xclusi v ely used for further processing to other b ulk chemicals [3, 39, 42, 47].
T oday , raw methanol is e xclusi v ely purified by distillation. According to the amount
of impurities and the desired product quality , dif ferent distillation systems ha v e to be
applied [3, 39, 56]. In single-column systems, the light ends are remov ed o v erhead, while
the product methanol is withdra wn abov e the reflux inlet. The process water and any
hea vy ends are dischar ged from the column bottom. A single-column is only suf ficient
to produce fuel grade methanol.
A two-column system comprises a topping column (e xtraction column) and a refining
column and constitutes a cost-sa ving solution for the purification of crude methanol [6].
In order to reduce the heat duty of the reboiler , the refining column often is split into a first
stage operating at ele v ated pressure (pressure column) and a second atmospheric stage
(atmospheric column). The three-column system therefore represents an ener gy-sa ving
alternati v e. In this configuration the topping column serves for the o v erhead remo v al
of light ends and an y DME present in the crude product, while methanol, water and
hea vy ends are withdra wn from the bottom. The bottom product is subsequently fed
36

2.4 Crude Methanol Purification
to the pressure column to separate the water and the hea vier ends from the methanol
v apor , which is leaving the column at the top. The bottom waste product from the
pressure column may still contain a considerable amount of methanol, which is fed
to the atmospheric column for further reco v ery . Other by products of the atmospheric
column include ethanol and higher alcohols (fusel oil). A schematic flo wsheet of the
two-column system and the three-column systems is sho wn in Fig. 2.9. Depending on
the heat inte gration concept, the columns are heated by lo w or medium pressure steam.
Generally , the methanol vapor of the pressure column is condensed in the reboiler of the
atmospheric column to sa v e ener gy . T ypical features of the distillation columns can be
found in [39, 101].
Figur e 2.9: Flo wsheet of a three-column distillation system.
37

Chapter 3
Methodology
3.1 Thermod ynamic Analysis
This section presents the fundamentals of the ener gy analysis, the con v entional ex er getic
analysis, the economic analysis, the ex er goeconomic analysis and the adv anced ex er getic
analysis. Additionally , Appendix E contains informations in reg ard to the definitions
of the e xer gy of product and fuel as well as the calculation algorithm of the adv anced
e xer getic analysis.
3.1.1 Energetic Anal ysis
According to the first law of thermodynamics ener gy is a conserv ed quantity which can
be di vided in v arious occurring forms. The ener getic state v ariables can be determined
by solving the o v erall ener gy balance in Eq. 3.1. In this work, the ov erall ener gy balance
is simplified by assuming stationary processes and ne glecting dif ferences re garding the
potential and kinetic ener gy . The remaining balance for an open system of the k-th
component includes enthalpy flo w rates
˙
H j
of the inlet (inde x in) and outlet (index out)
streams, electrical or mechanical po wer
˙
W CV
and rate of heat transfer
˙
Q CV
of the control
v olume (inde x cv).
d ( U + K E + PE )
d τ = ˙
Q j , CV + ∑
j = 1
˙
W j , CV + ∑
j
˙ m j , in ( h + ke + pe ) j ,, in
− ∑
j
˙ m j , out ( h + ke + pe ) j , out (3.1)
38

3.1 Thermodynamic Analysis
0 = ˙
Q j , CV + ˙
W j , CV + ∑
j
˙
H j , in − ∑
j
˙
H j , out (3.2)
Dif ferent characteristic parameters are used for a dif ferentiation and e v aluation of the
processes in vestig ated in this work. The chemical ef ficienc y
η ch
is calculated as the
ratio of produced methanol and supplied fuel in Eq.3.3. Both, the energy content of
methanol and fuel is determined by the lo wer heating v alue
H i ∗
. Accordingly , the electric
ef ficienc y
η el
can be defined as the ratio of electric po wer output
˙
W el
and the supplied
fuel in Eq. 3.4. The processes may generate tw o products which are of dif ferent ener getic
quality . In general, electrical po wer is ener getically more v aluable than methanol, so that
the performance of the process cannot be e v aluated solely by using the definition of the
o verall ef ficiency in Eq.3.5.
η ch = ( ˙ m · ˙
H i ) Product
( ˙ m · ˙
H i ) Fuel
(3.3)
η el = ˙
W el,net
( ˙ m · ˙
H i ) Fuel
(3.4)
η tot = η ch + η el = ˙
W el , net + ( ˙ m · ˙
H i ) Product
( ˙ m · ˙
H i ) Fuel
(3.5)
.
The characteristics of turbomachinery components, such as turbines (index T) and
compressors (inde x C), are represented by the isentropic ef ficienc y
η s
. The term
h s , ou t
refers to the e xit specific enthalpy determined by the inlet specific entrop y and the outlet
pressure.
η s , T = ˙
W real
˙
W ideal = h in − h out
h in − h s , out (3.6)
η s , C = ˙
W ideal
˙
W real = h s , out − h in
h out − h in (3.7)
∗ also abbre viated as LH V
39

Chapter 3 Methodology
3.1.2 Con ventional Exer getic Anal ysis
The e xer getic analysis is considered as a con venient and po werful tool to identify and
quantify the inef ficiencies and losses within chemical and ener gy con v ersion processes
from a thermodynamic point of vie w . The methodology re veals the locations, the causes
and the magnitude of the real thermodynamic inef ficiencies. Ex er gy is defined as the
maximum theoretical useful work obtainable as the system is brought into complete
thermodynamic equilibrium with the thermodynamic en vironment while only interacting
with this en vironment [103]. The methodology is well established and has pro ven to
be adv ante gous for furthering the goal of a more ef fecti ve use of resources [104]. The
general e xer gy balance for an open (control-v olume) system is representd by Eq. 3.8.
d E cv
d τ = ∑
j  1 − T 0
T j  · ˙
Q j + ˙
W CV + ∑
j
˙
E j , in − ∑
j
˙
E j , out − ˙
E D (3.8)
The rate of e xer gy change
d E cv
d τ
is calculated as the dif ference between e x er gy transfer
across the boundary which is associated with mechanical or electric po wer
˙
E W
, heat
˙
E Q
, inlet and outlet material streams
˙
E j
, and the rate of e x ergy destruction
˙
E D
. The
temperature
T j
represents the a v erage temperature of the rate of heat transfer at the
boundary of the control v olume, while
T 0
refers to the temperature of the thermodynamic
en vironment. Under steady state conditions, the rate of ex er gy destruction within the
k
-th component
˙
E D , k
is gi v en by Eq. 3.9. The term quantifies the irre v ersibilities within
a component and is caused by mixing, chemical reaction, heat transfer and friction.
˙
E D , k = ˙
E Q + ˙
E W + ∑
j
˙
E j , in − ∑
j
˙
E j , out (3.9)
The e xer gy flo w rate of a stream of matter
˙
E j
is the sum of the physical, chemical,
magnetic, kinetic, and potential ex er gy flo w rate. In this w ork, the contrib ution of the
magnetic, kinetic and potential e xer gies is ne glected.
˙
E j = ˙
E CH
j + ˙
E PH
j (3.10)
˙
E PH
j = ˙ n ·  ( h − h 0 ) − T 0 · ( s − s 0 )  (3.11)
˙
E CH
j = ˙ n ·  ∑
i
x i · e CH
i + RT 0 · ∑
i
x i · ln ( x i )  (3.12)
40

3.1 Thermodynamic Analysis
W ith respect to the calculation of the chemical ex er gy , the model of Szargut is used
as a reference en vironment [105]. The determination of the chemical ex er gy flo w rate,
according to 3.12, is only valid for a mixture of ideal g ases. The chemical e xer gy of
a stream, including a gas and a liquid phase, is calculated by their phase fractions if
condensation occurs at ambient conditions (15
◦
C and 1 bar). The specific physical
and chemical e xer gies are directly calculated by using the simulation sofw are Aspen
Plus
R

[106]. On the base of internal fortran routines, each stream is flashed to ambient
conditions whereas the physical properties are tak en from the simulation database.
An e xer getic analysis often in v olv es the calculation of measures of performance, such
as the e xer gy destruction ratios and the e x er getic ef ficiency . The use of dimensionless
v ariables eases the interpretation of results, in particular when processes with dif ferent
capacities are compared. F or their determination, an e x er gy balance for the o v erall
system and the k-th component can be written in the form of ex er gy fuel / e x er gy product
in Eqs. 3.13 and 3.14. The ex er gy of fuel for the o v erall system
˙
E F , t ot
is equal to the
summation of the e xer gy of product
˙
E P , t ot
, the o verall e x er gy destruction
˙
E D , t ot
and the
e xer gy loss
˙
E L , t ot
. On a component le v el an e x er gy loss is not defined. The SPECO
approach is applied to define the e xer gy of fuel
˙
E F , k
and e xer gy of product
˙
E F , k
for a
component [107].
˙
E F , t ot = ˙
E P , t ot + ˙
E D , t ot + ˙
E L , t ot (3.13)
˙
E F , k = ˙
E P , k + ˙
E D , k (3.14)
The e xer gy destruction ratio
y D , k
in Eq. 3.15 represents the ratio of the ex er gy de-
struction rate of the
k
-th component
˙
E D , k
and the e xer gy rate of the total system fuel
˙
E F , t ot
. Furthermore,
y ∗
D , k
in Eq. 3.16 represents the ratio of the ex er gy destruction rate
of a component and the ex er gy destruction rate of the o v erall system. The ex er getic
ef ficienc y
ε k
defined in Eq. 3.17 constitutes a unique indicator for the performance and
comparability of a system component [108].
y D , k = ˙
E D , k
˙
E F , t ot
(3.15)
y ∗
D , k = ˙
E D , k
˙
E D , t ot
(3.16)
ε k = ˙
E P , k
˙
E F , k
(3.17)
41

Chapter 3 Methodology
The e xer getic analysis re veals information which is in general inaccessible through
application of a thermodynamic analysis using ener gy and mass balances. Thereby , it
pro vides a base to deri v e possible impro v ements, by identifying the major inef ficiencies.
Ho we ver , no information is a v ailable, whether the suggested system modifications indeed
will result in an impro ved o v erall system. Considering each component’ s impro v ement
potential and the interactions among v arious system units in an adv anced ex er getic
analysis will gi v e realistic options for an impro v ement.
3.2 Economic Anal ysis
An economic analysis pro vides information about the economic feasibility of a construc-
tion and its operation while it is also a tool for the estimation of the cost associated
with the generated products of a facility . In this thesis the method of the total re v enue
requirement ( TRR ) is applied. The le v elized v alue
T RR L
consists of the le v elized annual
carrying char ges
C C L
, the operating and maintenance costs
OMC L
and the cost of the
purchase of the fuel F C L .
T RR L = C C L + F C L + OM C L (3.18)
The le v elized carrying char ges represent the annuity of the total capital in vestment
(
T C I
) calculated with the capital reco very f actor (
C RF
). The
T C I
in v olves the fix ed
capital in vestment (
F C I
), the sum of allo w ances for funds during construction time
(
AF U D C
) and start up costs. According to the approach described by Bejan et al. [104],
the fix ed capital in vestment (
F C I
) is calculated as the sum of direct and indirect cost. The
direct costs include the onsite costs and of fsite costs. Onsite costs (also referred to as bare
module cost
BMC
or module cost
C M
) comprise the purchased equipment cost (
PE C
),
cost of installation, piping, instrumentation and control as well as cost for electrical
equipment and materials. The of fsite costs include e xpenditures for ci vil, structural
and architectural work as well as service f acilities. The calculation of the
BMC
is a
challenging task, since a wide v ariety of data relating to cost and capacity of dif ferent
types of components is a v ailable in the literature. The cost of equipment is determined
by the capacity , the materials of construction, the design pressure and temperature. In
the best case v endor quotations should be used for a cost estimation. When vendor cost
quotations are missing, the
PE C
of v arious equipment items often are gi v en in the form
of estimating charts. In this thesis, the
PE C
for heat e xchangers, compressors, expanders,
42

3.2 Economic Analysis
and pumps is based on data from [109–111], whereas the costs for distillation columns
are also obtained from [112, 113]. Cost estimations for the SMR, A TR and DMR unit
were obtained from [109, 110, 113]. For a calculation of the module cost
C M
, ef fects
of material, design temperature, design pressure and equipment features are taken into
account. These effects are considered by f actors (
f m , f T , f p , f d
) which correct the base
cost
C B
(
PE C k
) of a component (modular method). In addition, the bare module f actor
f BM
can be applied to include the cost of an y supporting equipment and connections as
well as an y indirect costs related to the equipment. Estimating charts in general already
illustrate the bare module cost.
C M = C B f d f m f T f p f BM (3.19)
• C M
= module cost of purchased equipment
• C B = base cost of purchased equipment
• f d = design factor
• f m = material factor
• f T = T emperature factor
• f p = Pressure factor
An alternati v e to estimating charts are tables that pro vide component unit cost together
with a scaling e xponent
α
and a bare module factor
f BM
. Although the scaling exponent
for the same equipment can v ary with the size and the reference year , the e xponent
is often assumed to remain constant. Due to the economies of scale ef fect, the v alues
generally are belo w one, sho wing that the percentage increase in capacity is bigger than
the percentage increase in cost [104, 105].
C PE , Y = C PE , W  X Y
X W  α (3.20)
W ith Equation 3.20 the unkno wn cost of a purchased equipment item
C PE , Y
at a
specified capacity
X Y
can be calculated from the kno wn purchase cost of the same
equipment item
C PE , W
at a dif ferent capacity
X W
. The calculated cost of each equipment
item must be brought to a common reference year , which is used as the base for all
cost calculations. This is conducted by applying cost factors which represent inflation
indicators for selected equipment at both, the year of the calculated kno wn costs and the
reference year [105]. In this w ork, the chemical engineering plant cost index (CEPCI)
was used [114].
In contrast to the TCI, annually repeated e xpenditures for operation and maintenance
as well as for fuel are subject to a cost increase. F or their con version to an annuity , the
43

Chapter 3 Methodology
constant escalation le v elization f actor (CELF) is applied, assuming a in v ariable nominal
escalation rate r n and ef fecti ve interest rate i ef f .
C E L F = k ( 1 − k n )
( 1 − k ) C RF with k = 1 + r n
1 + i ef f (3.21)
The operation and maintenance costs at the be ginning of the operation are assumed to
be 5% of the T C I and therefore mainly depend on the capacity of the plant.
OMC L = T C I · 0 . 05 · C E LF gen (3.22)
The fuel costs are calculated from fuel parameters, such as the mass flo w rate, the
heating v alue and the annual operation time
τ
. The le v elization is conducted by applying
a fuel specific le v elization f actor C E LF fuel .
F C L = ˙ m fuel · H i · τ · f c · C E LF fuel (3.23)
Further specifications and assumptions of the economic analysis are summarized in
T able 3.1. The construction is assumed to start in 2013. 40% of the in vestment is made
in the mid-year 2014 and the remaining 60% in mid-year 2015. The plants are operated
with 8000 full load hours a year . The prices for of the fuel natural gas and of the products
methanol and electricity are subject to market fluctuations. Therefore, the profitability of
the plants will be analysed in the conte xt of typical price ranges. Regarding the fuel, the
range is based on the Henry Hub spot price [115], while the price per ton of methanol is
taken from w orld lar gest supplier Methane x
R

, Canada [28]. For an economic assessment
of processes with co-generation, prices for electricity are obtained from [116–118].
Expenditures for utilities such as catalysts and absorption materials are based on [119].
The economic estimation will be conducted using US$ 2016 as a currenc y .
The sum of the le v elized carrying char ges, operation and maintenance cost and fuel cost
results in the annual total re v enue requirement. The economic analysis pro vides a mean
to e v aluate the impact of market prices for fuels and products on the re v enues and the
economic feasibility . Ho we v er , a clear distrib ution of the costs on the products is only
possible through an e xer goeconomic analysis.
44

3.3 Exer goeconomic Analysis
T able 3.1: Specifications of the economic analysis.
V ariable Unit V alue V ariable Unit V alue
A verage capacity
factor
% 91.3 Service facilities % of C M 30.0
CEPCI 2004 [114] - 444.0 Architectural work % of C M 30.0
CEPCI 2010 [114] - 532.9 Contingencies % of C M 15.0
CEPCI 2016 [114] - 541.7 OMC % of TCI 5.0
Ef f. interest rate % 8.0 MSR catalyst $/kg 2.2
Gen. inflation rate % 2.5 SMR catalyst $/kg 10.0
Esc. rate fuel % 1.0 A TR catalyst $/kg 10.0
Esc. rate comp. % 4.0 WGS catalyst $/kg 3.0
Natural gas spot price $ 2016
/GJ
2.9 Pre-reform. cat. $/kg 10.0
3.3 Ex er goeconomic Anal ysis
The e xer goeconomic analysis combines the results of an e x er getic- and an economic
analysis and pro vides information which is not obtainable through the con ventional
analyses b ut essential for the design of a cost-ef fecti v e system. The purpose of an
e xer goeconomic analysis is to identify the relati ve cost importance of each component.
This is achie v ed by application of the e x er gy-costing principle, which requires costing
equations for all components [105]. These equations consider cost rates
˙
C j
of entering
and e xiting ex er gy streams by assigning a specific monetary v alue
c j
to each stream as
well as cost rates due to capital in vestment
˙
Z C I
k
and operation and maintenance e xpenses
˙
Z OM
k
[104]. T o obtain more precise information on the comple x chemical systems, the
cost rate of a stream is further split into the cost of the physical e x er gy
˙
C PH
j = c PH
j ˙
E PH
j
and the cost of the chemical e xer gy ˙
C CH
j = c CH
j ˙
E CH
j .
n
∑
j = 1
( c j · ˙
E j ) k , in + ( C C L + OM C L ) · PE C k
∑
n
PE C · τ
| {z }
˙
Z k = ˙
Z CI
k + ˙
Z OM
k
=
m
∑
j = 1
( c j · ˙
E j ) k , out (3.24)
The v alues of
˙
Z C I
k
and
˙
Z OM
k
are calculated by apportioning the
C C L
and the
OMC L
among the system components according to the contrib ution of the
k
-th component to the
45

Chapter 3 Methodology
purchased equipment cost for the o verall system. Heat losses ov er system component’ s
surface could also e xhibit a cost stream, b ut are ne glected in this w ork. If an e x er getic
fuel
˙
E F , k
and product
˙
E P , k
can be defined for a component, the cost balance can also be
noted as in Eq. 3.25.
c P , k ˙
E P , k = c F , k ˙
E F , k + ˙
Z k (3.25)
The terms
c F , k
and
c P , k
denote the a v erage cost per e xer gy unit of fuel and product
while
˙
Z k
refers to the cost rate associated with the in vestment and the operation and
maintenance e xpenses. F or dissipati v e components, such as condensers and v alves, an
e xer gy product is not defined. Howe ver , since these components are indispensible for
the process operation, their costs, associated with the in v estment, the operation and
maintenance and the e xer gy destruction (termed as
˙
C Dif f
) will be char ged to the o v erall
systems’ products.
A system of linear equations is obtained by setting up the component balances. In case
the system is under -determined (the number of unkno wn cost streams is lar ger than the
number of balances) auxiliary equations are required to assign a v alue to the unkno wn
v ariables. In general, the inlet streams of a component are assumed to be kno wn, while an
information about the outlet streams is missing. If the number of exiting streams is higher
than one (
m > 1
) a formulation of m-1 auxiliary equations is necessary to determine
specific a v erage cost of all streams. The F-principle (Fuel rule) on the fuel side and
the P-principle (Product rule) are used to determine the auxiliary equations [105]. On
one hand, the F-rule states that the specific cost associated with e xer gy remo v ed from a
component is equal to specific cost at which the e x er gy was supplied to the same stream
in the upstream components. On the other hand, the P-rule states that each unit of ex er gy
is supplied to all product streams at the same a v erage cost c P , k .
After the balance equations are solv ed, an ex er goeconomic e v aluation can be conducted
by considering dif ferent e x er goeconomic v ariables. One important outcome refers to the
cost rate that is associated with the e xer gy destruction ˙
C D , k of the k -th component.
˙
C D , k = c F , k · ˙
E D , k (3.26)
The e xer goeconomic e v aluation of a design re v eals the components with the highest
cost impact which includes the cost rate associated with the e x er gy destruction, as
calculated in Eq. 3.26, and the in vestment cost of the respecti v e component
˙
Z k
. At
first, the components are ranked in descending order according their cost significance
46

3.4 Adv anced Exer getic Analysis
˙
C D , k + ˙
Z k
. By application of the e x er goeconomic factor
f k
in Eq. 3.27, the contrib ution of
the in vestment cost (non-e x er gy related cost) to the cost rate
˙
C D , k + ˙
Z k
can be identified.
Thus, the factor provides an information about the internal cost distribution within a
component.
f k = ˙
Z k
˙
Z k + ˙
C D , k
(3.27)
The second important v ariable used in an e x er goeconomic e v aluation is the relativ e
cost dif ference, denoted as
r k
. The v ariable expresses the relati ve cost increase in the
a v erage cost per ex er gy unit between the product
c P , k
and the fuel
c F , k
of a component
[104]. As sho wn in Eq. 3.28, the relati v e increase depends on the cost rate associated
with the e xer gy destruction and the in vestment ˙
C D , k + ˙
Z k .
r k = c P , k − c F , k
c F , k
= ˙
C D , k + ˙
Z k
c F , k · ˙
E P , k
(3.28)
The e xer goeconomic factor
f k
depends on the operation of a component and therefore
has an indi vidual v alue for each unit. Howe ver , for classes of components, such as
heat e xchangers, turbines and compressors, there are some common v alue ranges that
usually apply . Consequently , an assessment by the factor must take into consideration
the component type. Depending on the v alue of the ex er goeconomic factor , a trade-of f
between in vestment cost and e x er gy destruction has to be made. High
f
v alues suggest a
reduction in the in vestment costs of a component at the expense of its e x er getic ef ficienc y .
On the other hand, lo w v alues indicate cost-sa vings that might be achie v ed for the o v erall
system by a decrease of the irre v ersibilities (impro ving the e x er getic ef ficienc y) e ven if
the component’ s capital cost will increase.
The results of the e xer goeconomic e v aluation are used to determine changes to the
initial design in order to impro ve it from both, an economic and a thermodynamic point
of vie w . The compromises between the in vestment cost and the cost rate of the e xer gy
destruction can gi v e indications for an iterati ve design impro v ement of the in vestig ated
system.
3.4 Ad v anced Ex er getic Analysis
The con ventional e x er getic analysis indentifies the location, the magnitude and the cause
of irre v ersibilities and indicates a general strate gy for a thermodynamic impro v ement.
47

Chapter 3 Methodology
Ho we ver , the method is not able to identify the interdependcies among the plant compo-
nents or to estimate the real potential for improv ement. The adv anced e x er getic analysis
concept [120] pro vides detailed information on the interactions between each system
component and on their real improv ement potential. T o comply with this, the e x er gy
destruction within each component is split into its endogenous and exogenous parts
as well as a v oidable and una v oidable parts, respecti v ely . A combination of these two
splitting approaches pro vides further v aluable information for an impro v ement of the
o verall ef ficiency . A schematic of all options for splitting the ex er gy destruction of a
component is presented in Fig. 3.1, which was adopted from [121].
Figur e 3.1: Overvie w of splitting options of the ex er gy destruction within a component in an
adv anced e x er gy analsyis (adopted from [121]).
Splitting the e x er gy destruction of the
k
-th component into its a v oidable (inde x A V) and
una v oidable (inde x UN) parts allo ws for an e v aluation of the real improv ement potential
of a component within a fixed process arrangement (Eq. 3.29). The main drawback
of this splitting approach is subjecti vity: the calculation of una v oidable v ariables is
associated with the estimation of v alues based on the best kno wledge and e xperience of
48

3.4 Adv anced Exer getic Analysis
an engineer . The unav oidable e x er gy destruction of a component
˙
E D , k
can be directly
calculated from a simulation in Aspen and is weighted by the default e x er gy rate of the
product (Eq. 3.30).
˙
E D, k = ˙
E A V
D, k + ˙
E UN
D, k (3.29)
˙
E UN
D, k = ˙
E P , k · ( ˙
E D, k / ˙
E P , k ) UN (3.30)
The una v oidable part of the e x er gy destruction refers to the amount of e xer gy de-
struction that cannot be further reduced because of technoeconomic limitations, such
as manufacturing methods, cost of materials, and a v ailability . The remaining a v oidable
e xer gy destruction represents the sa vings in irrev ersibilities of the
k
-th component that
can be achie v ed through a technically feasible design or operational impro v ement.
The thermodynamic interdependencies among the components can be e v aluated by
splitting the e xer gy destruction of the
k
-th component into the endogenous and exoge nous
parts.
˙
E D, k = ˙
E EN
D, k + ˙
E EX
D, k (3.31)
The endogenous e xer gy destruction
˙
E EN
D, k
is associated with the irre v ersibilities oc-
curring within the
k
-th component operating with default e x er getic ef ficienc y while
all other components are assumed to be re v ersible [108]. In contrast, the exogenous
e xer gy destruction
˙
E EX
D, k
within component
k
is caused by the irre v ersibilities of other
system components. The endogenous e xer gy destruction can be calculated from a set of
equations without the need for additional simulations. Se v eral approaches for calculating
the endogenous e xer gy destruction ha v e been suggested [122–124]. These approaches
face theoretical shortcomings and computational problems for chemical reaction
systems [122].
The combination of the two splitting approaches results in four parts of the e x er gy
destruction which are included in the final results of an adv anced e x er getic analysis.
Only the endogenous a v oidable and e xogenous a v oidable e x er gy destruction should be
considered to find promising modifications for an impro v ement of the ov erall system
[123]. The calculation algorithm for obtaining the una v oidable endogenous e xer gy
destruction ˙
E UN,EN
D, k is gi v en in Appendix F.
˙
E D, k = ˙
E UN,EN
D, k + ˙
E UN,EX
D, k + ˙
E A V ,EN
D, k + ˙
E A V ,EX
D , k (3.32)
49

