Thermodynamic Analysis of Coal Fired

Power Generation Cycles with

Integrated Membrane Reactor

and CO2 Capture

Dissertation

zur Erlangung des Grades

Doktor­Ingenieur

der Fakultät für Maschinenbau

der Ruhr­Universität Bochum

von

Frank Sander

aus Wimbern / Wickede (Ruhr)

Bochum 2011

Dissertation eingereicht am: 24. März 2011

Tag der mündlichen Prüfung: 17. Juni 2011

Erster Referent: Prof. Dr.­Ing. Roland Span

Zweiter Referent: Prof. Dr.­Ing. Viktor Scherer

Danksagung III

Danksagung

Die vorliegende Arbeit entstand im Rahmen meiner Tätigkeit am Lehrstuhl für

Thermodynamik und Energietechnik an der Universität Paderborn bzw. am Lehrstuhl

für Thermodynamik an der Ruhr­Universität Bochum.

Mein besonderer Dank gilt Herrn Prof. Dr.­Ing. R. Span, der diese Arbeit

ermöglicht und sie stets gefördert hat. Die internationale Ausrichtung seines

Lehrstuhles machte die Mitarbeit an dem europäischen Forschungsprojekt „ENCAP“

erst möglich. Die Mitarbeit an diesem Projekt hat mir viel Freude bereitet und ich

konnte in vielerlei Hinsicht vom Arbeiten in einem internationalen Umfeld profitieren.

Herrn Spans persönliches Engagement und seine kontinuierliche Unterstützung trug

maßgeblich zum Erfolg dieser Arbeit bei. Darüber ist Herr Span im Laufe der Zeit zu

einem persönlichen Freund geworden und war in manch anderen Fragestellungen ein

guter Ratgeber.

Ich bedanke mich bei Herrn Prof. Dr.­Ing. V. Scherer für das Interesse an dieser

Arbeit und der Übernahme des Zweitgutachtens.

Ebenso danke ich meinen Kolleginnen und Kollegen für die freundliche

Zusammenarbeit an den beiden Lehrstühlen. Besonders bedanke ich mich bei Frau

Dr.­Ing. Mandy Gerber für das freundschaftliche Verhältnis während der gesamten

Zeit an beiden Lehrstühlen. Die vielen Diskussionen mit ihr und die angenehme

Büroatmosphäre haben mich in vielerlei Hinsicht unterstützt. Darüber hinaus bedanke

ich mich bei Herrn Dipl.­Ing. Stephan Kotthoff für fachfremde aber nichtsdestotrotz

hilfreichen Diskussionen über meine Arbeit. Herrn Dr. Robin Payne danke ich für das

sorgfältige Lesen des englischen Manuskriptes und für die wertvollen

Korrekturvorschläge.

Abschließend bedanke ich mich bei meiner Familie, meinen Eltern und meiner

Frau Selda und unseren Kindern, die mich während meines gesamten

Ausbildungsweges und der Bearbeitungszeit dieser Arbeit unterstützt haben.

Der größte Teil dieser Arbeit wurde durch das europäische Forschungsprojekt

ENCAP (Vertragsnummer: SES6­CT­2004­502666) im Rahmen des sechsten

Rahmenprogramms finanziert, wozu ich zu Dank verpflichtet bin.

Baden(CH), im November 2011 Frank Sander

IV Acknowledgements

Acknowledgements

This work is the result from my work at the Chair of Thermodynamics and Energy

Technologies at the University of Paderborn and the chair of thermodynamics at the

Ruhr­Universität Bochum.

In particular, I am indebted to Prof. Dr.­Ing. R. Span, who made this work possible

and who supported it. Prof. Span’s international orientation resulted in the opportunity

for me to contribute in the European research project “ENCAP”. The collaboration in

this project gave me a lot of pleasure and I was able to benefit from the work in such

an international environment in many ways. Prof. Span’s dedication and his

continuous personal support was important for the success of this work. Moreover,

Prof. Span has become a personal friend with a listening ear and good advice.

Special thanks to Prof. Dr.­Ing. V. Scherer for his kind interest in this work and for

acting as second reviewer.

Furthermore, I would like to thank all colleagues from both universities for the

friendly working atmosphere. In particular, I am grateful to Dr.­Ing. Mandy Gerber

who has become a close friend of mine over the last years. Mandy supported me in

many ways, not only with a lot of discussions, but also by creating a pleasant

atmosphere in our shared office. Thanks to Dipl.­Ing. Stephan Kotthoff also for not

only technical but helpful discussions about my work. I want to express my thanks to

Dr. Robin Payne for carefully reading the entire manuscript and for his helpful

suggestions for improving the English style.

Finally, I wish to thank my family, my parents and my wife Selda and our children,

for supporting me not only during the time of working and writing on this thesis, but

also during the period of education.

Most of this work was funded by the European research project ENCAP (contract

no. SES6­CT­2004­502666) as part of the sixth framework programme which is

gratefully acknowledged.

Baden(CH), November 2011 Frank Sander

Table of Contents V

Table of Contents

Nomenclature .................................................................................................... IX

Summary ........................................................................................................ XIII

1 Introduction ................................................................................................... 1

1.1 Motivation .............................................................................................. 1

1.2 The ENCAP Project ............................................................................... 5

1.3 Scope of the Work .................................................................................. 6

2 Carbon Dioxide Capture Technologies ...................................................... 9

2.1 Overview ................................................................................................ 9

2.2 Pre­combustion CO2 Capture ............................................................... 12

2.3 Oxyfuel Cycles (Integrated CO2 Capture)............................................ 13

2.4 Post­combustion CO2 Capture .............................................................. 14

2.5 Power Generation Cycles with Integrated Membrane Reactors .......... 16

3 Membranes for Gas Separation ................................................................ 17

3.1 Background Information ...................................................................... 17

3.2 Membranes for Oxygen and Hydrogen Separation .............................. 21

3.2.1 Oxygen Transport Membranes ............................................................21

3.2.2 Hydrogen­selective Membranes ..........................................................28

4 Analysed Power Generation Processes ..................................................... 33

4.1 Common Framework for Modelling of Power Generation Cycles ...... 33

4.2 Integrated Gasification Combined Cycles (IGCC) .............................. 36

4.2.1 IGCC process without CO2 Capture ....................................................36

4.2.2 IGCC process with Cryogenic Air Separation Unit and CO2

Capture .................................................................................................42

4.2.3 IGCC process with Integrated Oxygen Transport Membrane

(OTM) and CO2 Capture .....................................................................46

VI Table of Contents