Chapter 3 Methodology
˙
E UN,EX
D, k = ˙
E UN
D, k − ˙
E UN,EN
D, k (3.33)
˙
E A V ,EN
D, k = ˙
E EN
D, k − ˙
E UN,EN
D, k (3.34)
˙
E A V ,EX
D, k = ˙
E A V
D, k − ˙
E A V ,EN
D, k (3.35)
The remaining parts of the e xer gy destruction of the
k
-th component in Eq. 3.32 can
be calculated by using equations 3.33 to 3.35. Based on simultaneous consideration of
the four parts of e xer gy destruction within the k -th component, measures for impro ving
the component and the o verall system can be deri ved [120].
3.5 Sim ulation and Software
The simulations were conducted using the software Aspen Plus
R

(Aspen) v ersion 9.1
[106]. Furthermore, for data analysis, data management, and additional calculations,
Matlab
R

has been used. The analyses ha v e been carried out at steady-state conditions.
In Aspen, the property method Redlich-Kwong-Soa v e (RKS) w as used to model the
gas path. W ith respect to the acid gas remov al system, the property method based on
the Perturbed Chain Statistical Association Fluid Theory (PC-SAFT) equation of state
was applied to the glycol Dimeth yl Ether of Polyethylene Glycol (DEPG) representing
the solv ent in the Acid gas remo v al (Sele xol
R

). This method is adequate to model
liquid systems with lar ge molecules for single components or in mixtures. Re garding
the material properties, Aspen uses information from the NIST databank [125]. The
properties of water and steam were calculated based on the steam table formulation
IAPWS’95 [126].
In Aspen, the simulations ha v e been performed in sequential modular mode. Internal
routines named design sepcification are used to determine the conditions of se v eral
outlet streams. Furthermore, the embedded sensitivity function w as used to determine
fa v ourable operation parameters for the key units of the process (see Section 4.1). The
internal calculator module is used for a direct calculation of characteristic parameters,
such as the stoichiometric measures and feed ratios. In particular , the module helps to
monitor the operation of the sensiti v e chemical components. Another internal module
called tr ansfer can be useful for the calculations of the a v oidable and una v oidable
e xer gy destruction as well as for the simulations of loops which are induced by rec ycle
streams. F or all kinetically limited reactions, kinetics are implemented in the internal
module r eactions . This can be a challenging task, as the gi v en reactions kinetics often
50

3.5 Simulation and Software
not suit the form that is required by Aspen. The program pro vides dif ferent types of
kinetics in order to describe the reaction mechanism in the most suitable manner . Of
particular importance are Langmuir -Hishle w ood-Houston-W atson kinetics (LHHW) [82]
and Po wer -La w kinetics. It may be challenging to implement kinetics from the literature
in the form required by Aspen. A detailed description on the used kinetics and their
implementation into Aspen is pro vided in Appendix B.
On a component le v el, se v eral reactors are simulated by using the RPLUG reactor
model, which is suitable to simulate a plug flow reactor . The model allo ws for an
implementation of kinetics for an heterogenous catalysis and assumes a perfect mixture
in the radial direction and no mixture in axial direction. A v ariety of kinetic models is
implemented into the reactors to describe the occurring reactions properly . Power -Law
kinetics based on a study by Luyben et al. [96] are implemented into the SMR and DMR
model. The same kinetic type is used to model the WGS by using kinetic parameters
from Choi and Stenger [127]. Re garding the methanol synthesis, the reaction kinetic
inputs to the RPLUG model are taken from a study by v an den Bussche et al. [128] and a
study by Graaf et al. [129, 130] , in which a
Cu
/
ZnO
/
Al 2 O 3
and a
Cu
/
Zn
/
Al
/
Zr
catalyst
is used for con version, respecti vely . Finally , after conducting the sensiti vity analyses, the
model of v an den Bussche w as chosen for implementation.
An y comb ustion process was simulated by using the RGIBBS reactor model which
calculates the outlet stream at chemical equilibrium conditions by minimizing the Gibbs
free ener gy . W ith respect to pre-reformer and the desulphurizer , a fixed con version rate
is implemented into a RST OICH model. The distillation columns are simulated by using
the RADFRA C model.
51

Chapter 4
System Design and Modelling
The synthesis of a process design for a chemical plant is an important task. The decisions
made in a design process significantly influence the future profitability , the en vironmental
impact, fle xibility in operation, and changes in system design. Therefore, in the first part
of this chapter , parameter studies are conducted for the major units to generate ne w and
adv anced process designs for high con version and ef ficienc y . In the second part, six final
process designs will be introduced which subsequently will be subject to the analyses.
4.1 P arameter Stud y f or the Synthesis of Pr ocess
Designs
The reforming unit and the methanol synthesis reactor are decisi v e for the o v erall
performance of a methanol plant, since the y are highly integrated from a chemical
and thermal point of vie w . Due to a v ariety of impact parameters in conjunction with
a comple x reaction mechanism, the setting of fa v ourable operation conditions is a
challenging task. Sensiti vity analyses are conducted for these units to sho w the impact
of uncertain operation parameters. In conclusion, beneficial operation parameters for
high con version and methanol yield are identified taking into account e xisting technical
limitations. Furthermore, information about the need for additional process steps can be
deri v ed from the results.
4.1.1 Ref orming of Methane
The reforming unit is the ke y component within a natural gas-based methanol pro-
duction process. The choice of a specific reforming technology influences the rate of
con version, the product yield, the ov erall ef ficiency of the process and consequently
52

4.1 Parameter Study for the Synthesis of Process Designs
also the economies of the plant. Therefore, profound in vestig ations of the dif ferent
syngas generation technologies are indispensable. In the follo wing, the influence of
operation parameters, such as the S/C ratio, O/C ratio,
CO 2
/C ratio, temperature and
pressure on the con version and the syng as composition will be in vestig ated. The results
of the sensiti vity analyses are based on the assumption, that the reforming reactions
reach chemical equilibirum. For all in vestigations, the natural gas feed has the same
composition which is gi v en in Appendix A.
4.1.1.1 Steam Reforming
W ithin the sensiti vity analysis of the steam reformer , the influence of the S/C-ratio and
the reactor temperature and pressure on the composition and the methane con version is
in vestig ated.
Figur e 4.1:
Sensiti vity analysis of the steam methane reforming technology: a) and b) sho w the
S module and the
CH 4
-con version at 5 bar and c) and d) refer to the v alues at 30 bar .
53

Chapter 4 System Design and Modelling
Figure 4.1 sho ws the results of the stoichiometric S module and the
CH 4
-con version
at 5 and 30 bar , respecti vely . In general, when selecting the operation conditions of a
reformer , a priority should be gi v en to the con version rate, since lar ge methane slips are
technologically not feasible for the synthesis unit. Considering the
CH 4
-con version in
Fig 4.1 b) and d), the contourplots suggest an operation at a temperature of at least abo v e
900
◦
C and a S/C ratio abo ve 2 for a high con version rate. In this operation range the S
module is fix ed to a v alue of 3.
A comparison of a) with c) and b) with d) sho ws that the pressure has a less significant
influence on the composition and the
CH 4
-con version. W ith increasing pressure at
constant temperature, hydrogen tends to react with
CO 2
to produce
CO
according to
the R WGS, resulting in lo wer v alues of the S module. On the other hand, the methane
con version decreases with increasing pressure due to the non-equimolar character of the
SMR reaction. Thus, at ele v ated pressure a higher operation temperature is required to
achie v e a lo w methane slip. In the context of the o v erall process design the pressure
should not be selected too lo w in order to limit the ener gy demand of the syngas
compression unit.
4.1.1.2 A utothermal Ref orming
The sensiti vity of syngas composition and the
CH 4
-con version to the supply of the
reforming agents
H 2 O
and
O 2
as well as to the operation pressure is in vestig ated. The
results of the analysis are sho wn for a pressure of 10 bar in Figs. 4.2 a) and b) and for a
pressure of 40 bar in Figs. 4.2 c) and d), respecti v ely . F or an O/C ratio of 0, pure SMR is
presented, while PO X is sho wn for an S/C ratio of 0. Based on the v ertical lines, it can
be concluded that comb ustion reactions in general has a higher impact than SMR on the
composition and the con version.
Also for the autothermal reformer , the operating conditions for a stoichiometric optimum
and maximum con version are dif ferent. As for the other reforming technologies, a high
con version rate has priority when selecting the operation parameters. Consequently , the
reformer should be operated with an O/C-ratio of 0.6, which is currently the technological
minimum [22, 66, 131]. F or a high con version of
CH 4
, the syngas composition can be
adjusted via the O/C ratio. Accordingly , an increase in the O/C ratio will result in a
syngas with reduced hydrogen content. Even if the Figs. 4.2 a) and c) could lead to the
conclusion that the influence of the S/C-ratio is lo w , the composition strongly depends
on the amount of inte grated steam. A closer look at the composition re v eals, that a S/C
ratio of 0.5 - 0.7 (at an O/C ratio of 0.6) generates a syngas with fa vourable composition.
54

4.1 Parameter Study for the Synthesis of Process Designs
Figur e 4.2:
Sensiti vity analysis of the autothermal reforming technology: a) and b) refer to the S
module and the CH 4 -con version at 10 bar and c) and d) to the v alues at 40 bar .
4.1.1.3 Dr y and Combined Ref orming
The sensiti vity of the S module and the CH 4 -con version to the S/C ratio and the CO 2 /C
ratio as well as to the operation temperature is presented for an operation pressure of 5
bar and 30 bar in Fig. 4.2. The contourplots a) and b) refer to the S module, while c) and
d) sho w the con version rate at the respecti v e pressure. The plots in each Figure refer to
an e xtremum in terms of the
CO 2
/C ratio - the solid lines represent the results for a lo w
ratio of 0.5, while the dashed lines sho w the results for a high ratio of 2.5. The results of
a pure operation as a DMR are sho wn on for a S/C ratio of zero.
F or high
CO 2
/C ratio, 99%
CH 4
-con version is achie ved for an operation temperature
in the range of 900 and 1050
◦
C. The course of the graphs in Figs. 4.2 b) and d) sho ws,
that reforming by
CO 2
is preferred at a high temperature le v el while the impact of the
SMR dominates between 300 - 600
◦
C. Considering the syngas composition in a) and
55

Chapter 4 System Design and Modelling
c), the lo w S module sho ws that for high
CO 2
/C ratio a lar ge amount of carbon dioxide
lea v es the reformer uncon verted. For a lo w ratio syng as with a satisfactory composition
in the range of S equals 1.6 - 1.7 can be produced. A comparison of 4.2 a) with c) and b)
with d) sho ws that a pressure increase at constant operation temperature increases the
con version rate and the v alue of the S module.
Figur e 4.3: Sensiti vity analysis of the dry methane reforming technology: a) and b) refer to the
S-Module and the CH 4 -con version at 5 bar and c) and d) to the v alues at 30 bar .
In conclusion, the DMR should be operated in a temperature range of 900 - 1100
◦
C
at a lo w pressure of 5 bar and a
CO 2
/C ratio of one. F or an operation of a combined
reformer (CMR), a slightly lo wer
CO 2
/C ratio of 0.8 and a S/C ratio between 0.6 and 0.7
is recommended.
56

4.1 Parameter Study for the Synthesis of Process Designs
4.1.1.4 Comparative Assessment of the Syngas Characteristics
The sensiti vity analyses for the single reforming technologies ha v e sho wn that the
operation parameters for a stoichiometric optimum and a high con version of
CH 4
are
dif ferent. Consequently the single reforming technologies are not capable of pro viding a
syngas with an optimal composition for the methanol synthesis. Ho wev er , as mentioned
earlier in Section 3.1, the S module is not fully reliable for an assessment of the syngas
quality . Therefore, the syngas composition shall also be assessed by using the other
stoichiometric parameters.
024 6 8 10 12
PO X
A TR
DMR
SMR
CO/CO 2 ratio
S module
H 2 /CO ratio
Figur e 4.4: Range of the syngas composition for the commercially applied methane reforming
technologies. The composition is expressed by three common stoichiomeric
measures.
The bar chart in Fig. 4.4 sho ws the range of the common stoichiometric measures for
the basic reforming technologies. The data has been recorded for the typical operation
parameters which ha v e been described in Section 3.1. Furthermore, the results are based
on a high
CH 4
-con version abo v e 95%. The de viation of the stoichiometric measures
from their ideal v alues means losses in terms of methanol yield and causes a voidable
costs in the synthesis unit. The syngas from PO X, A TR and DMR is highly reactiv e due
to high
CO
/
CO 2
ratio, b ut is deficient in hydrogen. In contrast, the syngas from SMR is
characterized by e xcess hydrogen and comparati v ely lo w reactivity .
57

Chapter 4 System Design and Modelling
Although the characterization of the syngas indicates trends in re gard to the ef ficienc y
of the synthesis reactor , no concrete conclusions can be drawn. Therefore, the output data
of the sensiti vity analyses for the reforming Section is used as an input information to an
equilibrium synthesis reactor model with fix ed temperature and pressure. The ternary
plot in Fig. 4.5 sho ws the sensiti vity of the methanol yield to the inlet composition of the
synthesis gas at equilibrium conditions of 250
◦
C and 50 bar . Since the synthesis is also
kinetically restricted by the catalyst, the methanol yield is much lo wer for the rate based
models (see Fig. 4.6).
Figur e 4.5: T ernary diagram for the methanol yield at equilibrium conditions of 50 bar and
250 ◦ C for a gi v en inlet composition.
The selected equilibrium conditions are typical for the industrial used lo w-pressure
methanol synthesis process. The ellipses indicate the agglomeration of recorded syngas
compositions, that are obtained from the basic reforming technologies under the typical
operation conditions. The ke y refers to the operation pressure of the respecti ve reformer
technology . A high methanol yield may be obtained by feeding a syngas generated by
PO X. Ho we v er , this technology requires a large air separation unit for oxygen supply
and has disadv antages in terms of heat inte gration [132]. The industrially most common
SMR technology yields relati v ely small amounts of methanol due to hydrogen e xcess. In
58

4.1 Parameter Study for the Synthesis of Process Designs
re gard to a syngas generated by A TR, the methanol yield can v ary widely depending on
the reformer operation. The methanol yield obtained for a typical syngas feed from DMR
is in a moderate range. Ho we ver , the analysis sho ws that DMR has a potential to be a
competiti v e alternati v e to common reforming technologies. The results obtained from
the sensiti vity analysis suggest, that a combination of dif ferent reforming technologies
can potentially impro ve the con version and therefore also the ef ficienc y of the o v erall
system.
4.1.2 Methanol Synthesis
The synthesis reactor is the second core unit of a methanol process as it has high
impact on the design and the performance of other system units. Do wnstream of the
reactor , the number of distillation columns and flashing units, as well as the reboiler and
condenser duties are determined by the con version and the formation of by-products.
Furthermore, the implementation of syngas conditioning technologies is determined by
the performance of the reactor .
The pre vious analyses ha v e sho wn that particularly the feed composition influences
the con version rate and thereby the product yield. As described in Section 2.3.2, the
con version is limited by the kinetics of the catalyst and the chemical equilibrium. So
far , only chemical equilibrium has been considered as a restriction in the synthesis. T o
obtain more precise information on the synthesis, kinetics of the reactions hav e to be
taken into account. A variety of models is a vailable to describe the kinetics in methanol
synthesis [47]. In this w ork, the two well-kno wn kinetic models of Graaf et al. [129,
130] and V anden Bussche et al. [128] hav e been selected for implementation. The first
model describes the kinetics o ver a
Cu
/
Zn
/
Al
/
Zr
catalyst for a pressure range of 1 - 50
bar and a maximum temperature of 250
◦
C. The model of V anden Bussche et al. uses a
commercially a v ailable
Cu
/
ZnO
/
Al 2 O 3
catalyst that is applied in se v eral modern plants.
Both models are implemented into the same reactor model ( RPLUG -Model in Aspen) for
analysis of the methanol yield. The underlying operation conditions are v ery similar to
those of the equilibrium model. The use of two models also serv es as a v alidation of the
models and as a control of their correct implementation in Aspen. The results of the two
models sho w a v ery good agreement. Due to the widespread use of the catalyst, V anden
Bussche’ s model finally is used for the simulation of the processes.
The sensiti vity of the methanol fraction to the inlet composition is depicted in the
ternary composition diagrams in Fig. 4.6. In comparison with Fig. 4.5, the kinetic
limitations become ob vious. Furthermore, dif ferent conclusions can be dra wn in re gard
59

Chapter 4 System Design and Modelling
to the optimal syngas composition. As reported in [6, 46, 57, 133, 134] a maximum
con version results for a lo w
CO 2
content of 2 - 10 % and a
H 2
/
CO
ratio of 2. Thus, for
the design of ne w and ef ficient processes, syngas conditioning units ha v e to be considered
to decrease the
CO 2
content while adjusting the
H 2
/
CO
ratio to a v alue of tw o. It should
be noted that this sensiti vity analysis identifies trends to mak e design decisions. The
actual methanol yields of the processes may dif fer because the reaction rates depend on
the components partial pressures and the temperature, and thus on the circulating mass
flo w , the reactor geometry and the heat flux.
Since methanol is synthesized from a mixture comprising primarily
CO
,
CO 2
and
H 2
, it is of v aluable kno wledge which components acti vely participate in the synthesis.
Therefore, the sensiti vity of the component con version rates to the feed composition
was analysed for both models at approximately 250
◦
C and 50 bar . The obtained results
which are sho wn in Appendix correspond to the findings of other studies [128–130, 134,
135].
Figur e 4.6: Sensiti vity of the methanol yield to the inlet composition of the syngas for the
kinetic models of Froment and Graaf under pseudo-isothermal conditions.
60

4.1 Parameter Study for the Synthesis of Process Designs
In the last step, one of the presented reactor models introduced in Section 2.3.2
had to be chosen for a system inte gration. All processes were simulated using both
reactor models, including the synthesis loop for recirculation of uncon verted syngas.
T ypical industrial design features and operation parameters were implemented into the
models. W ith respect to the four-bed quench reactor system, the distribution of the
quench gas w as subject to an optimization to maximize the product yield at the reactor
outlet [136]. The composition of the supplied syngas was tak en from the simulation
of the reformer technologies. Even if the characteristics of the synthesis ha v e to be
considered indi vidually depending on the synthesis g as, the methanol yields correspond
to the findings of other studies [86, 87, 137].
SMR A TR DMR SMR-A TR SMR-DMR
260 280 300 320
0
2
4
6
8
10
T emperature [ ◦ C]
mole-%
260 280 300 320
0
1
2
3
4
T emperature [ ◦ C]
mole-%
Figur e 4.7: Methanol yield from the isothermal reactor and the adibatic quench reactor for
dif ferent processes.
The results of the methanol yield for the isothermal reactor and the adiabatic quench
reactor are sho wn in Fig D.3. Obviously , for all test cases, a lar ger amount of methanol is
obtained for the isothermal reactor (left in Fig. D.3). The in-situ cooling minimizes the
temperature rise inherent to the reaction, therefore allo wing an adv antegous operation
characteristic (see Fig. 2.8). Thus, the isothermal reactor was chosen for all processes,
which are subject to further in vestig ations.
61

Chapter 4 System Design and Modelling
4.2 Overvie w of Pr ocesses and Subsystems
The thermodynamic and economic assessment is conducted for the processes in T able
4.1, which are further described in the follo wing sections. Six different process designs
are asssigned to the two cate gories Base Pr ocesses and Advanced Pr ocesses . The base
processes are characterised by a single reforming technology - the SMR, A TR and DMR
design use a stand-alone steam reformer , an autothermal reformer and a dry reformer ,
respecti v ely . Except for the SMR case, no syngas conditioning steps are integrated into
these designs.
The adv anced process designs are the result of a carefully conducted design synthesis,
which is based on detailed sensiti vity analyses and heuristic optimization approaches,
that were pre viously introduced in Section 4.1. In addition to a more comple x design
of the reforming unit, syngas conditioning steps may be inte grated into synthesis path.
In the CMR process, combined reforming by hydrogen and carbon dioxide is applied
of the syngas stoichiometry . A
CO 2
absorption based on the Sele xol
R

process is
implemented for further syngas adjustment. The SMR-A TR process includes a serial
two-step reforming concept with do wnstream syngas conditioning by a re v erse water
gas shift. The SMR-DMR process features a parallel configuration of a steam methane
reformer and a dry methane reformer .
T able 4.1: Specifications of the analysed processes.
Base Pr ocesses SMR A TR DMR
Reformer SMR A TR DMR
Syngas conditioning CO 2 -hydrogenation none none
Acid gas remo v al none none none
Synthesis reactor isothermal isothermal isothermal
Carbon utilization yes no yes
Advanced Pr ocesses CMR SMR-A TR SMR-DMR
Reformer CMR SMR and A TR SMR and DMR
Syngas conditioning WGS R WGS none
Acid gas remo v al Sele xol R
 none none
Synthesis reactor isothermal isothermal isothermal
Carbon utilization yes no yes
62

4.3 Basic Assumptions
In general, a methanol plant consists of the follo wing subsystems: a pretreatment unit
for the remo v al of sulfur impurities from natural gas, a reforming unit serving for the
production of synthesis gas in conjunction with a syng as conditioning, a synthesis unit
for the con version of syngas to methanol, a purification section for refining of crude
methanol in a set of distillation columns and a steam c ycle for the heat inte gration of these
subsystems (detailed information is pro vided in Chapter 2). The heat inte gration man-
agement system is designed for maximum electricity generation and is highly integrated
into the other process units. Depending on the configuration, the steam c ycle features a
dif ferent design, whereby heat is reco v ered as steam in a tw o-pressure le v el HRSG. In
general, steam must be pro vided to se v eral process units, including the hydrogenation
unit, the pre-reformer and in some cases to the reforming unit (SMR, CMR) and the
syngas conditioning units (WGS). The amount of comb ustion fuel required for co v ering
the heat duty mainly depends on the selection of the reforming unit. Further details will
be gi v en within the description of the configurations.
4.3 Basic Assumptions
Some of the major assumptions and design specifications used for the simulation of each
process are presented in T able 4.2. The SMR, A TR, DMR, CMR process are designed
for a medium capacity , while SMR-A TR and the SMR-DMR generate methanol on a
lar ge scale. In all process designs, a high calorific natural gas with a molar composition
of of 94.9%
CH 4
, 2.5%
C 2 H 6
, 0.2%
C 3 H 8
, 0.06%
C 4 H 10
, 0.02%
C 5 H 12
1.6%
N 2
, 0.7%
CO 2
, 0.02%
COS
[138] is used as a feedstock. The lo wer heating v alue
H i
and the higher
heating v alue
H s
of the natural gas is 48.78 MJ/kg and 54.10 MJ/kg, respecti v ely . Further
assumptions for particular processes are gi v en in the follo wing sections.
63

Chapter 4 System Design and Modelling
T able 4.2: Basic assumptions for all processes.
Subsystem / Component Unit V alue
General
ambient temperature [139] ◦ C 15.0
ambient pressure [139] bar 1.013
ambient air fractions mole-% 21 / 79
mechanical ef ficienc y of turbomachinery % 99.0
electric generator ef ficienc y % 99.0
electric motor ef ficienc y % 96.0
syngas compressor isentropic ef ficienc y % 85.0
rec ycle compressor isentropic ef ficienc y % 87.0
Pr etr eatment
H 2 S absorption coef ficient [113] % 99.9
con version rate of COS [113] % 99.5
Methanol Synthesis
catalyst density [96] kg/m 3 1,775
v oid fraction [96] - 0.5
diameter of tubes [96] m 0.04
heat transfer coef ficient [96] kW/m 2 K 0.3
Purification
topping column
condenser pressure [101] bar 1.3
reboiler pressure [101] bar 1.6
pr essur e column
condenser pressure [101] bar 1.9
reboiler pressure [101] bar 2.8
Steam Cycle
steam turbine isen. ef f. LP/MP/HP [140, 141]
% 90/93/92
pumps isentropic ef f. % 85.0
condenser pressure bar 0.05 - 0.4
max. li v e steam temperature ◦ C 620.0
pinch point temperature dif ference liq-
uid/liquid,liquid/gas,gas/g as
◦ C 5,10,15
pressure drop liquid/gas per 100 ◦ C % 2.0,3.0
pressure loss e v aporation % 4.0
64

4.4 SMR Process
4.4 SMR Pr ocess
Steam methane reforming is the most widely used technology for the industrial production
of methanol. The first Base process (SMR process) therefore comprises a configuration
which is based on commercial a v ailable processes for a capacity of 2,500 mtpd. The data
was mainly obtained from [50, 96, 142]. A simplified flo w diagram is presented in Fig.
4.8 and the corresponding data for a selection of flo ws from the simulation is gi ven in
T able 4.6. Further detailed information on the chemical composition of each stream is
pro vided in T able D.1. The general assumptions in T able 4.2 are completed by the design
specifications pro vided in T able 4.3.
Natural gas is entering the system with stream 1 and is first compressed to a pressure
of 30 bar by CM-01. After desulphurization in R-01 and R-02, the sweet gas is entering
an adiabatic pre-reformer , where higher hydrocarbons are cracked into methane. The
product gas e xits the reactor with a temperature of 510
◦
C and a pressure of 27 bar .
Do wnstream, the ef fluent (stream 4) is mix ed with steam at a S/C-ratio of 3.5 before
entering a re generati v e heat e xchanger for preheating. The plant is equipped with a
tub ular fired steam reformer , where the heat (250 MW) for the endothermic steam
reforming reaction is pro vided by the comb ustion of 28.7 kg/s natural gas (stream 7) and
16.5 kg/s tail gas. The tubular reformer consists of 3,500 tubes, each ha ving a diameter
of 0.01 meter and a length of 12 meter . The design specifications based on [142] were
slightly adjusted to the required capacity . A large flo w of combustion air (stream 8) is
compressed in steam-turbine dri v en compressors to a pressure of 3.5 bar . The syngas
lea v es the steam reformer with a temperature of 950
◦
C and a pressure of 20.5 bar . The
reaction parameters of the SMR section were selected according to the results of the
sensiti vity analyses in Section 4.1.
The air mass flo w and the required fuel supply are determined by the high temperature
heat demand of the SMR unit. Do wnstream, the lo w temperature heat needs to be
inte grated to increase the ef ficienc y of the process. The combustion gase s (stream 5)
lea v e the furnace with a temperature of 884
◦
C and are fed to a heat recov ery steam
generator (HRSG). A part of the heat (approximately 30%) is inte grated into the process
by preheating the syngas in E-01 and by supply of steam to R-02, R-03 and R-04. The
remaining heat (70%) is reco vered as steam on two pressure le v els. The lo w pressure
le v el of 30 bar was selected for steam supply to the components within the syngas track.
The high pressure le v el of 70 bar w as the result of an iterati v e optimization process for
maximum po wer generation of the steam turbines T -01 and T -02.
65

Chapter 4 System Design and Modelling
T able 4.3: Design specifications of the SMR process.
Subsystem / Component Unit V alue
Ref orming unit
pr e-r eformer
reactor pressure [50] bar 28.0
reactor temperature [50] ◦ C 510.0
steam r eformer
catalyst density [142] kg/m 3 2,000
v oid fraction [142] - 0.5
number of tubes - 3,500
length of tubes m 12.0
diameter of tubes m 0.01
heat transfer coef ficient kW/m 2 K 0.5
S/C-ratio [33] mole/mole 3.5
Methanol synthesis
number of tubes - 10,000
length of tubes [143] m 18.0
pur ge ratio [96] % 5.0
Purification
topping column
number of stages [101, 143] - 42
reflux ratio - 0.6
distillate to feed ratio - 0.3
pr essur e column
number of stages [101] - 81
reflux ratio - 0.8
distillate to feed ratio - 0.9
atmospheric column
number of stages - 56
reflux ratio - 0.9
distillate to feed ratio - 0.7
The synthesis gas lea ving the reforming unit is compressed to 50 bar in multiple syng as
compressor (CM-04 and CM-05) with intermediate cooling and water deduction. The
compressed gas contains a surplus of hydrogen, which is represented by a stoichiometric
module of S equal 3. Before entering the isothermal synthesis reactor , the fresh syngas
(stream 6) is blended with three rec ycle streams of uncon verted syng as. A special
feature of this process is the stoichiometric adjustment of the syngas composition by
66

4.4 SMR Process
inte gration of a pure
CO 2
stream for carbon utilization. The integration of
CO 2
for direct
hydrogenation in the synthesis unit w as discussed in section 2.2.1. The increase of the
carbon dioxide fraction in the syngas feed causes an increased methanol yield. Ho we ver ,
an e xcessi v e amount of water is also produced via the re verse water -gas shift reaction,
increasing the size and the ener gy demand of the distillation unit.
T able 4.4: Simulation results for selected flo ws of the SMR process.
Flo w T ype T ◦ C p [bar] ˙ m [ kg/s ] ˙
E [ MW ]
1 Natural gas 15.0 10.0 16.9 838.5
2 Natural gas 124.6 29.0 16.9 841.6
3 Steam 550.0 29.5 1.3 2.0
4 Clean gas 509.5 24.6 18.2 852.4
5 Exhaust gas 883.7 1.6 1631.9 996.6
6 Syngas 148.2 52.5 42.9 1012.6
7 Natural gas 15.0 16.0 28.7 1425.8
8 Air 15.0 1.0 1586.8 9.4
9 CO 2 45.0 52.5 13.4 8.8
10 Syngas 250.0 49.6 391.8 7099.4
11 Syn. + CH 3 OH 38.0 42.9 391.8 6979.8
12 Syngas 38.6 42.9 313.2 5558.5
13 Syngas 35.5 3.5 16.4 284.5
14 Crude methanol 38.0 2.0 60.5 1124.0
15 AAA methanol 71.0 1.3 30.0 672.3
16 Steam 440.0 58.7 358.0 501.8
17 Steam 300.0 27.1 81.1 93.5
18 Condensate 60.1 0.2 439.1 160.2
19 W ater 15.0 1.0 65.4 3.26
20 Of fgas 146.3 1.2 1630.5 107.8
Before entering the synthesis reactor , the syngas is preheated to 250
◦
C by recuperation
of the thermal ener gy of the synthesis product. The use of a preheater is indispensable,
since a minimum temperature of 220
◦
C is required for the acti v ation of the reaction
mechanism in the lo w-pressure synthesis. Due to moderate con version rates and lar ge
rec ycle flo w rates, the size of the reactor is relati v ely lar ge. The reactor vessel contains
10,000 tubes, ha ving a length of 18 meter and a radius of 0.04 meter each [143]. On the
shell side, the heat of the exothermic reactions is reco v ered as steam ha ving a constant
67

Chapter 4 System Design and Modelling
Figur e 4.8: Process flo wsheet of the methanol plant with steam reforming.
68

4.5 A TR Process
temperature of 262
◦
C and a pressure of 50 bar . The heat is returned to the process in the
reboilers of the distillations columns. The ef fluent from the reactor (stream 11) consists
of crude methanol, water and uncon verted synthesis gas. The v apor phase and the liquid
phase are separated in two flash drums D-01 and D-02. Three centrifugal compressors
(CM-06, CM-07 and CM-08) are used to rec ycle uncon verted syngas to the reactor inlet.
Due to a pressure drop of 7 bar within the synthesis loop, the energy demand of the
compressors is high. In order to a v oid a b uild up of inert components in the cycle, 5 %
of the mole flo w rate of the fresh inlet gas is remo v ed from the synthesis. The tail gas
(stream 13) is used as a b urner fuel in the SMR unit to reduce the consumption of natural
gas.
Containing 73 mole-% methanol, the crude product still contains a considerable
amount of water and other impurities, which are remo v ed in a three-column distillation
unit. T ypical operation parameters were taken from [101, 143] and adjusted for high
methanol purity . In the first column C-01 the light ends,
H 2
,
CH 4
,
CO
and
N 2
, are
remo ved from the crude product and rec ycled to the reactor inlet. The bottom product is
further purified in a refining and an atmospheric column (C-02 and C-03), which mainly
serv e to separate the lar ge amount of water from the product.
4.5 A TR Pr ocess
The process based on autothermal reforming in general features a less comple x design,
b ut requires an air separation unit (ASU) as an additional subsystem. In comparison
to the SMR process, the heat integration system is dif ferent, since the reforming unit
represents a heat source instead of a heat sink. As a consequence, a combustion of the
rec ycled tail gas is suf ficient to co v er the heat and electricity demand of the process.
The specific assumptions for the simulation of the A TR process are gi v en in T able 4.5.
A flo wsheet of the design is presented in Fig. 4.9 and the corresponding data for a
selection of streams is gi v en in T able 4.6. Further detailed information on the chemical
composition of each stream is pro vided in T able D.2.
A flo w of 22.8 kg/s natural gas is compressed to a pressure of 40 bar before desul-
phurization. Based on the sensiti vity analyses in Section 4.1, a high operation pressure
is required for high con version in the comb ustion and catalytic zone of the A TR (R-03
and R-04). Despite of a lo w S/C ratio of 0.6, the flo w rate of the syng as exiting from the
reforming unit is higher than for the SMR process. Ho we v er , due to elev ated reforming
pressure the electricity demand in the syngas compression unit is kept in a moderate
69

Chapter 4 System Design and Modelling
range. The con version in the A TR is completed by supplying oxygen to the combustion
zone at a O/C ratio of 0.55. As an additional subsystem an air separation unit (ASU) is
required to pro vide oxygen with a purity of 95 mole-% (Stream 5).
T able 4.5: Design specifications of the A TR process.
Subsystem / Component Unit V alue
Ref orming unit
pr e-r eformer
reactor pressure bar 34.8
reactor temperature ◦ C 550.0
autothermal r eformer
catalyst density kg/m 3 2,500
v oid fraction - 0.5
length m 7.0
diameter [144] m 1.0
O/C ratio [63] mole/mole 0.55
S/C ratio [63] mole/mole 0.6
Methanol synthesis
number of tubes - 10,000
length of tubes m 12.0
pur ge ratio % 11.0
Purification
topping column
reflux ratio - 0.6
distillate to feed mole ratio - 0.1
r efining column
reflux ratio - 0.75
distillate to feed ratio - 0.95
atmospheric column
reflux ratio - 0.8
distillate to feed ratio - 0.95
The dischar ge from the reformer is first cooled in E-02 and E-03 before it is compressed
in CM-04 and CM-05 to the operation pressure of the methanol synthesis. Liquid w ater is
rejected from the system after the intercooler (E-05) to reduce the b urden on do wnstream
components. Due to a high con v ersion rate in the synthesis reactor (R-03) and a pur ge
70

4.5 A TR Process
ratio of 11%, the total circulating mass flo w is only half as big as for SMR process. This
is beneficial in re gard to the size and cost for the reactor , the recuperator E-10 and the
rec ycle compressors (CM-06, CM-07, CM-08). For instance, the length of the reactor
tubes is only two-thirds of those in the SMR process. Apart from that, the design and
operation of synthesis reactor is similar for the SMR and the A TR process.
Despite the moderate concentration of water in the crude product (8 mole-%), a three
column distillation unit is required for the production of grade AA methanol. In compari-
son to the SMR process ha ving a water content of 25% in the crude product, the heat
duty of the reboiler heat e xchangers is lo w .
T able 4.6: Simulation results for selected flo ws of the A TR process.
Flo w T ype T [ ◦ C ] p [bar] ˙ m [ kg/s ] ˙
E [ MW ]
1 Natural gas 15.0 10.0 22.8 1131.9
2 Natural gas 155.3 39.8 22.8 1137.7
3 Clean gas 558.8 34.8 36.7 1171.8
4 Syngas 922.63 33.3 60.2 1096.4
5 Oxygen 277.0 34.8 23.6 10.5
6 Syngas 157.0 25.2 60.2 1018.7
7 Air 15.0 1.0 418.3 2.5
8 Syngas 250.0 50.7 207.2 3763.3
9 Syn. + CH 3 OH 38.0 44.8 207.2 3690.5
10 Pur ge gas 38.2 4.0 18.9 319.2
11 Crude methanol 38.0 2.0 33.8 711.4
12 Bottom product 133.0 2.8 0.9 0.1
13 AAA methanol 70.9 1.3 28.4 637.7
14 W ater 15.0 1.0 13.9 0.7
15 Steam 600.0 175.3 61.3 104.9
16 Steam 218.3 15.0 130.2 131.2
17 Condensate 76.0 0.4 130.2 54.7
18 Of fgas 110.0 1.9 437.2 35.1
19 Air 15.0 1.0 101.6 0.6
20 Nitrogen 17.9 3.3 76.9 9.7
21 Exhaust Gas 817.1 2.5 437.2 251.8
71

Chapter 4 System Design and Modelling
Figur e 4.9: Process flo wsheet of the methanol plant with autothermal reforming.
72

4.6 DMR Process
Process heat is mainly provided by cooling of the product stream from R-03 and by
comb ustion of the pur ge gas from the synthesis unit in R-09. In contrast to processes
with endothermic reforming, the capacity of the comb ustion unit is determined by the
total electricity demand instead of the process heat demand. W ithin the HRSG heat is
reco vered on a medium pressure le vel (80 bar) and a high pressure le v el (200 bar) for
electricity generation in a set of steam turbines (T -01 - T -03). Process steam with a low
pressure between 35 - 40 bar is used for preheating the feed to the A TR and for supply
to the components within the syngas track (R-01, R-02, R-03). The low-pressure ASU
using an intercooled multistage air compressor does not fa v or a heat integration.
4.6 DMR Pr ocess
The process design with dry methane reforming (DMR) is similar to the design of the
SMR process. The major design specifications and assumptions are gi v en in T able 4.7. A
process flo wsheet is presented in Fig. 4.10 and the corresponding data for a selection of
streams is gi v en in T able 4.8. Further detailed information on the chemical composition
of the streams is pro vided in T able D.3.
In contrast to the pre viously described processes, DMR is operated at a lo w pressure
to attain a high con version and a f a v ourable syng as composition. Thus, the design is not
equipped with a natural gas compressor . The dry reformer (R-04) is based on a tub ular
GHR design where con vecti ve heat is supplied by a flue gas from a comb ustion unit
(R-05). W ith 360 MW , the process heat demand of the DMR is very lar ge, since the
endothermic reactions are carried out between 900 - 1000
◦
C. T o decrease the demand
for natural gas (stream 7), a lar ge amount of tail gas is rec ycled from the synthesis unit
(stream 12). The proportion in the total amount of comb ustion fuel is considerably higher
than for the SMR process. A stream of 19.9 kg/s of natural gas is mix ed with 60 kg/s of
tail gas and comb usted with air (stream 8). Before entering the tubes of the DMR, a pure
CO 2
flo w (stream 5) of 65.6 kg/s is mix ed with desulphurized natural g as (stream 4) at a
CO 2
/C ratio of 1. The process parameters are selected to a v oid carbon deposition on the
catalyst surface.
The syngas from the reforming unit (stream 6) is characterized by a
H 2
/
CO
ratio of
1 and contains less than 3 mole-% of
CO 2
. Due to high pressure diff erence between
the reforming and the synthesis unit, the electricity demand for the syngas compression
(CM-04 and CM-05) is higher than in the alternati v e base processes. Furthemore, the
processing of a lar ge syngas flo w additionally increases the ener gy consumption.
73

Chapter 4 System Design and Modelling
T able 4.7: Design specifications of the DMR process.
Subsystem / Component Unit V alue
Ref orming unit
pr e-r eformer
reactor pressure bar 3.2
reactor temperature ◦ C 550.0
dry-r eformer
catalyst density kg/m 3 2,000
v oid fraction - 0.5
number of tubes - 7,500
length of tubes m 18.0
diameter of tubes m 0.01
heat transfer coef ficient kW/m 2 K 0.5
CO 2 /C ratio mole/mole 1.0
Methanol synthesis
number of tubes - 10,000
length of tubes m 18.0
pur ge ratio % 10.0
Purification
topping column
reflux ratio - 0.6
distillate to feed ratio - 0.1
pr essur e column
reflux ratio - 0.6
distillate to feed ratio - 0.95
The stoichiometrically inappropriate composition of the untreated syngas results in
a relati v ely lo w methanol yield. As a consequence, the circulated mass flo w rate of
uncon verted syng as is twice as lar ge design and three times lar ger than for the A TR
design. The large mass flo w rate leads to increased gas velocities, which in turn decrease
the con version rate. The components in the synthesis section accordingly are designed
for a lar ge capacity and therefore require a high capital in vestment. In comparison to the
other base processes, the crude product sho ws a high methanol purity of 97.2 mole-%.
Due to the lo w
CO 2
content of the fresh syngas, the formation of excessi v e w ater in
the synthesis unit is suppressed. The processing of the crude product to grade AAA
74

4.6 DMR Process
methanol can be conducted in a two-stage column system with comparati v ely lo w heat
demand.
As in the SMR design, the capacity of the comb ustion unit is also determined by the
process heat demand of the reforming unit. The comb ustion gases lea ving the shell of
the GHR still ha v e a temperature of 950
◦
C (stream 15). On a lo wer temperature lev el,
the e xcess heat is recov ered on tw o pressure le v els for electricity generation in a set of
steam turbines (T -03, T -02 and T -01). The li v e steam parameters are gi ven in T able 4.8
for stream 16, 17 and 18. A single reheating is used after T -02 to increase the steam
content to 87 % at a condenser pressure of 0.05 bar .
T able 4.8: Simulation results for selected flo ws of the DMR process.
Flo w T ype T [ ◦ C ] p [bar] ˙ m [ kg/s ] ˙
E [ MW ]
1 Natural gas 15.0 7.0 25.3 1256.4
2 Natural gas 155.2 6.5 25.3 1257.7
3 Steam 565.0 67.9 90.1 144.1
4 Clean gas 546.7 5.5 33.2 1281.4
5 CO 2 15.0 5.5 65.6 35.6
6 Syngas 282.8 52.0 93.7 1557.5
7 Natural gas 15.0 5.0 19.9 1036.0
8 Air 15.0 1.0 1442.5 8.6
9 Syngas 250.0 51.7 632.9 8339.6
10 Syn. + CH 3 OH 38.0 45.4 632.9 8231.6
11 Syngas 38.0 45.4 533.9 6647.2
12 Pur ge gas 29.9 4.0 59.3 723.4
13 Crude methanol 38.0 2.0 36.7 802.8
14 AAA methanol 80.4 1.9 31.5 706.3
15 Exhaust Gas 930.8 2.2 1521.8 1045.0
16 HP steam 600.0 178.6 306.3 524.3
17 MP Steam 414.8 64.5 388.5 532.6
18 LP steam 200.0 3.9 388.5 327.6
19 Condensate 32.9 0.05 388.5 68.0
20 W ater 15.0 1.0 7.9 0.4
21 Of fgas 120.0 1.6 1521.8 135.6
75

Chapter 4 System Design and Modelling
Figur e 4.10: Process flo wsheet of the methanol plant with dry reforming.
76

4.7 CMR Process
4.7 CMR Pr ocess
The bi-reforming concept (combined reforming) was introduced in Section 2.1.9 as an
alternati v e for dry reforming with stoichiometrically more f a v ourable syngas composition.
The design is based on the specifications gi v en in T able 4.9 and has methanol capacity
of 2,800 mtpd. The process flo wsheet and the corresponding data for a selection of
streams is presented in Fig. 4.11 and T able 4.10, respecti v ely . Further details on the
chemical composition of the flo ws is pro vided in T able D.4 in Appendix D. Moreo ver ,
the temperature profiles of heat transfer are presented in Fig. 4.12.
Similar to the DMR design, the CMR process dispenses with the need for a natural
gas compressor due to lo w operating pressure or the reformer . A steam flo w of 90.1
kg/s (stream 3) and a
CO 2
flo w of 65.6 kg/s (stream 5) is supplied at a S/C and O/C
ratio of 0.68 and 0.8, respecti v ely . Due to the supply of steam the reaction mechanism is
less endothermic than pure DMR, b ut still requires a considerable amount of heat. In
comparison to the DMR process, 100 MW less heat is required for a con version with
lo w methane slip. Accordingly , the fuel demand for natural gas (stream 20) and tail gas
(stream 10) is lo w . As in the other processes with endothermic reforming technology ,
the e xcess heat of the comb ustion gas (stream 15) is reco vered as steam for electricity
generation. The heat reco v ery steam generator comprises two pressure le vels. The lo w
pressure le v el mainly serv es for steam supply to the reforming unit and for preheating
of the reactants, while the high pressure le v el w as chosen for maximum electricity
generation. The flue gas (stream 21) lea v es the process with a temperature of 120 ◦ C.
The syngas e xiting the reformer (stream 5) is characterized by a
H 2
/
CO
ratio of 1.4
and contains 4 mole-%
CO 2
. Downstream, the gas is processed in a lo w temperature
water -g as shift (R-04) to adjust the
H 2
/
CO
ratio to the stoichiometric requirement (see
Section 2.2.2). Simultaneously , the
CO 2
content increases to more then 10%, which
would result in an unacceptable lo w con v ersion and lar ge rec ycle flo ws in the synthesis
unit. Therefore, after cooling in E-06 and compression to 31 bar in CM-01, the syngas is
fed to an A GR for
CO 2
capture by absorption through Sele xol
R

solv ent. A stream of 20
kg/s
CO 2
is rejected from the unit and is rec ycled to the reforming unit. The absorption
leads to a reduction of the mass flo w and therefore reduces the size and the in vestment
cost of do wnstream process equipment.
77

Chapter 4 System Design and Modelling
T able 4.9: Design specifications of the CMR process.
Subsystem / Component Unit V alue
Ref orming unit
pr e-r eformer
S/C ratio - 1.3
reactor pressure bar 3.2
reactor temperature ◦ C 550.0
bi-r eformer
S/C ratio - 0.68
CO 2 /C ratio - 0.8
Syngas conditioning unit
water -gas shift
gas temperature at inlet ◦ C 300.0
length m 6.0
diameter m 2.0
CO 2 -absorption
gas temperature at inlet ◦ C 31
lean solv ent temperature [145] ◦ C -1
solv ent/gas mole ratio CO 2 absorber [145] - 1.05
CO 2 compressor isen. stage ef ficienc y [146] % 77-78
Methanol synthesis
number of tubes - 10,000
length of tubes m 18.0
pur ge ratio % 10.0
Purification
topping column
reflux ratio - 0.6
distillate to feed ratio - 0.1
pr essur e column
reflux ratio - 0.9
distillate to feed ratio - 0.95
78

4.7 CMR Process
The syngas lea ving the A GR consists of 32.2 mole-% CO, 0.3 mole-% CO 2 and 66.7
mole-%
H 2
. The composition fa v ors a high con v ersion and therefore allo ws for very
small rec ycle flo ws in the synthesis loop (stream 8), reducing the size of the recycle
compressors and the heat e xchangers. The high methanol yield is also due to the generous
geometrical design of the synthesis reactor . The crude product entering the distillation
unit (stream 11) sho ws a high purity of 99.4 mole-%, since the formation of water is
pre v ented by the lo w
CO 2
content of the syngas. Howe ver , a two-stage column system
is required to obtain methanol with grade AA quality . Thus the heat requirement of the
distillation columns (C-01 and C-02) is comparati v ely lo w . In total, 32.5 kg/s methanol
are generated.
T able 4.10: Simulation results for selected flo ws of the CMR process.
Flo w T ype T [ ◦ C ] p [bar] ˙ m [ kg/s ] ˙
E [ MW ]
1 Natural gas 15.0 5.0 16.9 836.8
2 Natural gas 155.0 3.5 16.9 837.0
3 CO 2 550.0 8.0 35.1 28.6
4 Clean gas 800.0 2.9 65.8 909.8
5 Syngas 950.0 2.0 65.8 1088.5
6 Syngas 80.0 31.0 64.5 1033.5
7 Syngas 60.0 24.0 41.5 990.5
8 Syngas 248.3 50.5 112.2 2698.5
9 Syn. + CH 3 OH 38 45.5 112.2 2620.3
10 Pur ge gas 37.8 4.0 7.4 173.2
11 Crude methanol 38.0 2.0 37.9 846.3
12 AAA methanol 80.4 1.9 32.5 735.1
13 W ater 15.0 1.0 13.8 0.7
14 Steam 600.0 41.6 166.6 267.7
15 Condensate 32.9 0.05 206.9 37.5
16 W ater 20.0 1.0 549.0 27.6
17 Of fgas 120.0 1.6 651.2 85.6
18 Exhaust gas 958.0 3.2 651.2 493.1
19 CO 2 183.0 15.0 20.0 40.9
20 Natural gas 15.0 4.0 14.9 735.9
21 Air 15.0 1.0 628.9 3.7
79

Chapter 4 System Design and Modelling
Figur e 4.11: Process flo wsheet of the methanol plant with combined reforming.
80

4.8 SMR-A TR Process
Figur e 4.12: T emperature profiles of heat transfer within the CMR process.
4.8 SMR-A TR Pr ocess
The SMR-A TR process features a two-step reforming concept, comprising a serial
configuration of a con vecti ve heated steam methane reformer (primary reformer) and an
autothermal reformer (secondary reformer). The reforming concept has been applied on
a commercial scale and is designed for a lar ge capacity abo v e 5,000 mtpd. A flo wsheet of
the system design is pro vided in Fig. 4.14 and the corresponding stream data is gi v en in
T able 4.12. The stream data is supplemented by T able D.5 in Appendix D. Furthermore,
the temperature profiles of the heat transfer are presented in 4.13.
A flo w of 43.9 kg/s natural gas is compressed in CM-01 to the operating pressure of the
reformers (R-03 - R-05). After desulfurization and pre-reforming, methane is partially
con verted to syng as by SMR in a temperature range of 730 - 920
◦
C for an operating pres-
sure of 38 bar (7% of methane feed). The remaining methane is subsequently comb usted
in the second autothermal reformer (R-04) at an O/C ratio of 0.51 to adjust the syngas
composition for approaching the stoichiometric optimum. The inte gration of the sec-
ondary reformer reduces the load and the size of the primary reformer , ha ving a lo w S/C
ratio of only 0.5. The heat of the ef fluent (stream 6) from the secondary reformer is reco v-
ered as steam, which is provided to the shell side of the SMR to co v er the heat demand of
the endothermic reaction. In total, 70 MW of heat are transferred to the tubes of the SMR.
81

Chapter 4 System Design and Modelling
T able 4.11: Design specifications of the SMR-A TR process.
Subsystem / Component Unit V alue
Ref orming unit
pr e-r eformer
reactor pressure bar 38.6
reactor temperature ◦ C 680.0
steam r eformer
catalyst density kg/m 3 2,000
v oid fraction - 0.5
number of tubes - 5,000
length of tubes m 11.0
diameter of tubes m 0.01
heat transfer coef ficient kW/m 2 K 0.5
S/C ratio mole/mole 0.5
autothermal r eformer
catalyst density kg/m 3 2,500
v oid fraction - 0.5
length m 7.0
diameter m 3.0
O/C ratio mole/mole 0.51
W ater -gas shift
inlet temperature ◦ C 1,040
length m 14.0
diameter m 4.0
Methanol synthesis
number of tubes - 10,000
length of tubes m 18.0
pur ge ratio % 10.0
Purification
topping column
reflux ratio - 0.6
distillate to feed ratio - 0.1
pr essur e column
reflux ratio - 0.6
distillate to feed ratio - 0.95
atmospheric column
reflux ratio - 0.8
distillate to feed ratio - 0.6
82

4.8 SMR-A TR Process
The operation of an autothermal reformer requires the application of an ASU as an
additional and cost-intensi v e unit. A relativ ely lar ge oxygen stream of 41.8 kg/s is
supplied to the comb ustion zone of the autothermal reformer (R-04). As in the A TR
process, the lo w-pressure ASU does not f a v or heat inte gration.
At the outlet of the reforming unit the syngas composition is characterized by a small
hydrogen e xcess (
H 2
/
CO
ratio is 2.1). By water rejection in drum D-01 the composition
is shifted to the product side of reaction 2.24 to dri v e the re v erse w ater -gas shift reaction
(R WGS) in R-06. Accordingly , a lo w amount of hydrogen reacts with carbon dioxide
to increase the carbon monoxide content. The final syngas composition contains 0.8
mole-% CO 2 , 63.6 mole-% H 2 and 31.0 mole-% CO.
Due to high pressure drop along the gas path, the syngas compression also represents
a lar ge ener gy consumer in this process. The pressure dif ference at the inlet and outlet
of the syngas compression almost amounts to 40 bar . After synthesis is carried out at
50 bar and approximately 250
◦
C, the product stream from R-08 contains 10.4 mole-%
methanol, which is separated from the gaseous phase in a flashing unit (D-03 and D-04).
The corresponding rec ycle to feed ratio has a lo w v alue of 2.3. The crude product (stream
13) is characterized by a high purity of 97 mole-% methanol. Ho we v er , three distillation
columns are required to produce 63 kg/s methanol of the desired product quality .
Figur e 4.13: T emperature profiles of heat transfer within the SMR-A TR process.
83

Chapter 4 System Design and Modelling
Figur e 4.14: Process flo wsheet of the methanol plant with two-step reforming.
84

4.8 SMR-A TR Process
The units of the entire process chain are embedded in a heat inte gration system. A
moderate amount of additional natural gas (stream 9, 11.8 kg/s) is mix ed with a pur ge of
25.2 kg/s of rec ycled syngas (stream 12) and is subsequently comb usted with air . The air
is e xposed to a slight pressure ele v ation by CM-02. The heat of the combustion gases is
reco vered as steam in a tw o-pressure le v el HRSG for electricity generation, in order to
co ver the demand of the internal consumers. Only by using a single reheating, the steam
quality at the oulet of the LP turbine is kept abo v e a technological minimum feasible
le v el of
x = 87 . 1%
. The flue gas lea v es the HRSG with stream 21 and has a temperature
of 100 ◦ C.
T able 4.12: Simulation results for selected flo ws of the SMR-A TR process.
Flo w T ype T [ ◦ C ] p [bar] ˙ m [ kg/s ] ˙
E [ MW ]
1 Natural gas 15.0 10.0 43.9 80.0
2 Natural gas 144.2 44.8 43.9 2191.4
3 Clean gas 734.2 37.8 71.4 2287.4
4 Steam 1219.1 48.5 61.3 169.9
5 Oxygen 273.0 34.0 41.8 18.6
6 Syngas 1059.1 31 113.2 2205.0
7 Syngas 150.0 11.2 93.0 1996.9
8 Air 176.4 4.0 750.1 115.5
9 Natural gas 15.0 4.0 11.8 585.4
10 Syngas 250 50.6 323.0 6700.5
11 Syn.+ CH 3 OH 38.0 44.3 323.0 6535.9
12 Pur ge gas 36.9 4.0 25.2 485.0
13 Crude methanol 38.0 2.0 71.1 1565.3
14 Bottom product 100.8 3.0 2.7 47.9
15 AAA methanol 71.0 1.3 63.0 1423.9
16 W ater 15.0 1.0 27.5 1.4
17 Steam 600.0 167.4 261.2 446.9
18 Steam 420.0 45.0 325.6 440.5
19 Condensate 45.8 0.1 325.6 82.1
20 W ater 22.0 60.8 1062.9 59.9
21 Of fgas 100 1.9 787.1 87.6
22 Exhaust gas 1170.8 2.4 787.1 750.7
23 Nitrogen 17.8 3.3 133.3 16.8
24 Air 15.0 1.0 175.1 1.0
85

Chapter 4 System Design and Modelling
4.9 SMR-DMR Pr ocess
The last process design comprises a parallel arrangement of a tub ular gas-heated SMR
and a tub ular gas-heated DMR for a lar ge capacity abo v e 5,000 mtpd. A process flo wsheet
is presented in Fig. 4.16 and the corresponding data for selected streams is gi ven in
T able 4.14. Further details on the chemical composition of each stream are pro vided
in T able D.6 in Appendix D. The underlying assumptions and design specifications are
pro vided in T able 4.13. F or a better understanding of the heat integration management,
the temperature profiles of heat transfer are sho wn in Fig. 4.15.
A lar ge mass flo w rate of 45.6 kg/s natural gas is entering the system with stream 1
and is preheated and desulphurized in E-01 and R-01. The reforming unit is operated
on two pressure le vels, since the DMR (R-03) and SMR (R-04) f a v or a high con version
at significant dif ferent operating pressure. Only the subset of gas intended for the SMR
(stream 3) is compressed in CM-01 while the gas grid pressure at the inlet is suitable
for an ef ficient operation of the DMR. Do wnstream of R-03 the process g as requires a
compression in CM-02 to allo w a mixing with the syng as from the SMR. One quater of
the total mole flo w rate of natural g as is mixed with carbon dioxide (stream 5) at a
CO 2
/C
ratio of 2 before entering the DMR. The con version is carried out in a temperature range
of 700 - 1050
◦
C and a pressure of 9 bar . The remaining natural gas (75 mole-%) is fed
to the SMR unit and blended with steam at a S/C ratio of 1. The reactions are accom-
plished in a temperature range of 900 - 1100
◦
C and a pressure of 30 bar . The selected
split fraction is the result of an iterati v e optimization approach for determination of a
fa v ourable syngas composition. Howe ver , after mixing the resulting product gas (stream
13) has a H 2 /CO ratio of 1.8 and therefore is characterized by hydrogen deficienc y .
The process gas ha ving a pressure of 28 bar is subsequently dried and compressed in
CM-05 and CM-06. The non-stoichiometric composition of the feed gas only allo ws
a moderate con version within the synthesis reactor resulting in relati vely lar ge mass
flo ws in the rec ycle loop. Consequently , the rec ycle compressors and the product coolers
are designed for a lar ger capacity resulting in a high in vestment cost. It is worth to
mention, that the synthesis unit is operated with a high pur ge ratio of 20% to reduce the
size of the loop components. Furthermore, a lar ge recycle of syng as (stream 16) lo wers
the demand for comb ustion fuel (stream 12) to co ver the internal ener gy consumption.
Ho we ver , the process heat demand of both reformers still requires the supply of 27.9 kg/s
natural gas to the comb ustion unit. Lar ge air compressors CM-03 and CM-04 operat-
ing with lo w pressure ratio are used to o v ercome the pressure drop within the flue g as line.
86

4.9 SMR-DMR Process
T able 4.13: Design specifications of the SMR-DMR process.
Subsystem / Component Unit V alue
Ref orming unit
pr e-r eformer
reactor pressure bar 9.2
reactor temperature ◦ C 540.0
Steam r eformer
catalyst density kg/ m 3 1,775
v oid fraction - 0.5
number of tubes - 1,600
length of tubes m 15.0
diameter of tubes - 0.1
heat transfer coef ficient kW/m 2 K 0.5
S/C-ratio mole/mole 1.0
Dry methane r eformer
catalyst density kg/m 3 2,000
v oid fraction - 0.5
number of tubes - 550
length of tubes m 13.6
diameter of tubes m 0.1
heat transfer coef ficient kW/m 2 K 0.5
Methanol synthesis
number of tubes - 10,000
length of tubes m 18.0
pur ge ratio % 20.0
Purification
topping column
reflux ratio - 0.6
distillate to feed ratio - 0.1
r efining column
reflux ratio - 0. 75
distillate to feed ratio - 0.95
The crude product lea ving the synthesis unit with stream 17 has a high methanol
concentration of 96.3 mole-%. The desired product quali ty can be ensured by the
87

Chapter 4 System Design and Modelling
application of a two-stage column system. A flow of 65.9 kg/s methanol with Grade AA
quality is generated within the process.
Similar to the base processes with DMR and SMR, the firing duty within the com-
b ustion unit is determined by the process heat demand of the two reforming reactors
R-03 and R-04. A large flo w of exhaust gas (stream 9) is splitted and passes through
the shell-side of the tub ular reactors, which are manufactured from high temperature
resistant alloys. Do wnstream, the mixing stream still has a high temperature abov e 900
◦
C before entering the HRSG. The lo w-temperature heat is reco v ered on two pressure
le v els for electricity generation (see Figure 4.15). A single reheating is used to k eep the
steam content abo ve
x = 87 . 3%
at the outlet of the LP steam turbine T -02. Exiting the
HRSG the e xhaust gas (stream 25) still has a temperature of 120 ◦ C.
Figur e 4.15: T emperature profiles of heat transfer within the SMR+DMR process.
88

4.9 SMR-DMR Process
Figur e 4.16: Process flo wsheet of the methanol plant with steam and dry reforming.
89

Chapter 4 System Design and Modelling
T able 4.14: Simulation results for selected flo ws of the SMR-DMR process.
Flo w T ype T [ ◦ C ] p [bar] ˙ m [ kg/s ] ˙
E [ MW ]
1 Natural gas 15.0 10.0 45.6 2263.8
2 Natural gas 155.2 9.7 45.7 2266.5
3 Clean gas 590.1 33.0 73.8 1788.7
4 Syngas 1050.4 7.8 71.7 796.0
5 CO 2 550.0 9.2 59.3 48.7
6 Clean gas 700.0 8.8 71.7 632.4
7 Exhaust gas 1185.2 3.5 467.2 459.3
8 Syngas 1098.5 29.0 73.8 2191.6
9 Exhaust gas 1184.4 3.1 2172.9 2134.5
10 Exhaust Gas 923.8 3.1 2172.9 1544.7
11 Air 15.0 1.0 2077.3 12.3
12 Natural gas 15.0 5.0 27.9 1380.8
13 Syngas 693.6 28.7 145.5 2881.8
14 Syngas 250.0 50.2 417.7 7584.6
15 Syn. + CH 3 OH 38.0 46.4 417.7 7404.2
16 Pur ge gas 35.0 5.0 67.8 1121.7
17 Crude methanol 38.0 2.0 76.4 1664.8
18 Bottom product 77.5 1.6 68.5 1509.1
19 AAA methanol 80.4 1.9 65.9 1492.6
20 W ater 15.0 1.0 5.0 0.3
21 Steam 503.0 114.2 392.4 602.1
22 Steam 360.0 14.2 589.9 683.2
23 Condensate 32.9 0.05 589.9 103.7
24 W ater 24.8 15.0 703.0 36.6
25 Of fgas 120.0 1.8 2172.9 207.4
90

Chapter 5
Results and Discussion
This chapter presents the results obtained from the thermodynamic and economic analysis
for the six processes that ha v e been declared in the pre vious chapter 4. First, the process
performance will be in vestigated on the basis of se veral chemical and thermodynamic
indicators before an ener getic analysis is carried out. Particular attention will be de voted
to the
CO 2
utilization potential by in vestig ating the dif ferent inte gration concepts within
the processes. Subsequently , a con v entional e x er getic analysis is conducted to identify
the sources and the magnitude of the real inef ficiencies occurring within the process
units. An economic analysis is performed for an estimation of the in vestment cost and the
v ariable cost in volv ed in the projects. In this conte xt, the impact of market prices of fuel
and products on the re v enue and the contrib ution mar gin of the analysed systems is also
e xamined. Furthermore, an ex er goeconomic analysis is conducted to determine the cost
impact of each component on the o verall system and the le velized cost of the generated
products. Finally , for the most promising process, an adv anced ex er getic analysis is
conducted to analyse the interdependencies among the components and their potential
for an impro vement.
5.1 Pr ocess P erf ormance Anal ysis
In the follo wing, characteristic parameters will be discussed to understand the main
dif ferences resulting from the design specifications of the analysed processes. T able 5.1
sho ws a selection of results for a comparison of the thermodynamic characteristics of
the processes. Separately , T able 5.2 sho ws se v eral chemical parameters which assess
the performance of the reforming unit and the synthesis reactor . Furthermore, an energy
distrib ution of the processes is presented in Figure 5.1.
91

Chapter 5 Results and Discussion
T able 5.1: Selected results of the o v erall process.
Base pr ocesses Advanced pr ocesses
SMR A TR DMR CMR SMR-
A TR
SMR-
DMR
CH 3 OH-capacity mtpd 2590 2452 2719 2803 5440 5694
˙
W el,net MW 140.2 4.5 222.9 30.9 146.3 215.2
CH 4 process gas kg/s 16.9 22.8 25.3 16.9 43.9 45.6
CH 4 comb . fuel kg/s 28.7 - 19.9 14.9 11.8 27.9
Electric intensity MJ/kg 1.78 1.06 3.23 3.06 1.73 1.52
CH 4 -intensity 1 kg/MWh 202.0 126.1 175.1 149.5 126.4 154.9
HP li v e steam temp.
◦ C 440.0 600.0 600.0 600.0 600.0 503.0
MP li v e steam temp.
◦ C 300.0 423.0 415.0 - 420.0 -
LP li v e steam temp. ◦ C 152.0 218.0 200.0 250.0 144.0 360.0
HP li v e steam pres. bar 58.7 175.3 178.6 42 167 114.2
MP li v e steam pres. bar 27.1 62.8 64.5 - 45 -
LP li v e steam pres. bar 5.0 15.0 4.0 3.9 3.0 14.1
Exit temp. of fgas ◦ C 146.0 110.0 120.0 120.0 100.0 120.0
1 Ratio of total CH 4 intake to the total e x er gy of product
T aking into account technical limitations and scaling ef fects, the processes are designed
for dif ferent methanol production capacities gi v en in metric tons per day (mtpd). The
base processes are designed for a methanol production of 2,450-2,720 mtpd, which is the
typical capacity range for the majority of the plants in commercial operation. A single
A TR unit may also cov er a higher capacity of up to 5,000 mtpd [56]. The CMR process is
designed for a slightly higher capacity , while the other two adv anced prosesses represent
lar ge-scale applications abo ve 5,000 mtpd. Furthermore, in all processes, electricity
can be supplied to the grid as a second product. The net po wer
˙
W el,net
is particularly
high for the processes with endothermic reforming technology . Here, the electricity
generation is directly linked to the supply of process heat. In contrast, in the system with
autothermal reforming technology , the steam cycle can be designed to co v er only the
internal electricity demand.
F or all processes, heat is reco vered as steam within a tw o-pressure le v el HRSG.
Depending on the process design, the steam is expanded in tw o or three steam turbines.
The dif ferences in the steam c ycle design are mainly influenced by the internal steam
consumption of dif ferent units in the synthesis path and the pressure at which the steam
92

5.1 Process Performance Analysis
is required. In the SMR and SMR-DMR process, high temperature steam is primarily
used as a reforming agent and for preheating the reactants, thus reducing the HP li v e
steam temperature. The of fgas temperature is similar for most of the processes. In the
SMR process, lo w temperature heat is not inte grated into the synthesis loop, resulting in
a high of fgas temperature.
Furthermore, the electric intensity and the
CH 4
-intensity are pro vided in T able 5.1.
The electric intensity is calulated as ratio of the electrical duties, which are associated
with synthesis process chain to generate the chemical product, to the mass flow rate
of methanol. Thus, this number includes the po wer required for compression of the
natural gas, the syngas and the rec ycle streams. The electric intensity is particularly high
for the DMR and CMR process due to an ener gy-intensi v e syngas compression. The
CH 4
-intensity represents the ratio of the total natural gas intake, including the amount
for syngas generation and for comb ustion, to the e x er gy of product of the o v erall system.
F or obvious reasons, the processes with single endothermic reforming technology are
represented by a high
CH 4
-intensity . On the other hand, processes with autothermal
reforming are characterized by a lo w intensity .
SMR A TR DMR CMR SMR+A TR SMR+DMR
0
20
40
60
80
100
Relati v e amount of ener gy based on H i,NG [%]
CH 4 process gas CH 4 fuel Methanol product Net po wer
Air compr . Syngas compr . Recycle compr . Natural gas compr .
Pumps A GR ASU
Figur e 5.1: Ener gy distrib ution within the analysed processes.
93

Chapter 5 Results and Discussion
The bar chart in Fig. 5.1 sho ws detailed information on the ener gy distrib ution within
the processes. All v alues are gi ven as relati ve data based on the lo wer heating v alue of
the total methane intake. The left bar of each process represents the energy of the inlet
streams, which refer to the amount of natural gas used for syngas generation and process
heat supply , respecti v ely . Furthermore, the right bar shows the product distrib ution
(methanol and net electricity) and the internal electricity demand assigned to the main
consumers.
Considering the left bars, the structure of the energy supply v aries greatly for the
processes. Due to the exothermic character of the A TR, suf ficient process heat is
generated directly in the syngas track, thereby eliminating the need to supply natural g as
to the comb ustion unit. In contrast, the heat demand of the SMR and DMR technology is
co vered by the supply of na tural gas as a comb ustion fuel. Accordingly , the proportion of
natural gas intended for comb ustion is lo w for the SMR-A TR process. A clear assignment
of the methane intake to the respecti ve purpose is dif ficult, since the pur ge ratio varies
for the processes.
The amounts of ener gy output (right bar) should be treated with care, since products
of dif ferent ener getic quality are produced. Most of the supplied ener gy ob viously is
con verted into methanol, which is of lo wer energetic quality than electricity . In regard
to the po wer generation, for all processes, the major part of electricity is consumed
internally . The air compressor as the main consumer of electricity in particular has a
lar ge capacity for the processes with endothermic reforming technology . Here, the air
intake is higher , as the demand for process heat is larger . The syngas compressor as
the second lar gest consumer serv es to ov ercome the pressure dif ference between the
reforming unit and the synthesis unit. Especially for the processes with lo w-pressure
DMR and CMR, the pressure dif ference and thus the ener gy consumption is high. On the
other hand, an ener gy-intensi v e natural gas compressor is required for the high-pressure
reforming processes based on A TR and SMR technology . Due to relati v ely lo w volume
flo ws, a compression prior to reforming is less ener gy-intensi ve than do wnstream of the
syngas generation unit. Fig. 5.1 sho ws, that the ASU in the A TR and the SMR-A TR
process also constitutes a significant electricity consumer , caused by the multi-stage
compression of v ery lar ge air flo ws. Interestingly , the energy consumption of the rec ycle
compressors in all processes is comparati v ely lo w , despite a high pressure drop in the
synthesis loop. All other consumers, including the A GR in the CMR process, are also of
minor significance.
An interpretation of the pre vious results requires an in vestigation of the k ey chemical
components since their operation decisi v ely influences the o verall process characteristics.
94

5.1 Process Performance Analysis
T able 5.2 sho ws selected results for the reforming unit and the synthesis reactor . By
pro viding information on the composition of the syngas after the reformer and at the inlet
of the reactor , the impact of conditioning technologies and the recycle flo ws becomes
visible. Re garding the synthesis reactor , the con v ersion rates of
CO 2
,
CO
and
H 2
are
represented by
X CO 2
,
X CO
and
X H 2
. The mole ratio of the rec ycled flo ws to the syng as
make-up flo w is termed as recycle to feed ratio (R TFR). In general, the results are in line
with the information obtained by the parameter studies.
T able 5.2: Selected results for the reforming unit and the synthesis reactor .
Base pr ocesses Advanced pr ocesses
SMR A TR DMR CMR SMR-
A TR
SMR-
DMR
Syngas composition after r ef orming
CO 2 mole-% 4.9 4.4 2.8 10.8 2.9 2.3
CO mole-% 10.2 23.3 43.7 28.8 25.3 32.3
H 2 mole-% 49.0 56.1 42.4 58.4 58.2 57.6
CH 4 mole-% 0.3 1.1 3.6 0.0 0.0 0.9
H 2 O mole-% 35.2 14.1 7.5 1.6 12.8 6.5
Syngas composition bef or e the synthesis
CO 2 mole-% 16.4 10.1 10.3 0.9 3.0 6.9
CO mole-% 9.6 19.9 52.5 29.1 28.5 31.6
H 2 mole-% 68.5 61.4 28.2 66.7 63.5 56.8
CH 4 mole-% 1.4 2.9 6.5 0.3 0.3 1.9
H 2 O mole-% 0.5 0.4 0.6 0.0 0.2 0.8
Synthesis Reactor
R TFR mole/mole 5.9 2.7 5.8 1.7 2.3 3.1
X CO 2 % 8.0 2.1 6.9 17.3 2.4 6.0
X CO % 18.6 28.3 17.2 42.0 26.2 24.6
X H 2 % 10.9 19.4 31.1 37.9 23.4 24.4
CH 3 OH-yield mole-% 6.0 8.0 5.0 15.0 10.4 9.9
As presented in Section 4.1, the composition of the syngas generated by the single
reforming technologies in the base processes dif fers to a lar ge e xtent. The syngas
produced by SMR and A TR is characterized by a hydrogen surplus with reference to the
stoichiometry of the methanol synthesis. A R WGS unit would be an adequate mean to
95

Chapter 5 Results and Discussion
set a more fa v orable
CO
/
H 2
-ratio, while reducing the mole fraction of
CO 2
. In contrast,
an unconditioned syngas pro vided by DMR has a high surplus of carbon monoxide and
therefore would require a one-step WGS to impro v e the composition to w ards higher
methanol yield. For all base processes, the
CO 2
content is within an acceptable range.
In particular the syngas from SMR contains a lot of steam, which causes an additional
ener gy duty in the syngas compressor . In reg ard to the DMR process, a high methane
slip of 3.6% results in lar ge rec ycle flo ws of inert components in the synthesis loop. The
methane con version could be further increased by raising the mild operation conditions.
The syngas composition at the reactor inlet greatly de viates from the composition
of the reforming product. According to the con version rates
X CO 2
,
X CO
and
X H 2
, the
rec ycled syngas in the SMR and A TR is characterized by a lo wer mole fraction of
CO
and a higher mole fraction of
H 2
and
CO 2
(compared to the make-up syngas ). Thus,
through mixing of fresh and rec ycled syngas, the composition shifts to less fa v orable
stoichiometric ratios. In re gard to the DMR process gas, hydrogen is primarily acti ve in
the synthesis as indicated by the v alue of
X H 2
. Accordingly , the syngas composition at
inlet of the reactor is shifted to e v en lo wer v alues of
H 2
/
CO
ratio. Overall, com parati v ely
lo w methanol yields in the range of 5 - 8 mole-% result for the base processes. Thus, a
high R TFR of 5.9 and 5.8 results for the SMR and DMR process, while a moderate ratio
of 2.7 is obtained for the A TR process.
W ith respect to the adv anced processes, the reforming unit is designed for an im-
pro vement of the syng as composition for high methanol yield. At the reactor inlet the
process gas has a v alue for the
H 2
/
CO
ratio in the range of 1.8 - 2.3. In the CMR
process, the high
CO 2
content of 10.8% would result in v ery lar ge rec ycle streams. By
application of an A GR with physical absorption, the
CO 2
content is successfully reduced
to a minimum. The highest con version rates and accordingly the lo west R TFR result for
the CMR process. The methanol yield for the adv anced processes is in the range of 10 -
15 mole-%, reducing the R TFR to 1.7 - 3.1.
In general, a reduction of the circulated mass lo wers the in v estment cost for the com-
ponents of the synthesis unit. Furthermore, the amount of undesirable lo w temperature
heat from the synthesis is decreased, reducing the capacity of se v eral heat e xchangers.
96

5.2 Potential Analysis for Carbon Dioxide Utilization
5.2 P otential Analysis f or Carbon Dio xide
Utilization
The potential for carbon dioxide utilization by dif ferent inte gration strate gies is in vesti-
gated. The direct hydrogenation of
CO 2
as in the SMR process and the dry reforming by
CO 2
as in the DMR, CMR, and SMR-DMR process constitute the tw o main measures
for
CO 2
utilization in methanol synthesis [84, 147]. The integration not only abates
CO 2
-emissions by v alorization to a chemical intermediate, b ut also reduces the feedstock
and comb ustion fuel that is required for the production. For the calculation of the
CO 2
-
abatement and the
CH 4
-reduction, the processes are benchmarked against a con ventional
SMR process (Reference) without CO 2 -hydrogenation.
T able 5.3 sho ws the main results of the
CO 2
-abatement within the analysed processes.
F or a comparison of processes generating two products of dif ferent energetic quality , the
o verall e x er gy of product
˙
E P ,tot
seems to be the only reliable base for a comparison. The
ke y figures pro vided in T able 5.3 ha v e been defined in Section 2.2.1.
T able 5.3: K e y figures of the CO 2 -abatement potential within the analysed processes.
Ref. SMR DMR CMR SMR-DMR
˙
E P ,tot MW 693.78 812.54 929.13 765.97 1707.75
CH 4 -consumption t/h 159.13 148.00 162.78 114.26 238.42
CH 4 -reduction t/h 0 38.65 50.33 61.42 153.28
CO 2 -feed t/h 0 48.88 65.57 16.82 59.29
CO 2 -feed inten.
kg/MWh
0 60.15 70.58 21.96 34.72
CO 2 -emission t/h 286.91 333.88 504.10 174.93 580.67
CO 2 -emis. inten.
kg/MWh
413.54 410.91 542.55 228.37 340.02
CO 2 -abatement
kg/MWh
0 62.79 -58.43 207.13 108.24
CO 2 -abatement 1
kg/MWh
0 191.53 89.27 425.80 352.97
1 The ke y figure additionally considers the abated CO 2 associated with the CH 4 -reduction. An emission
factor of 55.9 t CO 2 /TJ is used for the calculation.
The results in T able 5.3 sho w that in all processes
CO 2
can be successfully integrated
for a reduction of the natural gas intak e. Ho we ver , other ef fects such as design impro v e-
ments also ha v e a significant influence on the fuel reduction and the other k e y figures. In
particular for the SMR and the DMR process, a lar ge amount of
CO 2
is inte grated per
MWh e xer gy of product. Furthermore, the number of the total
CO 2
-emissions is pro vided
to assess the true abatement potential of the processes. The amount of emitted
CO 2
by
97

Chapter 5 Results and Discussion
far e xceeds the amount of inte grated
CO 2
. The
CO 2
-abatement per MWh of ex er gy of
product is calculated by substracting the dif ference of the emission intensity and the feed
intensity of the considered process from emission intensity of the reference system. The
inte gration measures by direct hydrogenation and partial
CO 2
reforming successfully
lead to an abatement of
CO 2
per MWh of e xer gy of product. The highest
CO 2
-abatement
with 207.1 kg/MWh results for the CMR process, follo wed by the SMR-DMR process
and the SMR process with 108.3 kg/MWh and 62.8 kg/MWh, respecti v ely . The DMR
process of fers no abatement potential since dry reforming requires a lar ge amount of
process heat (pro vided by comb ustion gases), thus resulting in high
CO 2
emissions. If
the abated
CO 2
associated with the
CH 4
-reduction is taken into account with a f actor of
55.9
t CO 2
/TJ [148], an abatement potential results for all processes. Depending on the
CH 4 -sa vings, the CO 2 -abatement can reach considerably higher v alues.
5.3 Con ventional Exer getic Anal ysis
The results obtained from the ex er getic analysis for the o v erall system are gi v en in
T able 5.4. The processes are designed for dif ferent methanol production capacities and
co-generate electricity in some cases. The ex er gy concept is the only reasonable base for
a comparison of processes generating products of dif ferent ener getic v alue. A distrib ution
of the inef ficiencies occurring within the major subsystems (
˙
E D,sys
) is presented in Fig.
5.2. Since the processes ha ve the same feedstock b ut produce different products, the
inef ficiencies for the aggre gated subsystems are preferably related to the total e x er gy
of fuel
˙
E F ,tot
of the respecti v e process. The corresponding amounts of the inef ficiencies
are sho wn in Figs. E.1 - E.6 and their absolute numv ers are gi v en in T ables E.1 - E.6.
Furthermore, T able 5.5 sho ws a selection of the components with the highest e x er gy
destruction rate for each process.
The base processes with single SMR, A TR and DMR technology hav e an e x er getic
ef ficienc y of 31.1%, 56.9%, and 39.8%, respecti vely . The SMR and DMR design
necessarily feature a co-production of electricity . Ob viously , the ex er gy of fuel
˙
E F , t ot
for
processes with endothermic reforming technology is high due to the additional intake of
natural gas as a comb ustion fuel. The e x er gy losses which are mainly associated with the
of fgases are correspondingly high. In contrast, the A TR design features a lo w ex er gy of
fuel and relati v ely lo w ex er gy losses. Compared to the SMR and DMR design, ha ving an
e xer gy destruction ratio
y D
of 63.5% and 53.3%, respecti v ely , the major part of the ex er gy
of fuel in the A TR process is used to generate the ex er gy of product ( y D,A TR = 39 . 5%).
98

5.3 Con v entional Ex er getic Analysis
The adv anced processes ha v e a comparati vely high e x er getic ef ficienc y , indicating
a successful design selection from a thermodynamic point of vie w . The ef fi ciency for
the CMR, the SMR-A TR and the SMR-DMR design is 48.6%, 57.3% and 46.3%, re-
specti v ely . Interestingly , the CMR process is of high ex er getic ef ficienc y , despite the
fact that endothermic reforming technology is applied. This is due to the syngas condi-
tioning technologies, which ha v e a lo w e x er gy destruction, b ut allo w a relati vely lar ge
e xer gy of product. The SMR-A TR design has the highest ex er getic ef ficiency of all
studied processes. The thermoneutral combination of SMR and A TR only requires a
small amount fuel and at the same time generates a synthesis gas which enables a high
synthetic turno ver and a lar ge e xer gy of product. Re garding the SMR-DMR process, the
numbers in T able 5.4 indicate that the equipment in general is of v ery lar ge size. As a
result, the process might require a lar ge financial in vestment.
T able 5.4: Results obtained from the con ventional e x er getic analysis for the o verall system.
No.
Process ˙
E F ,tot [MW] ˙
E P ,tot [MW] ˙
E D,tot [MW] ˙
E L,tot [MW] ε [%]
1 SMR 2334.39 812.54 1327.88 194.81 34.8
2 A TR 1144.30 651.15 453.20 39.95 56.9
3 DMR 2336.87 929.13 1245.47 162.27 39.8
4 CMR 1576.45 765.97 653.66 156.83 48.6
5 SMR-A TR 2772.20 1587.02 1083.34 101.84 57.3
6 SMR-DMR 3689.79 1707.75 1735.50 246.54 46.3
F or a comparison of the ex er gy destruction rate on a subsystem le v el, e x er gy destruc-
tion ratios
y D,sys
are presented in Fig. 5.2. On a subsystem le v el, all processes are
characterized by the same e xer getic hierarch y: The majority of the irre v ersibilities occurs
within the reforming unit and the steam c ycle, while the e xer gy destruction rate for the
syngas compression, the methanol synthesis and the purification section is comparati v ely
lo w . Irre versibilities related to the pretreatment, the ASU and the A GR are of minor
significance.
In each process, more than 50% of the ov erall e x er gy destruction is attrib utable to the
reforming unit (
y ∗
D,sys
in the range
50 − 59%
). The subsystem includes the reforming
reactors, the comb ustion unit and the air compression. The subsystem’ s irre v ersibilities
are particularly high for the processes with endothermic technology . The main sources of
inef ficiencies refer to highly irre versible reactions occurring within the furnace (comb us-
99

Chapter 5 Results and Discussion
tion unit) and in the comb ustion zone of the autothermal reformer . The catalytic reactions
in the SMR and DMR ha v e a remarkable, albeit comparati v ely lo w e x er gy destruction.
The steam c ycle which serves the o v erall heat inte gration e xhibits the second highest
irre v ersibilities, accounting for approximately 25-30% of the o v erall e x er gy destruction.
The subsystem comprises v arious heat e xchangers including the HRSG, a set of steam
turbines and se v eral pumps. Splitting the steam c ycle into its components re v eals that for
the processes with endothermic reforming technology more than 60% of
˙
E D,sys
is caused
by the heat transfer within the HRSG. In re gard to the A TR and the SMR-A TR process,
the e xer gy destruction within the HRSG is much lo wer (around 30% of the steam c ycle’ s
˙
E D
) due to lo wer process heat demand of the reforming unit. In general, irre v ersibilities
within a heat e xchanger are also caused by the consideration of pressure drop, which
depends on the heat transfer area and the phsyical state of the flo ws. Based on high
isentropic ef ficiencies, the steam turbines ha v e a minor contrib ution to irre versibilities
within the steam c ycle. A significant ex er gy destruction rate occurs within the condenser
since useless lo w temperature heat, coming from the cooling of the crude product stream
in the synthesis loop, is rejected to the en vironment. This concerns in particular the
processes with a high R TFR (SMR, DMR and SMR-DMR).
The syngas compression unit in general consists of se v eral pre-coolers, tw o compres-
sors and one intercooler and has an e xer gy destru ction ratio
y ∗
D,sys
in the range of 4 -
7%. The irrev ersibilities are primarily caused by heat e xchange at high temperature
dif ference, by the pressure losses in the tubes of the heat exchangers and by friction
during compression. In particular for the pre-coolers, the temperature dif ference between
the process gas and the steam is high. The irre v ersibilities of the compressors are tak en
into account by an isentropic ef ficienc y of 85%. The extent of irre v ersibilities for this
subsystem is primarily determined by the quantiti v e influence of the mass flo w rate of
syngas and the pressure dif ference between the reforming unit and the synthesis unit.
100

5.3 Con v entional Ex er getic Analysis
0
5
10
15
20
25
30
35
40
˙
E D,sys / ˙
E F ,tot [%]
Pretreatment Reforming unit Syngas compression
Syngas conditioning Methanol synthesis Steam c ycle
Distillation ASU A GR
SMR A TR DMR CMR SMR-A TR SMR-DMR
Figur e 5.2: Results obtained from the con ventional e x ergetic analysis for the aggre g ated
subsystems.
The methanol synthesis unit has a small contrib ution to the o v erall inef ficiencies.
y ∗
D,sys
is in the range of 2 - 8%. The subsystem includes a recuperator and se v eral
heat e xchangers for crude product cooling as well as three recycle compressors and the
synthesis reactor . Interestingly , the synthesis reactor only has a lo w e x er gy destruction
rate which likely results from the heat transfer at lo w temperature dif ference between the
process gas and the boiling w ater and from the lo w con version rate. The major part of
the system’ s irre v ersibilities occurs within the product coolers due to a lar ge heat transfer
rate. Because of a high isentropic ef ficienc y of 87% and a moderate pressure drop in the
synthesis unit, the e xer gy destruction of the rec ycle compressors is lo w .
Depending on the reformer technology and the corresponding con version in the syn-
thesis , the amount of impurities in the crude product v aries. The number and size of
the distillation columns in the purification section as well as the ener gy demand in the
reboilers is particularly dri v en by the w ater content in the crude product. T o a minor
e xtent longer chained hydrocarbons (light ends) also play a role. In general, irre v ersibili-
ties within a distillation column are caused by mixing of streams with non-equilibrium
compositions and by dif ferences of temperature and pressure within the column trays
[149]. Additionally , heat transfer in the reboiler and the condenser also is associated
with e xer gy destruction. High irre v ersibilities particularly occur in the SMR and A TR
101

Chapter 5 Results and Discussion
process, due to a lar ge w ater content in the crude product. T wo-column distillation
systems ob viously ha v e a lo wer e x er gy destruction. Depending on the column features
and the crude product composition, the e x er gy losses which are dischar ged with the
bottom product v ary .
Due to minor significance, the ASU, the A GR and the conditioning unit will be
discussed briefly . Irre versibilities within the ASU are mainly caused by friction and heat
transfer within the intercooled multi-stage air compressor . W ith respect to the A GR, the
absorption process in the column has the highest irre v ersibilities. The catalytic shift
reactors in the syngas conditioning unit serv e for a slight adjustment of the composition
and therefore only cause a lo w e x er gy destruction.
The components in T able 5.5 are ranked by descending order of their e x er gy destruction
rate. The comb ustion units ha v e outstanding high irre v ersibilities. Additionally , se v eral
heat e xchangers of the HRSG are ranked among the components. Their e xer gy destruction
is caused by the lar ge heat transfer rate at high temperature dif ference between the hot
and the cold side. The distillation columns are also listed in the T able, while the recycle
and syngas compressors are not associated with a high e x er gy destruction.
102

5.3 Con v entional Ex er getic Analysis
T able 5.5: Results obtained from the con ventional e x er getic analysis for the components with the highest e xer gy destruction.
SMR A TR DMR CMR SMR-A TR SMR-DMR
˙
E F =2334.4 MW ˙
E F =1144.3 MW ˙
E F =2336.9 MW ˙
E F =1576.5 MW ˙
E F =2772.2 MW ˙
E F =3689.8 MW
Comp. k y D , k [ % ] Comp. k y D , k
[%]
Comp. k y D , k
[%]
Comp. k y D , k
[%]
Comp. k y D , k
[%]
Comp. k y D , k
[%]
R-04 29.97 R-09 11.64 R-05 25.54 R-05 18.62 R-07 12.92 R-05 21.77
HP-EV A 4.77 R-03 6.67 HP-ECO 3.16 HP-EV A 3.41 R-04 5.28 HP-EV A 3.56
HP-SH 1.59 C-02 1.79 HP-SH 2.46 R-03 2.58 HP-SH 2.11 LP-EV A 2.37
T -02 1.29 HP-EV A 1.16 E-09 2.09 HP-SH 1.90 E-13 1.37 HP-SH 1.65
E-05 1.32 E-03 1.12 R-04 2.08 E-07 1.51 C-02 1.30 E-11 1.30
C-02 1.27 E-08 0.96 LP-EV A 1.49 C-02 1.45 HP-EV A 1.29 LP-SH 1.29
CM02 0.95 C-03 0.92 E-05 1.36 E-05 1.13 HP-ECO 1.18 E-08 1.25
C-01 0.90 HP-SH 0.85 LP-RH 1.12 LP-SH 0.76 E-06 1.04 T -02 1.19
LP-SH 0.89 R-04 0.84 T -01 0.94 E-11 0.76 E-04 0.84 C-02 1.16
LP-EV A 0.59 E-09 0.70 C-02 0.89 CM-02 0.71 E-11 0.70 R-03 1.13
E-04 0.49 LP-EV A 0.69 T -02 0.59 T -01 0.67 E-03 0.51 RH 1.11
C-03 0.35 CM-02 0.65 E-06 0.58 E-06 0.60 C-01 0.49 E-05 0.85
E-08 0.34 T -01 0.62 E-03 0.52 LP-EV A 0.59 CM-02 0.48 R-06 0.59
E-01 0.24 E-02 0.55 MP-SH 0.47 C-01 0.50 T -03 0.48 CM-04 0.57
E-02 0.22 E-05 0.55 C-01 0.31 T -02 0.39 E-08 0.45 CM-03 0.57
103

Chapter 5 Results and Discussion
5.4 Economic Anal ysis
An e v aluation of the processes from a thermodynamic point of vie w allo ws a first
classification and e xclusion of unattracti v e designs. A holistic consideration, ho we v er ,
also requires an estimation of the major costs in v olved in a project. An economic
assessment of the in vestig ated processes is conducted by application of the total re v enue
requirement method (
T RR
). The assumptions underlying the analysis are gi v en in
Section 3.2. All monetary values are gi ven in US-$ 2016. In regard to the economic
sensiti vity analysis, a pricing of the main product methanol and the by-product electricity
is assumed. Furthermore, the costs for integrated and emitted
CO 2
is first assumed to be
zero, since the majority of methanol plants is located in regions without carbon pricing.
The influence of CO 2 pricing on the le v elized product cost is discussed in Section 5.5.
The results obtained from the economic analyses for the six processes are presented in
T able 5.6. The v alues of the
BMC
, the
F C I
and the
T C I
mainly depend on the capacity
of the respecti v e system. The breakdown of the
BMC
sho ws that the in vestment is mainly
assigned to the reforming unit (RE), ha ving a share of 60 - 70% in the
BMC
. If the
BMC
of the ASU is added to the reforming unit, the proportion of the subsystem is similar
high for the A TR and SMR-A TR process. The in v estment for the syng as compression
unit (COMP) is determined by pressure dif ference between the chemical units and by
the syngas v olume flo w . The synthesis unit (SYN) accounts for 7 - 13% of the
BMC
,
whereby the in vestment costs are proportional to the circulating mass flo w rate. In re gard
to the steam c ycle, an increased in v estment share is e xhibited for the processes with
endothermic reforming technology .
The distrib ution of the le v elized carrying char ges
C C L
, operation and maintenance
cost
OMC L
, and fuel costs
F C L
sho ws that the annual re v enue requirement
T RR L
and
therefore the economic feasibility is significantly influenced by the fuel costs. Consider -
ing the relati v e contrib ution per MWh of e x er gy of product, a dominating impact of the
specific le v elized fuel costs
f c L
particularly is gi v en for the processes with endothermic
reforming technology . Especially for the SMR process, the fuel cost respresent a very
high proportion in the cost of the end-products. The A TR and the SMR-A TR process
profit from a high ef ficienc y and therefore are associated with the lo w specific fuel cost.
At this point it should be emphasized that the specific costs per unit of e x er gy form
an a v erage v alue o v er two products and therefore cannot be equated with the le velized
cost for electricity (LCOE) and methanol (LCOM). In most process designs, the specific
carrying char ges
cc L
and operation and maintenance costs
omc L
account for less than
50% of the total le v elized cost per unit of e x er gy product. Slightly increased specific
104

5.4 Economic Analysis
costs for the in vestment arise for the SMR and the DMR process. The relati v ely lo w
specific carrying char ges for the lar ge-scale processes SMR-A TR and SMR-DMR can be
deri v ed from ef fect of the economies of scale. Finally , the economic feasibility of the
processes not only depends on the fuel cost b ut also on the future de v elopment of the
market prices for the generated products methanol and electricity .
T able 5.6: Results obtained from the economic analysis for the o v erall systems.
SMR A TR DMR CMR SMR-
A TR
SMR-
DMR
BMC 10 6 US$ 507.37 308.87 505.24 358.29 594.93 642.03
PRE 1 % BMC 2.09 5.06 0.92 0.71 4.87 0.88
RE 2 % BMC 64.51 41.88 60.43 51.46 44.65 68.90
COMP
3 % BMC 4.89 7.95 9.12 17.85 5.77 4.42
CON
3 % BMC - - - 0.08 5.30 -
SYN
4 % BMC 13.04 11.47 8.35 7.89 10.07 7.33
SC 5 % BMC 10.87 8.33 17.70 11.13 12.27 15.13
DISTL
6 % BMC 4.61 7.83 3.47 4.49 3.94 3.34
ASU % BMC - 17.48 - - 13.13 -
A GR % BMC - - - 6.39 - -
F C I 10 6 US$ 636.42 396.47 633.19 450.39 743.89 804.37
T C I 10 6 US$ 674.73 420.34 671.31 477.50 788.66 852.79
C C L 10 6 US$/a 68.72 42.81 68.37 48.63 80.37 86.86
OMC L 10 6 US$/a 62.76 40.52 62.46 47.61 76.47 122.73
F C L 10 6 US$/a 249.99 132.38 247.97 174.00 305.46 402.72
T RR L 10 6 US$/a 381.5 215.72 378.80 270.25 462.26 612.31
cc L 7
US$/
MWh 10.57 8.21 9.19 7.94 6.33 6.36
omc L
US$/
MWh 9.65 7.78 8.40 7.77 6.02 8.98
f c L
US$/
MWh 43.14 25.41 33.36 28.40 24.06 31.72
1
Subsystem Pretreatment
2
Subsystem Reforming Unit
3
Subsystem Reforming Unit
3
Subsystem
Syngas Conditioning
4
Subsystem Synthesis Unit
5
Subsystem Steam Cycle
6
Subsystem Distillation
7 The specific cost are related to ˙
E P ,tot
105

Chapter 5 Results and Discussion
In general, these economic decision parameters are subject to temporal and re gional
fluctuations. Thus, the estimation of the economic feasibility of a methanol plant is a
comple x task and dif ficult to predict. In the follo wing, sensiti vity analyses are carried
out to sho w the impact of the le v elized fuel cost on the re venues and on the contrib ution
mar gin. From the in vestig ations, the interdependence of the minimum selling prices of
electricity and methanol for cost reco very can be deduced. Furthermore, the analyses
result in a determination of break e v en points from which an economic f a v orability of a
specific process can be deri v ed under gi v en mark et conditions.
An economic blackbox model is applied to analyse the relations between the inlet cost
streams, the in v estment and the re venues from the selling of the products. Since one
fuel stream and two product streams mean three de grees of freedom, two parameters
must be specified - the specific fuel cost
f c
and the sale price for electricity . Information
about the e xact cost distrib ution on the products is later obtained by the application of
the e xer goeconomic analysis.
All cost parameters included in the sensiti vity analyses refer to le velized v alues to take
into account temporal changes o ver an operation time of 20 years. The annual product
cost for methanol and electricity (
C M eoh
and
C el ec
) result from the calculation of the total
annual re v enue requirement and are obtained by solving the Eq. 5.1 -5.3.
T RR L = C C L + F C L + OM C L (5.1)
C M eoh + C E l ec = C C L + F C L + OM C L (5.2)
MC M eoh = C C L + ˙ m fuel · LH V · τ · f c · C E LF fuel + OMC L − ˙
W net · τ · M C el ec
˙ m M eoh ∗ τ (5.3)
The range of whole sale prices per MWh of generated electricity
MC el ec
was selected
based on information reported in [116–118], while the cost range per GJ of natural
gas w as determined based on data obtained from [150, 151]. For an assessment of
the economic feasibility , the calculated minimum price per metric ton of methanol is
compared with market data published by the w orld’ s lar gest methanol production and
distrib ution compan y Methane x Corporation (Methanex), Canada [28, 98]. The data
basis includes monthly , continental wholesale contract prices for the last fi ve years. The
comparison with real market data also serv es to v alidate the results of the economic
analysis.
The Figs. 5.3 - 5.8 sho w the sensiti vity of the minimum methanol price to the cost of
natural gas and the selling price of electricity . The solid lines represent the process
106

5.4 Economic Analysis
specific isolines of constant methanol price to co ver the TRR. Accordingly , their slope
is a measure for the dependenc y of the minimum methanol price on the sale price of
electricity . The dotted and dashed graphs represent the market data and sho w the 5-year
high and lo w of the methanol wholesale price in selected continents. The graphs only
serv e as a reference for the isolines of the processes and should therefore not be related
to the axis. Furthermore, the e xact product cost distrib ution (for the reference price of
2.9 US
$ 2016
/GJ), which is obtained from the ex er goeconomic analysis, is brok en do wn
for the reference cases. In order to understand the graphics, the reference case of the
SMR process will be e xemplarily e xplained. For the sensiti vity analysis, it is assumed
that the respecti v e plant participates in a mark et with electricity pricing.
In principle, a distinction can be made between two process types - those with en-
dothermic reforming technology , in which a considerable amount of excess electricity is
generated, and those with exothermic or mix ed reforming technology . Reg arding the first
group of processes, Figs. 5.3, 5.5, 5.7, and 5.8 sho w that the wholesale price of electricity
decisi v ely influences the minimum product price for methanol to co v er the
T RR L
. Fo r
the processes with A TR and CMR technology in Figs. 5.4 and 5.6, the mar ginal influence
results from the lo w po wer generation.
Figur e 5.3:
Impact of the natural gas cost and the sales price of electricity on the minimum sale
price of methanol for the SMR process.
107

Chapter 5 Results and Discussion
As sho wn in T able 5.6, the products generated in the SMR process are associated with
high specific a v erage cost per MWh. In addition to relati v ely lar ge carrying char ges,
the product costs are high due to the lar ge
CH 4
-intensity and thus to the fuel costs.
F or the reference case, the le v elized fuel cost of approximately 3.9 US$/GJ result in
le v elized electricity cost (LCOE=
MC el ec , re f
) of 110 US$/MWh and le v elized methanol
cost (LCOM=
MC M eoh , re f
) of 295 US$/mt. In case the le v elized fuel cost
f c , A
increase to
7.8 US$/GJ at constant re v enues from electricity
MC el ec , re f
, the minimum methanol price
MC M eoh , A
would amount to 590 US$/mt to co v er the
T RR
. Under these conditions an
economic operation in Europe most likely is not possible, while an economic feasibility
of the project could be gi v en in Asia and North America. W ith respect to case B, an
increased le v elized sale price of 168 US$/MWh for electricity at constant
f c , B
would
result in reduced minimum cost of methanol MC M eoh , B of 216 US$/mt.
In general, for fuel cost abov e 7.5 US$/GJ, a ton of methanol cannot be produced at an
economically competiti v e price. As for the other processes with endothermic reforming
technology (e.g, DMR process in Fig. 5.5), the relativ ely steep incline of the isolines
indicates a lar ge dependence on selling price of electricty . T aking into account European
electricity procurement prices between 50-80 US$/MWh [116] and an a v erage European
natural gas price of 4.5 - 5.5 US$/GJ [152] (6 - 7.5 US$/GJ on a le v elized base), a
cost-co vering operation w ould require a minimum methanol price of 600 - 650 US$/mt.
Based on this market data and the historical methanol prices [28], a domestic production
would be associated with great economic uncertainties. W ith respect to the Asian and
North American market, an economic feasibility could be gi ven at le velized natural gas
cost of up to 7 US$/GJ.
Fig.5.4 sho ws the results obtained from the sensiti vity analysis of the A TR process.
Due to lo w po wer generation, the re v enues from the selling of electricity nearly ha v e
no influence on the minimum methanol price. In comparison with the endothermic
reforming processes (SMR, DMR, CMR, SMR-DMR), methanol can be produced at
relati v ely lo w cost. Ev en for le v elized g as prices of up to 12 US$/GJ the plant could still
be operated economically . The currently lo w natural g as prices of less than 3 US$/GJ in
North America (4.1 US$/GJ on a le v elized base) w ould result in methanol production
cost of 250 - 260 US$/mt. Based on the European natural gas price of 4.5 - 5.5 US$/GJ
in 2016 [152] (6 - 7.5 US$/GJ on a le v elized base), minimum methanol cost of 330 - 400
US$/mt would result. A European production of methanol from natural gas by means of
A TR technology therefore may be feasible.
The results obtained from the sensiti vity analysis of the DMR process are presented
in Figure 5.5. The process generates a lar ge amount of e xcess electricity , which could
108

5.4 Economic Analysis
Figur e 5.4:
Impact of the natural gas price and the sales price of electricity on the minimum sale
price of methanol for the A TR process.
Figur e 5.5:
Impact of the natural gas cost and the sales price of electricity on the minimum sale
price of methanol for the DMR process.
109

Chapter 5 Results and Discussion
constitute an additional source of income, thus influencing the minimum methanol price.
Under the terms of a high electricity sale price abo ve 80 US$/MWh and a lo w natural
gas mark et price in the range of 0 - 3 US$/GJ, methanol could be of fered for 0 US$/mt
to co ver the TRR of the plant. Thus, the plant could generate profits as a pure electricity
market participant. For a
MC el ec
belo w 40 US$/MWh, the
MC M eoh
is in the same range
as for the SMR process in Figure 5.3. W ith increasing electricity prices, cost adv antages
for the DMR process arise due to the lar ger po wer capacity (see T ab . 5.1). Compared
with the A TR process, the DMR process is economically only competiti v e for lo w cost
of natural gas or a high sale price of electricity . The reference plant with fuel cost 2.9
US
$ 2016
/GJ generates electricity and methanol at le v elized cost of 100 US$/MWh and
220 US$/mt. Considering current market conditions in Europe (see SMR process), a
metric ton of methanol could be sold for a minimum price in the range of 400 - 500
US$/mt. The results of the sensiti vity analysis sho w that a methanol production with
CO 2
-v alorization can be economically feasible and competiti v e. Howe ver , the economic
feasibility is essentially determined by the re v enues from selling of electricity . In addition,
the le v elized product costs will increase if costs are assigned to the inte grated CO 2 .
Figur e 5.6:
Impact of the natural gas cost and the sales price of electricity on the minimum sale
price of methanol for the CMR process.
Figure 5.6 illustrates the results of the sensiti vity analysis for the CMR process. The
slope of the isolines is less steep compared with the SMR and DMR process since the
110

5.4 Economic Analysis
po wer output is significantly lo wer . Abo v e a natural gas price of 11 US$/GJ, an operation
of the plant may not be economically viable. Under the current market conditions, the
production cost per ton or methanol are in the range of the respecti v e continental contract
price of Methane x. At
MC el ec
belo w 40 US$/MWh, the ton of methanol can be sold for
a lo wer price in comparison with the DMR process. Thus, bi-reforming constitutes a
good alternati v e (to the DMR process) for
CO 2
-v alorization at competiti ve product cost
and with increased economic independence. The reference plant with fuel cost of 2.9
US
$ 2016
/GJ generates electricity at high le v elized cost of 165 US$/MWh and methanol
at le v elized product cost of 250 US$/mt.
Figur e 5.7:
Impact of the natural gas cost and the sales price of electricity on the minimum sale
price of methanol for the SMR-A TR process.
The SMR-A TR process was designed for a lar ge capacity in order to use the ef fect of
the economies of scale for decreasing the product cost. Re garding the product-related
fix ed cost (carrying char ges
cc L
) in T able 5.6, the process clearly sho ws a cost benefit
o ver the processes with single reforming technology . Considering Fig. 5.7, the sale price
of electricity has a considerable influence on
MC M eoh
as indicated by the steep slope.
Ob viously , the SMR-A TR process can generate methanol at the lo west cost, re gardless of
the sale price for electricity . Similar to the DMR process, methanol could e v en be of f ered
free of char ge at lo w natural gas prices and high selling prices of electricity . For the
111

Chapter 5 Results and Discussion
reference plant, the le v elized cost of electricity amount to 75 US$/MWh while methanol
is generated at product cost of approximately 200 US$/mt.
Figur e 5.8:
Impact of the natural gas cost and the sales price of electicity on the bottom price of
methanol for the SMR-DMR process.
As sho wn in T able 5.1, the SMR-DMR process was also designed for a lar ge capacity
to reduce the capital cost by the ef fect of the economies of scale. The product-related
carrying char ges in T able 5.6 sho w that this attempt w as successful. Ho we ver , the a v erage
product cost are high due to the high fuel-intensity of the process. An operation abov e a
natural gas price of 11 US$/GJ most lik ely would result in financial losses. Considering
European market conditions, a minimum methanol price of 400 - 500 US$/mt would
be necessary to reco ver the total cost. For the reference case, electricity is generated at
le v elized cost of 122 US$/MWh while the le v elized product cost for methanol amount to
210 US$/mt.
In the economic consideration it may be useful to consider only the v ariable costs, while
ne glecting the fixed costs associated with the in vestment (sunk costs). The contrib ution
mar gin
CM
is calculated as the dif ference between the re venues from the selling of the
products
C M eoh + C E l ec
and the v ariable cost associated with the fuel procurement
FC L
and the operation
OMC L
. A positi v e contrib ution mar gin therefore represents the annual
amount of income that is a v ailable to co v er the fix ed cost. In this analysis the operation
and maintenance cost are e xclusi v ely added to the v ariable cost. The Figs. 5.9-5.11 sho w
112

5.4 Economic Analysis
the threshold, at which the re venues can be used to pay of f the fix ed costs. For each
process, this threshold is represented by the isoline for a contribution mar gin of zero.
The isolines are obtained by solving Eq. 5.4 and Eq. 5.5.
CM = C M eoh + C E l ec − FC L − OMC L with CM = 0 (5.4)
MC M eoh = ˙ m fuel · LH V · τ · f c · C E LF fuel + OMC L − ˙
W n et · τ · MC el ec
˙ m M eoh ∗ τ (5.5)
Depending on the selling prices of the products, one area of the positi ve contrib ution
mar gin results abo ve and one area of the ne g ati v e contrib ution mar gin belo w the respec-
ti v e isoline. The Figs. 5.9 - 5.11 therefore sho w at which le velized product market prices
each of the plants can be operated economically under gi v en le v elized fuel cost of 3
US$/GJ, 7 US$/GJ, and 11 US$/GJ, respecti v ely . Furthermore, break ev en points can be
read from the intersections of the isolines.
Figur e 5.9: Sensiti vity of the contrib ution mar gin depending on the minimum methanol price
and the selling price of elecrtcity for fuel cost of 3 US$/GJ.
113

Chapter 5 Results and Discussion
Figur e 5.10: Sensiti vity of the contrib ution mar gin depending on the minimum methanol
price and the selling price of elecrtcity for fuel cost of 7 US$/GJ.
W ith rising fuel cost
f c
, the minimum price per ton methanol
MC M eoh
increases, as-
suming a constant le v elized electricity selling price. Depending on the market conditions
for methanol and electricity , dif ferent processes are economically more attracti ve. For
fuel costs of 3 US$/GJ, the methanol can be sold in a range of 0 to 470 US$/mt to obtain
a positi v e contrib ution mar gin for each process. The wide price range is determined by
the fuel intensity and the pricing for electricity . F or lo w re venues from the selling of
electricity , the A TR and SMR-A TR process clearly ha v e an economic adv antage. On
the other hand, for high re v enues from the selling of electricity (abo v e 100 US$/MWh),
methanol from processes with endothermic reforming technology (DMR, SMR-DMR)
could be of fered for a cheaper price to reco v er the v ariable cost. Under consideration
of historical methanol prices, an economic feasibility might be gi v en for all considered
processes.
114

5.4 Economic Analysis
Figur e 5.11: Sensiti vity of the contrib ution mar gin depending on the minimum methanol price
and the selling price of elecrtcity for fuel cost of 11 US$/GJ.
F or fuel cost of 7 US$/GJ (see Fig. 5.10), methanol can be produced at cost of 300
- 800 US$/mt. The SMR and DMR process can only generate a positi v e contrib ution
mar gin for v ery high product prices. Already for an electricity sale price of 40 US$/MWh
the SMR-A TR reaches a break-e v en with the A TR process. Furthermore, the economic
adv antage of the A TR and SMR-A TR process increases compared to the designs with
endothermic reforming technology . For fuel cost of 11 US$/GJ, only the processes
with autothermal reforming (A TR and SMR-A TR) could possibly achie v e a positi v e
contrib ution mar gin. No w , these processes hav e a clear cost-adv antage ov er the processes
with endothermic reforming technology . In case of such high long-term prices, alternati ve
syngas generation processes based on coal most lik ely would be the preferred option
for methanol production. In these considerations, it should be noted, that a year -round
sale of electricity at prices abo ve 100 US$/MWh constitutes an unrealistic scenari o.
In conclusion, a clear economic hierarchy results for the in vestigated processes. The
follo wing ranking is obtained for ascending methanol product cost to co v er the v ariable
cost of the project: A TR process, SMR-A TR process, CMR process, followed by the
SMR-DMR process, the DMR process and finally the SMR process. Cost adv antages of
the processes with endothermic technology especially arise for high electricity prices at
lo w fuel cost.
115

Chapter 5 Results and Discussion
5.5 Ex er goeconomic Anal ysis
An e xer goeconomic analysis is conducted in order to understand the formation and the
flo w of costs within the processes. Information on the cost distrib ution, the real le v elized
product cost and the economic significance of each component can be deri ved. Details
on the application of the methodology are pro vided in Section 3.3. For each process, the
results obtained from the e x er goeconomic analysis are presented for a reference case
using a natural gas is 2.9 US$
2016
/GJ. In Fig. 5.12 a relati v e cost distrib ution is gi ven for
the o verall processes, whereby the in vestment cost
˙
Z t ot
and the cost rates associated with
the inef ficiencies ˙
C L , t o t and ˙
C D , t o t are related to the total e xer gy product ˙
E P , t ot .
SMR A TR DMR CMR SMR-A TR SMR-DMR
0
20
40
60
4 . 6
3 . 7
4 . 3 5 . 4
2 . 9
5 . 1
21 . 9
9 . 3
17 . 8 16 . 5
9 . 4
13 . 9
20 . 2 16 17 . 6 15 . 7 12 . 4 15 . 3
Relati v e cost rate [US$/MWh]
˙
Z t o t / ˙
E P , t ot
˙
C D , t o t / ˙
E P , t ot
˙
C L , t o t / ˙
E P , t ot
Figur e 5.12: Relati v e cost rate associated with the in vestment and the inef ficiencies of the
ov erall processes.
The relati v e cost rate of the in vestment
˙
Z t ot
˙
E P , t ot
includes the carrying char ges and the
operation and maintenance cost. In particular , the SMR and DMR process are capital
intensi v e due to the lar ge in vestment associated with the reforming unit. Re garding the
adv anced processes, the v alue of
˙
Z t ot
˙
E P , t ot
is decreased due to a design impro vement and the
ef fect of the economies of scale.
T urning ne xt to the relati v e cost rate associated with the e xer gy destruction
˙
C D , t o t
˙
E P , t ot
,
the irre v ersibilities ob viously e xhibit the major cost source for the base processes with
endothermic reforming technology . In reg ard to the A TR and the SMR-A TR process,
116

5.5 Exer goeconomic Analysis
˙
C D , t o t
˙
E P , t ot
has a lo wer cost importance. The processes are of higher ex er getic ef ficiency due
to reduced irre v ersibilities within the reforming unit. Furthermore, the relativ e cost rate
associated with the e xer gy losses of the ov erall system
˙
C L , t o t
˙
E P , t ot
is sho wn in Fig. 5.12. The
relati v e cost rate is particularly high for the SMR, DMR, CMR, SMR-DMR process
since the e xer gy losses are mainly attrib uted to the of f gases which are large for the
endothermic reforming technologies. Accordingly , the influence of the cost rate is lo wer
for the A TR and the SMR-A TR process.
In order to impro ve the cost ef fecti veness of a system, a detailed thermoeconomic
e v aluation has to be conducted on a component le v el. F or each process, the results for
a selection of components with the highest cost importance are presented in T ables 5.7
- 5.12. Accordingly , the components are ranked by descending order of their cost rate
associated with the e xer gy destruction and the in v estment (
˙
C D , k
+
˙
Z k
). The ex er goe-
conomic factor
f k
and the relati v e cost dif ference
r k
are used to e v aluate the internal
cost distrib ution and the relati v e cost increase between
c P , k
and
c F , k
. By analysing the
e xer goeconomic v ariables, recommendations concerning changes of the design and of
operation parameters can be made in order to decrease the le v elized product cost within
an iterati v e impro v ement process. From the results of all processes in T ables 5.7 - 5.12,
some general conclusions can be dra wn:
1.
The comb ustion unit, the air compressors and se v eral heat e xchangers within the
synthesis train and the steam c ycle ha v e a high cost impact. .
2.
Re garding the heat e xchangers, the e xer gy destruction
˙
C D , k
represents a dominant
cost source, as indicated by a lo w e x er goeconomic factor .
3.
The compressors ha v e a high v alue for the e x er goeconomic factor , indicating that
their in vestment ˙
Z k constitutes the major cost source.
4.
The circulating streams within the synthesis unit are associated with high a v erage
cost per unit of e xer gy .
5.
The processes with
CO 2
inte gration could be economically adv antaged, since
CO 2
was assumed to be free of cost.
6.
The a v erage cost per unit of e x er gy of the syngas strongly v aries with the type of
reforming agent. High av erage cost particularly result from the supply of costly
oxygen in the process with autothermal reforming.
117

Chapter 5 Results and Discussion
Re garding the reference SMR process, the detailed results obtained from the ex er -
goeconomic analysis are presented in T able 5.7. The relativ e cost dif ference
r
for the
o verall plant amounts to 338.5%, whereby the a v erage cost per unit of e x er gy product
c p , t ot
includes the cost of dissipati v e components and the cost rate associated with e x er gy
losses. The relati v e cost dif ference for electricity (
r el ec = 747 . 1%
) is much higher than
the one for methanol (
r M eoh = 253 . 1%
). The ex er goeconomic f actor
f t ot
has a v alue
of 46.7%, indicating that the le v elized product cost can potentially be decreased by an
impro vement in the e x er getic ef ficiency at the e xpense of higher in vestment cost.
On a component base, the steam reformer R-04, the air compressor CM-02 and the
high temperature syngas cooler E-06 ha v e the highest v alue of the sum
˙
Z k + ˙
C D , k
and
therefore represent the most important components from a thermoeconomic point of vie w .
In re gard to the reformer , the lo w e x er goeconomic f actor sho ws that the costs are almost
e xclusi v ely due to the e x er gy destruction. Thus, the cost impact of the SMR can be
reduced by an increase in the e x er getic ef ficienc y . Corresponding measures concern the
use of an adv anced b urner technology to reduce the irre versibilities associated with the
comb ustion reactions (by lo wering the air to fuel ratio) and the reduction of the S/C ratio
to decrease irre v ersibilities due to mixing. Furthermore, a preheating of the comb ustion
reactants might also lo wer the e x er gy destruction within R-04. The air compressor unit
CM-02 is the most e xpensi v e component of the process and has the second highest v alue
of the sum
˙
Z k + ˙
C D , k
. The high v alue of the ex er goeconomic factor
f
suggests a reduction
of the capital in vestment at the e xpense of lar ger irre v ersibilities. For this purpose, the
high isentropic ef ficienc y (
η s
= 0.86) and the pressure ratio could be reduced. The
pressure ratio, ho we ver , is fixed by the outlet pressure of the of f g ases and the pressure
drop within the HRSG.
The high temperature syngas cooler E-06 in the syngas compression unit also has
a remarkable high v alue of
˙
Z k + ˙
C D , k
causing a high cost dif ference
r k
. Considering
the lo w
f
v alue, the cost increase is exclusi vely due to e x er gy destruction. Thus the
irre v ersibilities should be reduced at the e xpense of an increase of the capital in v estment
and the O&M cost. The cost rate
˙
C D , k
is due to lar ge irre v ersibilities and the high
a v erage cost at which e xer gy is supplied to the component
c F , k
. The e x er gy destruction
is mainly caused by heat transfer at high temperature dif ference of up to 400
◦
C. An
inte gration into the heat management system at a higher temperature le v el (increase in
thermodynamic a v erage temperature on the cold side) would decrease the impact of
˙
C D , k
,
b ut is impossible due to internal heat distrib ution of the process and the corresponding
design specifications of the steam c ycle. An extension of the heat transfer area or an
use of materials with a higher heat transfer coef ficient may reduce the cost rate
˙
C D , k
to a
118

5.5 Exer goeconomic Analysis
certain e xtent. The high-pressure e v aporator HP-EV A in the HRSG is also characterized
by lar ge irre v ersibilities, which constitute the main cost source for this component. An
increase in the steam pressure in order to reduce the irre v ersibilities w ould require major
changes within the heat inte gration management system, since the v apor fraction at the
outlet of the lo w-pressure turbine is already at the lo wer boundary . By an increase of
the heat supply , the steam parameters could be increased to reduce the irre v ersibilities
due to heat transfer . Ho we v er , the cost rate associated with the ex er gy losses of the of f
gases w ould increase. The recuperator E-04 in the synthesis unit also has a high cost
importance for the SMR process. According to the lo w e x er goeconomic factor of 28.5%,
a cost reduction might be achie v ed by an increase in the e x ergetic ef ficiency of the
component. Ho we ver , The lar ge cost rate ˙
C D , k is rather caused by the high a v erage cost
per unit of ex er gy of fuel, which results from the costly rec ycle streams of uncon verted
syngas. Thus, approaches to reduce the sum
˙
C D , k
+
˙
Z k
should consider a reduction of the
cost that were supplied upstream to the process gas.
The main condenser E-05 is rated to position number six and has a lo w
f
v alue. The
cost rate
˙
C D , k
results from the lar ge irre v ersibilities and the high a v erage cost
c F , k
of
useless steam. By changing the li v e steam parameters, the temperature of the inlet
steam could be reduced, thus decreasing the e xer gy destruction. The irrev ersibilities also
constitute the major cost source for the high-pressure superheater (HP-SH). The li v e
steam temperature is limited to 600
◦
C, while the comb ustion gases from the radiati ve
zone of the SMR enter the heat e xchanger with a temperature of approximately 900
◦
C.
The ke y decision v ariables include the steam pressure, the li ve steam temperature and the
inlet temperature of the comb ustion gases. By an increase of the HP pressure le vel and
the li v e steam temperature, the e x er gy destruction rate could be reduced. Interestingly ,
despite a high isentropic ef ficienc y
η s
, the irre v ersibilities also ha v e a high impact on the
cost of the high-pressure steam turbine T -02. A further increase of
η s
and a reduction of
the pressure ratio would reduce the e x er gy destruction.
The other components are only briefly discussed since their cost influence on the
o verall process is of lo wer significance. The highly inefficient distillation columns C-01,
C-02 and C-03 lead to a significant increase in the a v erage cost per unit of e x er gy as
indicated by a relati v e cost dif ference of each abov e 250%. Their cost importance is
mainly attrib uted to the inef ficiencies which result from the heat transfer within the
condenser and the reboiler . By reducing the formation of water in the synthesis, the
impurities of the crude product and thus the irre v ersibilities could be lo wered. Due to a
lo w steam mass flo w rate within the lo w pressure le v el of the HRSG, the sum of
˙
C D , k
+
˙
Z k is comparati v ely lo w for the LP-SH and LP-EV A.
119

Chapter 5 Results and Discussion
T able 5.7: Results obtained from the e x ergoeconomic analysis for a selection of components of the reference SMR process.
Comp. k ˙
E F , k ˙
E P , k ˙
E D , k ε k y D , k c F , k c P , k ˙
C D , k ˙
Z k ˙
C D , k + ˙
Z k f k r k
[MW] [MW] [MW] [%] [%] [$/GJ] [$/GJ] [$/h] [$/h] [$/h] [%] [%]
R-04
1,669.71
970.20 699.51 58.1 30.0 3.97 7.53
10,004.95
2,410.65 12,415.60 19.4 89.5
CM-02 203.86 181.70 22.16 89.1 1.0 25.49 40.87 2,033.07 8,027.91 10,060.98 79.8 60.4
E-06 97.77 40.14 57.64 41.1 2.5 38.61 94.32 8,012.08 36.47 8,048.55 0.5 144.3
HP-EV A 436.16 324.87 111.29 74.5 4.8 12.95 17.40 5,187.03 21.73 5,208.76 0.4 34.4
E-04 64.69 53.25 11.44 82.3 0.5 54.99 71.52 2,264.60 904.02 3,168.62 28.5 30.1
E-05 34.64 - 30.84 - 1.3 21.33 - 2,368.42 9.15 2,377.56 0.4 -
HP-SH 129.89 92.85 37.04 71.5 1.6 12.95 18.18 1,726.46 21.87 1,748.34 0.0 40.4
T -02 380.36 350.22 30.14 92.1 1.3 22.77 25.49 2,470.44 956.34 3,426.80 27.9 11.9
E-08 23.14 15.20 7.94 65.7 0.3 54.99 84.42 1,571.29 39.00 1,610.29 2.4 53.4
C-02 42.53 12.95 29.58 30.5 1.3 12.10 45.73 1,288.26 279.83 1,568.09 17.9 278.0
LP-SH 28.14 7.36 20.78 26.2 0.9 12.95 50.35 968.28 22.97 991.25 0.0 288.9
C-01 26.21 5.14 21.07 19.6 0.9 12.10 64.86 917.66 58.22 975.89 6.0 436.1
E-02 50.54 45.45 5.09 89.9 0.2 38.61 43.36 707.25 69.0 776.21 8.9 12.3
CM-08 15.97 14.01 1.95 87.8 0.1 23.95 37.56 168.52 518.07 686.59 75.5 56.8
LP-EV A 79.17 65.30 13.86 82.5 0.6 12.95 15.80 646.05 23.99 670.03 3.6 22.0
T otal
2,334.39
812.54
1,327.88
34.8 56.9 3.72 16.31 *
17,776.30 16,435.73
34,212.04 46.7 338.5
* includes ˙
C L and ˙
C Di f f
120

5.5 Exer goeconomic Analysis
The results obtained from the e xer goeconomic analysis of the A TR process are pre-
sented in Fig. 5.8. The relati vely lo w v alue of the sum
˙
C D , t o t
+
˙
Z t ot
indicates that the
process enables a more cost-ef fecti ve pro vision of e x er gy of product compared with the
SMR or DMR process. Both, the capital in vestment and the cost of the irre versibilities
are essentially lo wer than for the other base processes with similar methanol capacity .
Accordingly , the relati v e cost dif ference r t o t only amounts to 204.5%. A special feature
of this process is the high a v erage cost of the ex er gy streams which are due to the supply
of costly oxygen to the A TR.
The air compressor CM-02, the comb ustion unit R-09, the condenser E-09 and the
A TR R-03 ha v e the highest cost importance within the system. As for the other processes,
the air compressor e xhibits the highest
˙
Z k
among all components. The lar ge
f
v alue
of 86.9% clearly suggests a reduction of
˙
Z k
by accepting lar ger irre v ersibilities. K ey
design v ariables for a reduction of the in v estment cost include the pressure ratio and the
isentropic ef ficienc y . Since the of fgases lea v e the HRSG at 1.9 bar and the isentropic
ef ficienc y is high, significant sa vings could be achie v ed by a reduction of both parameters.
In contrast to all other processes, the comb ustion unit R-09 has a lo wer cost importance
for the o verall system due to smaller capacity . A part of the required process heat is
internally pro vided by the partial oxidation of methane in the autothermal reformer R-03,
compensating the heat supply of the comb ustion unit. Howe ver , R-09 has the second
highest v alue of the sum
˙
C D , k
+
˙
Z k
. Similar to the comb ustion units in the other systems,
the cost rate associated with the e x er gy destruction represents the major cost source (lo w
f
v alue). A reduction of the e xcess air not only might help to reduce cost rate associated
with the irre v ersibilities in R-09 b ut also the capital in vestment for CM-02. In regard
to the A TR, being ranked to position number 4, the irre v ersibilities also constitute the
major cost source as indicated by the lo w
f
v alue. The decision v ariables include the S/C
ratio and the O/C ratio. By a reduction of the O/C ratio, the ex er gy destruction due to
highly irre v ersible comb ustion reactions could be reduced. Ho we ver , parameter changes
should be conducted carefully as the y strongly influence the ex er gy destruction of other
components.
The steam c ycle condenser E-09 and the crude product cooler are also rated among
the components with the highest cost importance. Both heat exchangers serv e to remo v e
lo w temperature heat from the synthesis unit and the entire process. The respectiv e
v alue of the e x er goeconomic factor is lo w , due to a large cost rate associated with the
irre v ersibilities. The e x er gy destruction rate is in a moderate range, so that the costs are
rather caused by the high a v erage cost at which e xer gy is supplied to the components.
An ef fecti ve measure to decrease
˙
C D , k
is the reduction of the R TFR in the synthesis unit.
121

Chapter 5 Results and Discussion
This could be achie v ed by an increase of the synthesis unit pressure at the e xpense of a
higher in vestment for the syng as compression unit.
The refining column C-02 is rated position number 7, whereby the lo w
f
factor
indicates that the irre v ersibilities, related to the lar ge heat transfer rate in the reboiler and
in the condenser , are crucial to the cost of the component. The ke y decision parameters
include the water content in the crude product and the reflux ratio of the reboiler . The
water content could only be reduced by ma jor parameter and design changes in the
reforming and the synthesis unit. A decrease of the reflux ratio would ef fecti vely help in
decreasing the e xer gy destruction but also increase the e x er gy losses associated with the
rejection of the bottom product to the en vironment.
The HP-EV A and the intermediate cooler E-05 in the syngas compression unit are
ranked to position number 8 and 9, both ha ving a lo w
f
v alue. The cost rate
˙
C D , k
is
less due to e xer gy destruction by heat transfer at lar ge temperature dif ference, but rather
caused by the high a v erage cost per unit of e x er gy of fuel. A decrease of the a v erage
temperature dif ference between the hot and the cold side (thus of the outlet temperature
of the process gas) by an increase of the heat transfer area would ef fecti vely reduce
e xer gy detruction b ut inevitably increase the cost of the synthesis unit recuperator E-10,
which is also listed in T able 5.7. Similar to the other components of the synthesis unit, the
lar ge v alue of ˙
Z k + ˙
C D , k results from high a v erage cost per unit of e x er gy of the process
gas.
W ith respect to the turbomachinery , the specific cost for the internally supplied elec-
tricty (or mechanical work) are in the same range as for the SMR process. Also for
this process, the turbines sho w une xpected lo w
f
v alues. The low-pressure turbine T -01
is among the components with the highest cost significance. Although the isentropic
ef ficienc y is already high, a further increase would lead to a reduction of the e x er gy
destruction. Further measures to decrease
˙
C D , k
concern an impro vement of the li ve steam
parameters to decrease the irre v ersibilities due to friction by an increase of the steam
quality within the last stages of the turbine.
122

5.5 Exer goeconomic Analysis
T able 5.8: Results obtained from the e x ergoeconomic analysis for a selection of components of the reference A TR process.
Comp. k ˙
E F , k ˙
E P , k ˙
E D , k ε k y D , k c F , k c P , k ˙
C D , k ˙
Z k ˙
C D , k + ˙
Z k f k r k
[MW] [MW] [MW] [%] [%] [$/GJ] [$/GJ] [$/h] [$/h] [$/h] [%] [%]
CM-02 69.41 61.92 7.48 89.2 0.7 23.45 45.14 631.70 4,203.71 4,835.41 86.9 92.5
R-09 308.90 175.74 133.16 56.9 11.6 4.32 8.87 2,070.85 809.19 2,880.04 28.1 105.4
E-09 8.05 0.00 8.05 - 0.7 80.63 - 2,336.82 50.62 2,387.44 2.1 -
R-03 1,032.23 955.89 76.34 92.6 6.7 7.08 7.76 1,944.82 398.23 2,343.05 17.0 9.6
E-08 15.52 4.59 10.93 29.6 1.0 40.08 136.71 1,577.24 20.11 1,597.35 1.3 241.1
E-03 52.37 39.54 12.83 75.5 1.1 30.23 40.07 1,396.58 3.90 1,400.47 0.3 32.6
C-02 28.48 7.98 20.50 28.0 1.8 15.36 63.53 1,133.46 251.12 1,384.58 18.1 313.6
HP-EV A 61.15 47.84 13.31 78.2 1.2 18.85 24.12 903.40 3.88 907.28 0.4 28.0
E-05 14.34 8.07 6.27 56.3 0.6 32.91 58.89 742.43 12.14 754.57 1.6 78.9
E-06 27.60 23.53 4.07 85.3 0.4 43.15 51.94 632.38 112.49 744.87 15.1 20.4
CM-04 10.09 9.12 0.98 90.3 0.1 23.56 45.65 82.98 642.11 725.08 88.6 93.8
T -01 76.43 69.31 7.12 90.7 0.6 20.41 23.18 523.22 168.84 692.05 24.4 13.6
E-02 25.33 19.03 6.30 75.1 0.6 30.23 40.30 685.86 4.20 690.06 0.6 33.3
HP-SH 48.26 38.50 9.76 79.8 0.9 18.85 23.70 662.27 9.83 672.10 1.5 25.7
CM-01 7.27 5.78 1.48 79.6 0.1 23.45 52.55 125.37 480.42 605.79 79.3 124.1
C-03 17.14 6.65 10.49 38.8 0.9 9.50 33.53 358.87 216.31 575.18 37.6 252.8
MP-EV A 51.80 43.96 7.84 84.9 0.7 18.85 22.34 532.19 20.66 552.86 3.7 18.5
T otal 1,144.30 651.15 453.2 56.9 39.6 3.79 11.24 * 6,183.46
10,416.54 16,600.00
62.8 204.5
* includes ˙
C L and ˙
C Di f f
123

Chapter 5 Results and Discussion
Re garding the DMR process, the results obtained from the e xer goeconomic analysis
are gi v en in T able 5.9. The v alue of the sum
˙
C D , t o t
+
˙
Z t ot
is of a similar magnitude as
for the SMR process. The relativ e a v erage cost dif ference
r t ot
, ho we ver , only amounts
to 284.3% in comparison to 338.5% in the SMR process due to lar ger ex er gy rate of
product. In re gard to the products, the relati v e cost dif ference for electricity and methanol
is 640.1% and 171.4%, respecti v ely . The cost rates associated with the capital in v estment
(and O&M costs) and the e xer gy destruction ha ve a similar cost impact on the o v erall
process, as sho wn by an e x er goeconomic factor of f t ot = 49.8%.
In re gard to the cost distrib ution and the cost importance of the components, similar
characteristics as for the SMR process can be identified. Also for the DMR process, the
comb ustion unit R-05, the air compressors CM-02, and the condenser E-10 are the most
important components from a thermoeconomic point of vie w . Due to the high process
heat demand, these units are designed for a large capacity . The comb ustion unit has
the highest v alue of the sum
˙
C D , k
+
˙
Z k
and sho ws a typical lo w
f
v alue of 25%. The
cost rate
˙
C D , k
might be decreased by lo wering the air -to-fuel ratio and thus the e xcess
air for the comb ustion. Howe ver , the mass flo w rate of the e xcess air is determined by
the outlet temperature and therefore by the operation of the DMR reactor . A lo wer air
mass flo w rate w ould also result in a reduction of the capacity of the air compressor
CM-02, which is ranked to the second position among the components with the highest
cost significance. The high v alue of
f
sho ws, that the capital in vestment and the O&M
costs are the main source of cost. A reduction of the isentropic ef ficienc y would decrease
˙
Z k
at the e xpense of a lo wer e x er getic ef ficienc y . Furthermore, the pressure ratio might
be slightly decreased for a design impro v ement since the pressure of the of fg ases is still
high with 1.6 bar . In particular for the DMR process, the circulating mass flo w rate in
the synthesis unit is high. The remo v al of a lar ge amount of lo w temperature heat from
the synthesis unit is responsible for the lar ge irre v ersibilities within the crude product
cooler E-09 and the condenser E-10. Both components ha v e a lo w v alue of
f
, suggesting
an increase in the e xer getic ef ficiency . A reduction of the R TFR by an increase of the
synthesis unit pressure e xhibits the most ef fecti ve measure to decrease the irre v ersibilities
in these heat e xchangers. Additionally , the heat transfer area can be increased to reduce
the a v erage temperature dif ference of heat transfer and thus the component’ s e xer gy
destruction. The lar ge v alue of the relati v e cost dif ference for E-09 is due to the high
spcific cost of the circulating water between the condenser and the crude product cooler .
Another component with a high cost significance refers to the water -cooled high
temperature syngas cooler E-05. According to the lo w v alue of
f
, a reduction of
˙
C D , k
should be achie v ed by decreasing the a v erage temperature dif ference (up to 500
◦
C) of
124

5.5 Exer goeconomic Analysis
heat transfer . An intervention in the operation of the reformer could lead to a reduction of
the temperature of the process gas on the hot side. On the other hand, the syngas cooler
could be e xposed to a higher temperature le v el on the steam side, which w ould require a
rearrangement of the components in the steam c ycle.
Considering the v alue of the sum
˙
C D , k
+
˙
Z k
in T able 5.9, all other components (do wn
from position 6) ob viously ha v e a lo wer cost importance. In regard to the high-pressure
economizer HP-ECO, the superheater HP-SH and the reheater LP-RH in the HRSG, the
irre v ersibilities play a major role for their cost significance. The high e x er gy destruction
is predominantly caused by lar ge mass flo w rates of the high-pressure steam and the
e xhaust gas. A reduction of the operation temperature of the DMR would ef fecti vely
reduce the irre v ersibilities due to heat transfer within the HP-SH and thus the associated
cost rate ˙
C D , k .
The lo w and medium-pressure turbines T -01 and T -02 are also listed in the T able 5.9,
being ranked to position 8 and 11. As for the steam turbines in the other processes, the
irre v ersibilities also constitute the major cost source for these components. A change of
the li v e steam parameters and a slight decrease in the pressure ratio in order to increase
the steam quality at the outlet of T -01 constitute ef fecti v e measures to reduce ˙
C D , k .
As for the other processes, the refining column C-02 is also among the components
with a highest cost impact due to a lar ge cost rate associated with the irre v ersibilities. A
decrease of the reflux ratio would potentially lo wer
˙
C D , k
at the e xpense of an increase of
the cost rate associated with ex er gy losses. Additionally , the temperature of the steam in
the reboiler could be lo wered to reduce the e x er gy destruction due to heat transfer .
125

Chapter 5 Results and Discussion
T able 5.9: Results obtained from the e x ergoeconomic analysis for a selection of components of the reference DMR process.
Comp. k ˙
E F , k ˙
E P , k ˙
E D , k ε k y D , k c F , k c P , k ˙
C D , k ˙
Z k ˙
C D , k + ˙
Z k f k r k
[MW] [MW] [MW] [%] [%] [$/GJ] [$/GJ] [$/h] [$/h] [$/h] [%] [%]
R-05 1,709.69 1,112.97 596.72 65.1 25.5 4.07 7.00 8,753.34 2,966.85
11,720.18
24.8 71.3
CM-02 214.62 190.72 23.90 88.9 1.0 21.74 35.21 1,870.65 7,372.72 9,243.36 79.5 61.1
E-10 58.84 - 36.94 - 1.6 53.06 - 7,057.01 149.36 7,206.37 0.0 -
E-09 65.16 16.36 48.80 25.1 2.1 31.90
128.81
5,603.31 106.26 5,709.58 1.8 303.8
E-05 59.81 28.00 31.82 46.8 1.4 45.73 98.26 5,237.65 56.90 5,294.55 1.1 114.8
HP-ECO 341.42 267.66 73.75 78.4 3.2 11.01 14.07 2,924.37 20.86 2,945.23 0.7 27.8
HP-SH 254.46 196.94 57.52 77.4 2.5 11.01 14.29 2,280.66 39.84 2,320.50 1.7 29.7
T -01 259.59 237.54 22.04 91.5 0.9 19.72 22.29 1,564.99 635.45 2,200.44 28.5 13.0
CM-04 49.75 44.39 5.37 89.2 0.2 21.74 33.38 420.16 1,438.90 1,859.06 77.1 52.8
E-06 24.26 10.80 13.46 44.5 0.6 36.40 82.70 1,763.79 36.05 1,799.83 2.0 127.2
T -02 232.60 218.88 13.72 94.1 0.6 19.51 21.49 963.45 595.69 1,559.15 37.8 10.1
LP-EV A 108.94 74.16 34.78 68.1 1.5 11.02 16.21 1,380.49 4.82 1,385.31 0.3 47.1
LP-RH 53.60 27.54 26.06 51.4 1.1 11.01 21.86 1,033.16 42.50 1,075.66 3.9 98.4
C-02 27.12 6.28 20.84 23.2 0.9 10.23 57.16 767.67 293.21 1,060.88 27.4 457.1
R-03 68.29 65.31 2.98 95.6 0.1 81.32 85.06 873.37 4.59 877.96 0.5 4.6
CM-05 39.46 34.85 4.61 88.3 0.2 21.74 28.27 360.88 457.83 818.71 55.5 29.7
RECUP 69.15 53.64 15.50 77.6 0.7 11.01 15.22 614.64 197.13 811.77 23.9 37.9
T otal 2,336.87 929.13 1,245.47 39.8 53.3 3.68
14.14
*
16,487.07 16,350.55 32,837.63
49.8 284.2
* includes ˙
C L and ˙
C Di f f
126

5.5 Exer goeconomic Analysis
T able 5.10 presents the results obtained from the e xer goeconomic analysis of the CMR
process. F or a similar methanol capacity , the sum of the cost rates
˙
C D , t o t
+
˙
Z t ot
is lo wer
than for the SMR and DMR process. This reduction is mainly caused by a lo wer po wer
capacity of the system (see T able 5.1). According to the product capacities ( with a
focus on methanol), the av erage relati ve cost dif ference
r t ot
is comparati v ely lo w . A
special feature of this process is the inte grated Selexol
R

unit which causes relati v ely
high a v erage cost of the syngas by rec ycling
CO 2
of high specific cost to the reforming
unit. This in turn causes a sharp increase of the cost rate associated with the inef ficiencies
of the components in the synthesis unit. The combustion of uncon v erted and costly
synthesis gas (stream 10 in Fig. 4.11) results in high av erage cost per unit e x er gy of
the e xhaust gas and finally in high specific cost of the steam and electricity . Thus, in
comparison to other processes, electricity is pro vided internally to the consumers at high
a v erage cost
c F , k
of 27.0 US$/GJ. F or the ov erall process,
˙
Z t ot
has a lar ger impact than
the cost rate associated with the irre v ersibilities ˙
C D , t o t .
As with the other processes, the comb ustion unit R-05 and the air compressor CM-
02 also ha v e the highest cost importance for the CMR process. In comparison to the
endothermic DMR and SMR process, the demand for process heat is lo w , thus reducing
the size of these components. The most important design v ariables and their impact on
the components’ costs ha v e already been e xplained in detail for the other processes.
Interestingly , the refining column C-02 is rated to position number three. The steam
supplied to the reboiler of the column has lar ge a v erage cost per unit of e x er gy , causing a
high v alue of
˙
C D , k
. A reduction of the irre v ersibilities due to a decrease of the reflux ratio
would cause an unacceptable lo w product purity and high ex er gy losses. In this case a
third column would be required to guarantee AA methanol product quality . Furthermore,
the steam temperature on the hot side of the reboiler could be reduced to lo wer the e x er gy
destruction related to heat transfer .
The crude product cooler E-11 in the CMR process is comparati v ely small, since the
circulating mass flo w is lo w due to a high con v ersion rate in the synthesis unit. Ho we v er ,
the component is ranked to position number four , ha ving irre v ersibilities as the major
cost source. The high cost rate
˙
C D , k
is caused by a combination of a lo w e x er getic
ef ficienc y and high a v erage cost per unit of e x er gy of fuel. An increase of the synthesis
pressure could reduce the circulating mass flo w rate (and the heat transfer rate) and thus
the inef ficiencies within E-11. In addition, further cost reduction might be achie ved by
e xposing the component to higher steam temperatures to reduce the high temperature
dif ference of heat transfer .
127

Chapter 5 Results and Discussion
The syngas cooler E-07 is rated to position number 6 and has a lo w v alue for the
v ariable
f
due to a high cost rate associated with the irre v ersibilities. Considering Fig.
4.12, it is highly recommended to relocate the heat e xchanger within the heat integration
system to reduce the e xer gy destruction due to heat transfer . T urning next to the combined
reformer R-03, the relati v ely lo w v alue of the factor
f
suggests a decrease of the cost rate
associated with the irre v ersibilities. A decrease of the heat transfer rate achie v ed through
a reduction of the inlet temperature of the comb ustion gases results in a lo wer con v ersion
rate of methane. Alternati vely , the S/C ratio could be increased and the
CO 2
/C ratio
decreased to reduce the heat demand of the reforming reactions.
The syngas compressor CM-01 has a high cost significance due to a lar ge cost rate
associated with the capital in vestment (
f k = 83 . 7%
). The component is designed for
a lar ge capacity due to high pressure ratio of 15 between the A GR and the reforming
unit. The cost rate
˙
Z k
could be reduced by increasing the operation pressure of the
combined reformer or by a reduction of the isentropic ef ficienc y of CM-01. A change of
the operation parameters of the reformer , ho we v er , should be treated with care.
The high-pressure e v aporator (HP-EV A) is also important from a thermoeconomic
point of vie w due to a high v alue of the sum
˙
C D , t o t
+
˙
Z t ot
. The e xer gy destruction
constitutes the major cost source and results from the heat transfer at a temperature
dif ference of up to 400
◦
C (see Fig. 4.12). An increase of the steam pressure may
decrease ˙
C D , k , b ut requires major design changes in the steam c ycle.
The syngas cooler E-05 is also listed in T able 5.10. The component serv es to cool the
syngas for the WGS unit while preheating simultaneously steam for the shift reaction.
The lo w v alue for the variable
f
results from a lo w e x er getic ef ficienc y of 50.1% in
conjunction with relati v ely high a v erage cost of the e x er gy of fuel. Comparing the v alues
of the sum
˙
C D , k
+
˙
Z k
in T able 5.10, the other components obviously ha v e a relati vely
lo w cost significance. The water gas shift reactor R-04 is operated with lo w con v ersion
rate, thus e xhibiting lo w irre versibilities. Ho we v er , the
f
factor sho ws that the cost
rate associated with the irre v ersibilities is dominating. As with the other processes, the
irre v ersibilities are also the major cost source for the turbines T -01 and T -02.
128

5.5 Exer goeconomic Analysis
T able 5.10: Results obtained from the e x er goeconomic analysis for a selection of components of the reference CMR process.
Comp. k ˙
E F , k ˙
E P , k ˙
E D , k ε k y D , k c F , k c P , k ˙
C D , k ˙
Z k ˙
C D , k + ˙
Z k f k r k
[MW] [MW] [MW] [%] [%] [$/GJ] [$/GJ] [$/h] [$/h] [$/h] [%] [%]
R-05 890.55 597.04 293.51 67.0 18.6 4.30 7.07 4,538.87 1,421.47 5,960.34 23.9 64.6
CM-02 103.31 93.10 10.21 90.1 0.7 27.02 40.65 992.93 3,572.59 4,565.52 78.3 50.4
C-02 33.17 10.24 22.93 30.9 1.5 35.62 124.04 2,940.05 320.98 3,261.02 9.8 248.3
E-11 18.72 6.81 11.92 36.6 0.8 72.04 199.23 3,090.80 25.84 3,116.64 0.8 176.6
E-12 17.32 - 17.32 0.00 1.1 39.6 - 2,466.63 63.99 2,530.63 2.5 -
E-08 46.27 22.51 23.75 48.7 1.5 34.58 71.58 2,956.59 42.27 2,998.86 1.4 107.0
R-03 219.33 178.70 40.62 81.5 11.8 15.82 2.76 1,728.77 840.33 2,569.10 32.7 33.8
CM-01 82.08 78.13 3.94 95.2 0.3 27.02 35.37 383.74 1,962.64 2,346.38 83.7 30.9
HP-EV A 179.95 126.22 53.73 70.1 3.4 11.82 16.88 2,286.64 11.30 2,297.94 0.5 42.8
E-05 35.78 17.91 17.87 50.1 1.1 28.57 57.11 1,837.60 2.98 1,840.58 0.2 99.9
T -01 139.02 128.50 10.52 92.4 0.7 25.25 28.10 956.14 365.61 1,321.75 27.7 11.3
HP-SH 111.59 81.62 29.97 73.1 1.9 11.82 16.22 1,275.37 16.44 1,291.81 1.27 37.19
CM-03 11.23 10.08 1.15 89.8 0.1 27.02 56.67 111.86 964.17 1,076.02 89.6 109.7
C-01 8.90 1.04 7.86 11.7 0.5 35.62 320.16 1,007.48 58.82 1,066.29 5.5 798.9
T -02 127.28 121.16 6.12 95.2 0.4 22.83 24.77 502.92 343.65 846.58 40.6 8.5
E-06 18.55 9.08 9.47 49.0 0.6 24.21 49.56 825.21 3.44 828.66 0.4 104.7
R-04 87.58 85.20 2.38 97.3 0.2 72.80 74.84 623.95 3.21 627.16 0.5 2.8
T otal
1,576.45
765.97 653.66 48.6 41.5 3.83 12.41 *
12,621.10 12,033.39
24,654.48 56.9 224.0
* includes ˙
C L and ˙
C Di f f
129

Chapter 5 Results and Discussion
The results obtained from the e x er goeconomic analyis of the SMR-A TR process are
presented in T able 5.11. The process is designed for a large capacity , thus requiring a
higher capital in vestment and increased O&M costs (
˙
Z t ot
) compared to other processes.
Due to relati v ely lo w e x er gy destruction (see T able 5.4), the capital in vestment and O&M
cost ha v e a dominating impact on the cost of the ov erall process. The e xer gy streams
ha v e lo w a v erage cost, since the reforming agent oxygen can be supplied cheaply due
to the ef fect of the economies of scale concerning the ASU. As a result, the internally
consumed electricity is of lo w cost which generally decreases the cost rate associated
with the e xer gy destruction of the electricity consuming units. In comparison with other
processes the relati v e cost increase r t o t is the lo west, ha ving a v alue of 154.0%
Considering the component le v el, the comb ustion unit R-07, the air compressor CM-02
and the A TR R-05 are the most important components from a thermoeconomic point of
vie w . T aking into account the plant capacity and the process heat demand, the size of
R-07 and CM-02 is relati v ely small. The heat of the ef fluent from the A TR is reco v ered
in the primary tub ular steam reformer R-03, thus lo wering the heat duty that needs to be
co vered by the comb ustion of fuel in R-07. Furthermore, the process benefits from a lo w
heat requirement of the SMR, since the con version must be limited for stoichiometric
reasons. Similar to the comb ustion unit, the autothermal reformer R-04 is characterized
by a lo w f v alue, sho wing a dominating impact of ˙
C D , k . The irre versibilities, caused by
the highly irre v ersible PO X reactions, might be reduced by adjusting the S/C and the
O/C ratio.
Also for this process, the condenser E-13 has a high cost significance due to the cost
related to the irre v ersibilities. Ho we v er , compared with the other processes, the steam
has lo w a v erage cost per unit e x er gy , resulting in a comparati vely lo w v alue of
˙
C D , k
+
˙
Z k
.
All high-pressure heat e xchangers of the HRSG are present in T able 5.11 and ha v e a lo w
e xer goeconomic factor . The large v alue of
˙
C D , k
results from lar ge heat transfer rates and
high temperature dif ferences between the process gas, the e xhaust g as and the steam (see
Fig. 4.13).
Ob viously , the lar ge steam flo ws also cause lar ge cost rates
˙
C D , k
in the medium and
lo w-pressure turbine, which are rated to position ten and ele ven. Interestingly , the capital
in vestment and O&M costs play a more important compared to the turbines in the other
processes. The refining column C-02 also has a high cost importance in this process
gi v en by cost rate
˙
C D , k
. Although a three column distillation system is applied, a lar ge
heat transfer rate is gi v en for the reboiler of column C-02. A decrease of the reflux ratio
might reduce the v alue of the sum
˙
C D , k
+
˙
Z k
for C-02. Ho we v er , if the e xer gy losses
130

5.5 Exer goeconomic Analysis
should be maintained at the same le v el, a more intensi v e operation of the atmospheric
column C-03 would be required, thus shifting the irre versibilities do wnstream.
Due to lar ge syngas mass flo w rates, a v ariety of components within the syngas track
appears among the units with the highest cost importance. These include the main
rec ycle compressor CM-09, the syngas compressors CM-05 and CM-06 as well as the
recuperator of the shift unit E-06. Common to all these compressors is the dominant cost
influence of ˙
Z D , k .
The results obtained from the ex er goeconomic analysis of the SMR-DMR process
are presented in T able 5.12. The process has the highest methanol and electric capac-
ity among all analysed processes. Accordingly , the v alue of the sum
˙
C D , t o t
+
˙
Z t ot
is
much higher than for the other processes. The relati v e cost dif ference of the o v erall
system amounts to 229.3% which is lo wer than for the DMR and SMR process b ut
higher compared to the SMR-A TR, the A TR, and the CMR process. As for the other
systems, the relati v e cost dif ference for electricity is much higher than for methanol. The
corresponding v alues for r el ec and r M eoh amount to 777.7% and 150.2%, respecti v ely .
The comb ustion unit R-05, the air compressors CM-03 and CM-04, the high-pressure
economizer HP-ECO, and the lo w-pressure turbine T -02 are the most important compo-
nents from a thermoeconomic perspecti v e. Comparing the v alues of the sum
˙
C D , k
+
˙
Z k
,
an outstanding cost importance of the comb ustion unit is re v ealed. The component is
designed for huge capacity due to the high process heat demand of both endothermic
reforming units R-03 and R-04. The relati v ely lo w
f
v alue of 25.0% suggests an increase
of the e xer getic efficienc y at the e xpense of higher in vestment costs. The irre v ersibilities
are mainly caused by the chemical reactions, the heat transfer and the mixture of streams
at dif ferent temperature. A decrease of
˙
C D , k
can be achie v ed by lo wering the unit cost of
the e xer gy of fuel
c F , k
through an increase of the share of natural gas in the total amount
of fuel. In general, the rec ycled syngas from the pur ge (stream 16) has higher specific
cost than natural gas. Furthermore, actions can be taken to increase ε k . These measures
include a preheating of reactants and the reduction of the e xcess air for comb ustion.
The latter measure would simultaneously reduce the in vestment cost and therefore the
cost importance of the air compressors, which are rank ed to position two and three. In
addition, a further decrease of
˙
Z k
of the most capital-intensi v e units could be achie v ed by
reducing the isentropic ef ficienc y η s .
131

Chapter 5 Results and Discussion
T able 5.11: Results obtained from the e x er goeconomic analysis for a selection of components of the reference SMR-A TR process.
Comp. k ˙
E F , k ˙
E P , k ˙
E D , k ε k y D , k c F , k c P , k ˙
C D , k ˙
Z k ˙
C D , k + ˙
Z k f k r k
[MW] [MW] [MW] [%] [%] [$/GJ] [$/GJ] [$/h] [$/h] [$/h] [%] [%]
R-07
1,044.63
686.59 358.04 65.7 12.9 5.13 8.70 6,610.10 2,208.43 8,818.53 25.0 69.6
CM-02 124.45 111.04 13.42 89.2 0.5 16.46 30.43 794.83 4,793.02 5,587.86 85.8 85.0
R-04
2,087.82 1,941.35
146.48 93.0 5.3 5.10 5.63 2,687.60 1,053.41 3,741.01 28.2 10.5
HP-SH 220.19 161.60 58.60 73.4 2.1 11.29 15.43 2,382.59 24.05 2,406.64 1.0 36.6
E-13 37.93 - 37.93 - 1.4 13.44 - 1,834.77 79.07 1,913.84 4.1 -
HP-ECO 176.76 143.93 32.83 81.4 1.2 11.29 14.28 1,334.98 214.23 1,549.21 13.8 26.5
HP-EV A 139.51 103.87 35.65 74.5 1.3 11.29 15.18 1,449.48 4.81 1,454.29 0.3 34.4
C-02 52.28 16.30 35.98 31.2 1.3 7.12 31.33 922.77 497.75 1,420.52 35.0 339.8
CM-09 12.85 11.48 1.37 89.3 0.1 16.62 45.79 82.19 1,123.25 1,205.44 93.2 175.5
T -02 196.59 184.77 11.82 94.0 0.4 14.58 16.31 620.42 530.62 1,151.04 46.1 11.9
T -03 161.73 148.45 13.28 91.8 0.5 14.58 16.70 696.95 435.75 1,132.70 38.5 14.5
CM-01 12.60 11.43 1.18 90.7 0.00 16.46 43.85 69.70 1,057.18 1,126.88 93.8 166.5
CM-03 12.97 11.27 1.69 87.0 0.1 16.78 41.23 102.22 890.13 992.36 89.7 145.7
E-06 175.18 146.33 28.85 83.5 1.0 5.25 7.10 545.29 426.37 971.66 43.9 35.1
CM-05 26.03 24.31 1.73 93.4 0.1 16.62 27.19 103.23 821.66 924.89 88.8 63.6
CM-06 52.38 48.86 3.52 93.3 0.1 16.62 21.71 210.72 684.72 895.44 76.5 30.6
R-02 115.92 113.20 2.72 97.7 0.1 83.07 85.08 812.94 4.81 817.74 0.6 2.4
T otal
2,772.20 1,587.02
1,083.34
57.3 39.1 3.83 9.73 *
14,924.19 19,600.33
34,524.52 56.8 154.0
* includes ˙
C L and ˙
C Di f f
132

5.5 Exer goeconomic Analysis
Due to the high mass flo w rate of the e xhaust gas, v ery lar ge amounts of heat need to be
reco vered as steam in the HRSG in order to achie v e a high o v erall ef ficienc y . According
to the lar ge heat transfer rates, a v ariety of heat e xchangers (HP-EV A, LP-EV A, HP-SH,
LP-SH and RH) has a high cost significance caused by the cost rate associated with their
irre v ersibilities. For a deeper understanding of the inef ficiencies, Fig. 4.15 should be
considered. A reduction in the temperature dif ference between the hot and cold side
would reduce the e x er gy destruction within the HRSG and thus the cost rate
˙
C D , k
of
the heat e xchangers. The temperature profiles in Fig. 4.15 sho w that a third pressure
le v el or a pressure increase of the HP le v el could be an ef fecti ve measure to reduce the
irre v ersibilities. Furthermore, a temperature drop of the exhaust gas (
T 10
) would also
reduce
˙
C D , k
. Ho we ver , this would result in a high methane slip from the reformer and
increase the cost of do wnstream components.
Despite the high e xer getic efficienc y of the high- and lo w-pressure turbines T -01 and T -
02, the irre v ersibilities e xhibit the major cost source. The lo w-pressure turbomachinery is
designed for a lar ge po wer output of 530 MW , resulting in relati vely lar ge irrev ersibilities.
A change of the li v e steam parameters might decrease the e xer gy destruction caused by
friction of water droplets in the last e xpansion stages (impro v ement of steam quality).
The e xtraction column C-01 and the refining column C-02 are also among the components
with the highest cost significance. Large process streams ha ve to be processed at lo w
e xer getic ef ficiency , which causes a predominating cost rate of the irre v ersibilities (
f k
=
10 - 20%). Ef fecti ve measures for a reduction of the irre versibilities ha v e been introduced
before.
133

Chapter 5 Results and Discussion
T able 5.12: Results obtained from the e x er goeconomic analysis for a selection of components of the reference SMR-DMR process.
Comp. k ˙
E F , k ˙
E P , k ˙
E D , k ε k y D , k c F , k c P , k ˙
C D , k ˙
Z k ˙
C D , k + ˙
Z k f k r k
[MW] [MW] [MW] [%] [%] [$/GJ] [$/GJ] [$/h] [$/h] [$/h] [%] [%]
R-05
2,429.67 1,626.51
803.16 66.9 21.8 5.02 8.33
14,527.56
4,847.73
19,375.29
25.0 65.9
CM-04 207.25 186.20 21.05 89.8 0.6 26.20 39.19 1,985.32 6,717.43 8,702.76 77.2 49.6
CM-03 209.33 188.15 21.19 89.9 0.6 26.20 37.97 1,998.46 5,972.62 7,971.08 74.9 44.9
HP-EV A 490.40 358.98 131.42 73.2 3.6 14.32 19.58 6,773.87 28.30 6,802.17 0.4 36.8
T -02 579.51 535.75 43.76 92.5 1.2 23.62 26.42 3,721.07 1,671.93 5,393.00 31.0 11.8
LP-EV A 268.33 180.95 87.38 67.4 2.4 14.32 21.29 4,503.94 39.75 4,543.69 0.9 48.7
E-11 47.82 - 47.82 - 1.3 23.43 - 4,066.85 195.40 4,262.25 3.0 -
E-08 127.58 81.52 46.06 63.9 1.3 23.84 37.39 3,952.65 24.00 3,976.65 0.60 56.8
C-02 61.19 18.45 42.74 30.2 1.2 18.35 69.83 2,822.74 597.03 3,419.77 17.5 280.6
HP-SH 269.93 209.10 60.83 77.5 1.7 14.32 18.56 3,135.39 55.64 3,191.03 1.7 29.6
LP-SH 89.57 41.90 47.67 46.8 1.3 14.32 31.15 2,457.18 81.41 2,538.58 3.2 117.5
E-11 35.01 16.99 18.02 48.5 0.5 33.61 74.73 2,180.66 333.43 2,514.09 13.3 122.3
RH 68.22 27.28 40.94 40.0 1.1 14.32 36.48 2,110.23 66.04 2,176.27 3.0 154.8
E-05 54.28 22.88 31.40 42.2 0.9 17.89 42.78 2,022.80 27.16 2,049.96 1.3 139.1
T -01 201.68 190.05 11.63 94.2 0.3 23.25 25.65 973.26 663.85 1,637.11 40.6 10.3
E-10 46.28 39.42 6.86 85.2 0.2 33.61 41.73 829.55 322.49 1,152.04 28.0 24.2
C-01 16.91 2.75 14.16 16.3 0.4 18.35
123.85
935.42 109.26 1,044.68 10.5 575.0
T otal
3,689.79 1,707.75 1,735.50
46.3 47.0 3.79
12.48
*
23,678.22 26,199.11 49,877.33
52.5 229.3
* includes ˙
C L and ˙
C Di f f
134

5.5 Exer goeconomic Analysis
The calculated le v elized cost of electricity (LCOE) and methanol (LCOM) are consis-
tent with the results of the economic sensiti vity analysis in Section 5.4. Ho wev er , the
e xact breakdo wn of costs by allocation of monetary v alues to each unit of ex er gy of a
stream allo ws for adv anced and alternativ e conclusions re garding the economic use of the
processes. Figs. 5.13 and 5.14 present the le velized cost of methanol and the le v elized
cost of electricity in dependence of the specific fuel cost
f c
. The linear slope illustrates
the sensiti vity of the cost of the respecti ve product to a change of the fuel cost.
Considering the results in Fig. 5.13, the SMR process obviously cannot produce
methanol at competiti v e cost. Crucial is the supply of lar ge quantities of e xpensi ve high
temperature steam as a reforming agent, which increases the a v erage cost per unit ex er gy
of the syngas. Furthermore, this high temperature steam therefore cannot be used for
electricity generation. In addition, the process is not as highly inte grated into the steam
c ycle as the other synthesis routes, resulting in a higher cost rate associated with the
e xer gy losses. The cost of the losses are assigned to the products relati v e to their ex er gy
rate. A third reason refers to the direct hydrogenation of
CO 2
in the methanol synthesis.
The inte gration results in an excessi ve w ater production by the re v erse w ater gas shift
reaction (see reaction Eq. 2.27), which increases the capacity of the components in the
synthesis loop and the distillation unit. Do wnstream a lar ge amount of heat is pro vided
by costly steam within the reboilers, thus additionally increasing the cost of the main
product.
In re gard to the other processes, the lev elized methanol cost are in a similar range.
Ob viously , the SMR-DMR and the SMR-A TR process profit from the economies of scale
and the corresponding lo w fix ed cost. The SMR-A TR process can produce methanol at the
lo west le velized cost. As the specific fuel cost increase, the cost adv antage of producing
methanol increases. It should be noted that
CO 2
for dry reforming and hydrogenation
is assumed to be free of char ge (concerning the SMR, DMR, CMR, and SMR-DMR
process). Under these terms, the SMR-DMR process can generate methanol at the second
lo west cost. If costs are assigned to the incoming
CO 2
stream, the le v elized methanol
cost will be higher . For fuel cost of 8 US$/GJ the A TR and CMR process reach break
e v en with the DMR process. The product related lo w fix ed cost of the DMR process are
crucial for the cost adv antage at lo w fuel cost. On the other hand, the increased product
cost for the A TR at specific fuel cost of 0 US$/GJ can be attributed to the costly supply
of oxygen. At a high natural gas price abo v e 8 $/GJ, the fuel cost ha v e a higher influence
on the LCOM of the DMR (compared with the A TR and CMR process) due to high
CH 4
-intensity . The methanol product of the CMR and the A TR process sho ws a similar
price sensiti vity to changes in the fuel cost.
135

Chapter 5 Results and Discussion
024 6 8 10 12
0
200
400
600
800
1 . 000
Specific fuel cost, f c [US$/GJ]
Le v elized cost of methanol, LCOM [US$/mt]
SMR A TR
DMR CMR
SMR-A TR SMR-DMR
Figur e 5.13: Le v elized cost of methanol (LCOM) as a function of the fuel cost.
The le v elized cost of electricity (LCOE) directly depend on the a v erage cost at which
e xer gy is supplied to the steam c ycle. Thus, the LCOE depend on the a v erage cost of
the e xhaust gas from the comb ustion unit and on the specific cost of the process gas.
Depending on the system design, the combustion fuel contains a dif ferent proportion
of syngas which is determined by the pur ge ratio in the synthesis unit. The syngas is
of higher a v erage cost per unit ex er gy than the supplied natural gas. Consequently , the
LCOE of a process increases with an increasing supply of pur ge gas to the comb ustion
unit. Considering Fig. 5.14, a clear trend becomes apparent. For the processes with
endothermic reforming technology , the LCOE sho w a high sensiti vity to changes in the
specific fuel cost, while it is lo w for the processes with autothermal reforming.
The CMR process generates electricity at the highest le v elized cost among all processes.
The syngas rec ycled to the comb ustion unit is more e xpensi v e than for other processes
causing high specific cost of the e xhaust gases. The high price sensiti vity can be e xplained
by the high proportion of natural gas (66 mass-%) in the total amount of comb ustion fuel.
136

5.5 Exer goeconomic Analysis
A pecularity arises for the A TR process. Since exclusi vely syngas is used as a comb us-
tion fuel, the LCOE directly depend on the specific cost of the process gases. Therefore,
for lo w natural gas prices, the LCOE is already high due to e xpensi ve syngas generation
caused by the use of a costly reforming agent. F or the same reason a lo w price sensiti vity
results for the LCOE. The remaining processes (SMR, DMR, SMR-DMR, SMR-A TR)
ha v e lo wer LCOE at lo w specific fuel cost, since heat is rather provided by the comb us-
tion of natural gas instead of using the process gas. Accordingly the sensiti vity of the
LCOE to price changes is higher than for the A TR process. F or fuel cost of 5.3 US$/GJ
the A TR process reaches break e ven with the SMR and the SMR-DMR process. The
SMR-A TR process can clearly generate electricity at the lo west cost. The SMR-A TR
process sho ws a significant adv antage in terms of the le velized electricity generation cost.
Interestingly , the process has low LCOE at lo w specific fuel cost although costly oxygen
is used as a reforming agent. In comparison to the A TR process, the SMR-A TR process
ob viously can profit from the ef fect of the economies of scale of the ASU. In addition,
the LCOE sho w a lo w sensiti vity to changes in the specific cost due to lo w
CH 4
-intensity .
024 6 8 10 12
0
100
200
300
400
Specific fuel cost, f c [US$/GJ]
Le v elized cost of electricity , LCOE [US$/MWh]
SMR A TR
DMR CMR
SMR-A TR SMR-DMR
Figur e 5.14: Le v elized cost of electricity (LCOE) as a function of the fuel cost.
137

Chapter 5 Results and Discussion
W ithin the economic analysis, the cost rate of integrated
CO 2
streams was assumed to
be zero since it was considered as a w aste product (e.g., from a po wer plant with carbon
capture), which could cause additional cost if it would be emitted. Ho we ver , high-purity
CO 2
might constitute a resource of v alue since process equipment ine vitably would be
required for the supply . Consequently , the processes with
CO 2
inte gration could ha v e
been fa v ored in the cost estimation. A sensiti vity analysis is performed to in v estigate
the impact of possible
CO 2
cost on the le v elized product cost. Figure 5.15 presents the
le v elized cost of methanol as a function of the specific fuel cost, taking into account the
uncertainty of the cost of
CO 2
. Accordingly , Figure 5.16 sho ws the sensiti vity of the
le v elized cost of electricity . F or the analysis, the cost for
CO 2
were assumed to be in
the range of 0 - 100 US$/t. The graphs of the cases with maximum
CO 2
cost (SMR100,
DMR100 and SMR-DMR100) are sho wn in Figs. 5.15 and 5.16. Furthermore, the
graphs for the cases without
CO 2
cost are presented as a reference. Consequently , the
two related graphs (e.g. DMR and DMR100) limit the range of the cost influence of the
CO 2 on the methanol cost (e.g. ∆ LCOM SMR ) for a constant cost of natural gas.
024 6 8 10 12
0
200
400
600
800
1 . 000
Specific fuel cost, f c [US$/GJ]
Le v elized cost of methanol, LCOM [US$/mt]
SMR A TR
DMR CMR
SMR-A TR SMR-DMR
SMR100 DMR100
CMR100 SMR-DMR100
∆ LCOM SMR-DMR
∆ LCOM SMR
∆ LCOM CMR
∆ LCOM DMR
Figur e 5.15: Le v elized cost of methanol as a function of the fuel cost under uncertainty of the
cost for CO 2 .
138

5.5 Exer goeconomic Analysis
The increase of the le v elized cost of electricity and methanol by monetization of
CO 2
is proportional to the amount of inte grated
CO 2
. Consequently , the LCOM and
LCOE sho w the highest sensiti vity to a change in the
CO 2
cost for the DMR process,
follo wed by the SMR-DMR process, the CMR process, and the SMR process. Under the
assumption of a monetization of
CO 2
, a dif ferent cost hierarchy results for the processes.
In re gard to the LCOM, the processes with
CO 2
inte gration already lose their cost
adv antage for lo w costs of
CO 2
. For maximum
CO 2
cost of 100 US$/t, the LCOM
for the DMR process is higher than for the SMR process at lo w fuel cost. T aking into
account the cost increase
∆ LCOM CMR
and
∆ LCOM SMR-DMR
, the LCOM of the CMR
and SMR-DMR process is still lo wer than for the SMR process. In comparison to the
A TR process, methanol is generated at higher le v elized cost in the CMR and in the
SMR-DMR process.
024 6 8 10 12
0
100
200
300
400
Specific fuel cost, f c [US$/GJ]
Le v elized cost of electricity , LCOE [US$/MWh]
SMR A TR
DMR CMR
SMR-A TR SMR-DMR
SMR100 DMR100
CMR100 SMR-DMR100
∆ LCOE DMR
∆ LCOE CMR
∆ LCOE SMR-DMR
∆ LCOE SMR
Figur e 5.16: Le v elized cost of electricty as a function of the fuel cost under uncertainty of the
cost for CO 2 .
139

Chapter 5 Results and Discussion
W ith respect to the sensiti vity of the LCOE in Fig. 5.16, similar conclusions can
be dra wn. The LCOE in particular increase for the DMR process. At lo w fuel cost of
up to 4 US$/GJ, the LCOE is higher for the DMR process than for the SMR process.
Furthermore, electricity generated by the A TR process is of lo wer le v elized cost compared
to the electricity from SMR-DMR process. The CMR process still has the highest LCOE,
although the
CO 2
costs only ha v e a small impact. In regard to the SMR process,
CO 2
cost ha v e a mar ginal influence.
5.6 Ad v anced Ex er getic Analysis
The results obtained from the ex er getic, the economic and the e x er goeconomic analysis
sho wed that the SMR-A TR process is the most attracti v e from a thermodynamic and
economic point of vie w . In order to allo w a better understanding of the component
interactions and the improv ement potential, an adv anced e x er getic analysis is applied
to the system configuration. The methodology of the adv anced e x er getic analysis was
introduced in Section 3.4 and a flo wsheet of the process w as presented in Figure 4.14.
The results of the con ventional and adv anced ex er getic analysis on a component le vel are
pro vided in T able F .1 in Appendix F.
Una v oidab le and A v oidab le Exer gy Destruction
An impro vement of the operating conditions of the studied components in the scope of
technical and economic limitations helps to decrease their e xer gy destruction rate. Doing
so, the una v oidable part of a component’ s e x er gy destruction
˙
E UN
D,k
quantifies the amount
of e xer gy destruction that cannot be further reduced. Ev en though the assumptions made
are subjecti v e, general trends for an improv ement can be identified. For the present
analysis, a generous approach with increased heat exchange areas and high isentropic
ef ficiencies w as applied. Reg arding the reactors, the changes mainly refer to a decreased
pressure drop and an increased heat transfer coef ficient
k
. The una v oidable e xer gy
destruction of the columns is estimated through consideration of the minimum reflux
ratio [149] and minimum temperature dif ference in the utility heat e xchangers. The
complete set of assumptions is presented in T able 5.13. The results of the modifications
for a selection of components with the highest ˙
E D,k are presented in Figure 5.17.
140

5.6 Adv anced Exer getic Analysis
T able 5.13: Assumptions for the determination of the una v oidable e x er gy destruction.
Comp. k ˙
E D , k ˙
E UN
D , k
R-01 ∆ p = 0.2 bar ∆ p = 0 bar
R-02 ∆ p = 0.2 bar ∆ p = 0 bar
R-03 l = 11m, n = 5000,
k = 0.5 kW/m 2 K
l = 12.1 m, n = 6300,
k = 0.6 kW/m 2 K
R-04 ∆ p = 2 bar ∆ p = 0 bar
R-05 ∆ p = 1 bar ∆ p = 0 bar
R-06 ∆ p = 1 bar ∆ p = 0 bar
R-07 ∆ p = 1 bar , λ = 2.42 ∆ p = 0 bar , λ = 1.00
R-08 ∆ p pg = 3 bar , ∆ p st = 1 bar ∆ p pg = 0 bar , ∆ p st = 0 bar
C-01 ∆ T con = 40K
,
∆ T re = 180 K
,
R C-01 = 0 . 6
∆ T con = 10 K
,
∆ T re = 10 K
,
R C-01 = 0 . 45
C-02 ∆ T con = 40K
,
∆ T re = 160 K
,
R C-01 = 0 . 6
∆ T con = 10 K
,
∆ T re = 10 K
,
R C-02 = 0 . 45
C-03 ∆ T con = 20K
,
∆ T re = 170 K
,
R C-03 = 0 . 8
∆ T con = 40 K
,
∆ T re = 10 K
,
R C-03 = 0 . 6
E-01 - E-13 ∆ T min = 10 - 500 K,
∆ p = 0.2 - 4 bar
∆ T min = 0 . 5K ∆ p= 0
ECO, EV A, SH ∆ T min = 10 - 500 K,
∆ p = 0.2 - 5 bar
∆ T min = 0 . 5K ∆ p= 0
T -01,T -03 η s = 0.92, η mec h = 0.99 η s = 0.95, η mech = 1
T -02 η s = 0.93, η mech = 0.99 η s = 0.95, η mech = 1
CM-01,CM-05-06 η s = 0.88, η mech = 0.99 η s = 0.92, η mech = 1
CM-02 η s = 0.85, η mech = 0.99 η s = 0.92, η mech = 1
CM-07 - CM-09 η s = 0.87, η mec h = 0.99 η s = 0.92, η mech = 1
P-01 - P-03 η s = 0.85, η mech = 0.99 η s = 0.89, η mech = 1
Ev en though the process is of high ex er getic ef ficency , the results obtained from the
con ventional e x er getic analysis in Section 5.3 and T able F .1 sho w that the simulated
process still has a high impro vement potential. In general, the larger the absolute v alue of
the irre v ersibilities within a component is, the higher its impro v ement priority must be.
141

Chapter 5 Results and Discussion
The furnace R-07 and the comb ustion zone of the autothermal reformer R-04 ha v e the
highest e xer gy destruction based on the results of the con ventional e x er getic analysis.
In general, it is reasonable to assume that e v ery reactor system is unique in terms
of operating conditions and technical features. In particular , the catalyst acti vity , the
selecti vity of a reaction and the operation mode are lik ely to be kept constant. Therefore,
a lo w impro v ement potential of se v eral reactors only results from slight changes in terms
of operating conditions. W ith respect to R-07 and R-04, a lower pressure drop w as
assumed, reducing the thermodynamic irre v ersibilities due to friction. Furthermore, the
mass flo w rate of air and thus the comb ustion ratio
λ
has been adjusted to stoichiometric
consumption of the comb ustion reactions. By application of this measures, a lar ge
impro vement potential for R-07 is gi ven, which is represented by the relati v ely high share
of a v oidable e x er gy destruction. In contrast, R-04 only sho ws a mar ginal potential for
impro vement. A relativ e high impro v ement potential is obtained for the SMR unit R-03,
although the component is not sho wn in Figure 5.17. An e xtension of the tube length
l
in conjunction with an increase of the heat transfer coef ficient
k
and the number of
tubes
n
results in a reduction of irre v ersibilities due to heat transfer at lo wer temperature
dif ference.
Se v eral heat e xchangers are rated among the components with the highest e x er gy
destruction (see T able 5.5). The unav oidable part
˙
E UN
D , k
results after a reduction of the
pressure drop and a minimization of the temperature dif ference, respecti v ely . Se v eral heat
e xchangers of the high pressure le v el in the HRSG ha v e a lar ge potential for impro vement.
In particular for the HP-SH and the HP-EV A, the major amount of irre v ersibilities can be
a v oided (70% and 87%). In re gard to the high-pressure economizer HP-ECO and the
recuperator E-06, the a v oidable e x er gy destruction also has a high share of 53% and 55%
in total
˙
E D , k
, respecti v ely . The improv ement potential of the intercooler E-04 is restricted,
since a lar ge dif ference in the hot and cold thermodynamic a v erage temperature is still
present after modification. The e x er gy destruction of the crude product gas cooler E-11 is
primarily caused by a lar ge heat transfer rate. Thus, the major part of the irre v ersibilities
can be a v oided by application of the modifications.
Considering Figure 5.17, the refining column C-02 and the extraction column C-01
e xhibit a lar ge improv ement potential. In regard to the refining column C-02 82% of
the component’ s total ex er gy destruction can be a v oided. For the e xtraction column this
v alues e ven amounts to 89%. The turbomachinery is characterized by high ex er getic
ef ficiencies and thus only plays a subordinate role in re gard to the a v oidable e x er gy
destruction. Ho we v er , a moderate improv ement potential arises, although the isentropic
ef ficiencies
η s
are high in the design case. Increasing the isentropic stage ef ficiencies of
142

5.6 Adv anced Exer getic Analysis
the air compressors CM-02 and CM-03, the absolute ex er gy destruction decreases by
45% and 46%, respecti v ely . Among the steam turbines, the low-pressure e xpander T -03
e xhibits the lar gest potential for improv ement.
In conclusion, the av oidable part of e x er gy destruction is relati v ely lar ge for the
majority of components, which indicates a considerable impro vement potential for the
o verall process. In particular , the furnace R-04, the distillation columns C-01 and C-02,
and the high-pressure heat e xchangers of the HRSG of fer a potential for impro v ement.
Endogenous and Exog enous Ex ergy Destruction
The component interactions are quantified by dif ferentiation in endogenous and e x-
ogenous e xer gy destruction. The corresponding v alues for
˙
E EN
D , k
and
˙
E EX
D , k
are pro vided
in T able F .1. Furthermore, the definitions of the ex er getic ef ficiencies required by the
calculation of the endogenous e xer gy destruction are sho wn in T able E.1.
T wo-thirds of the o v erall e x er gy destruction is endogenous with a v arying distrib ution
among the system components. Therefore, an impro vement strate gy should first consider
the internally caused inef ficiencies, before modifications in the system design are studied.
A v ariety of components has a significant lar ge e xogenous e x er gy destruction, ha ving a
share of 35 - 40% in the components total ex er gy destruction. This characteristic results
from a highly integrated system design with se veral rec ycle streams and highlights the
importance of component interactions within the process.
The furnace R-07 and the autothermal reformer R-04 are the most important com-
ponents e xhibiting a high endogenous ex er gy destruction due to highly irre versible
oxidation reactions taking place. The mode of operation of both components has a lar ge
impact on the e xer gy destruction within the other components. In particular , the heat
e xchangers in the HRSG (HP-ECO, HP-EV A, HP-SH) are af fected by the performance
of the furnace R-07 and therefore by the heat demand of the SMR in R-03. The exoge-
nous e xer gy destruction of the components in the syng as compression unit (CM-05 and
CM-06), the synthesis unit (E-09 - E-10, CM-07 - CM-09) and the purification section
(C-01 - C-03) is lar gely determined by the mode of operation of the reformer units R-03,
R-04 and R-05. The irrev ersibilities within these components is strongly af fected by
the rec ycle gas flo w and the composition of the crude product, which in turn depend
on the composition of the syngas feed. In addition, the synthesis reactor R-08 also has
a significant impact on the e xogenous ex er gy destruction of se veral components as its
operating conditions and the catalyst type af fect the con version rate and thus the rec ycle
143

Chapter 5 Results and Discussion
flo w in the synthesis loop. The results suggest, that a careful selection and operation of
the reformer and synthesis reactor technology contrib utes to a reduction of the e xer gy
destruction of other components.
Combined Splitting of Exer gy Destruction
Based on combined splittings of the a v oidable and una v oidable as well as the en-
dogenous and e xogenous parts of e xer gy destruction (see Eq. 3.29- Eg. 3.35), the most
promising components for an impro vement within the e xisting design and for possible
modifications in system layout are identified. The av oidable and una voidable endoge-
nous and e xogenous parts of e xer gy destruction are illustrated in Figure 5.17 for the
components with the highest irre v ersibilities.
Considering only modifications in the operating range (
˙
E A V ,EN
D
), the autothermal
reformer R-07, the refining column C-02 as well as the high-pressure heat exhangers
in the HRSG (HP-EV A and HP-SH) exhibit an independent impro v ement potential as
sho wn by a lar ge v alue for the endogenous ex er gy destruction. Isolated improv ement
attempts should also be applied to the syngas cooler E-04, the recuperator E-06 as well
as to the crude product cooler E-03. The improv ement potential of the autothermal
reformer is of minor significance. T aking into account design modifications (
˙
E A V ,EX
D
),
reductions in e xer gy destruction can particularly be achie ved for the HP-SH, the HP-EV A
and C-02 by changing their system location and their mode of operation. Considering
the limitations arising from the highly inte grated system design, suggested modifications
ha v e to be in vestig ated carefully . One approach would be an analysis of the binary
component interactions.
144

5.6 Adv anced Exer getic Analysis
Figur e 5.17: Results of splitting the ex er gy destruction into its una v oidable and a v oidable
endogenous and exogenous parts (UN EN, UN EX, A V EN, A V EX).
0123456 7 8 10 12 14
R-07
R-04
HP-SH
C-02
HP-EV A
HP-ECO
E-06
E-04
E-11
E-03
C-01
CM-02
Ex er gy destruction ratio y D [%]
UN EN UN EX
A V EN A V EX
...
139.82 61.89
108.37 / 47.96
92.41 52.58 0.94 / 0.53
11.04 / 6.86 / 25.09 / 15.60
4.02 / 2.55 / 17.99 / 11.42
2.64 / 1.68 / 19.16 / 12.17
9.40 / 5.99 / 10.65 / 6.79
7.97 / 5.10 / 9.62 / 6.15
8.93 / 5.74 / 5.18 / 3.33
3.49 / 2.25 / 8.32 / 5.37
5.59 / 3.62 / 3.01 / 1.95
0.89 / 0.58 / 7.30 / 4.73
4.53 / 2.94 / 3.61 / 2.34
145

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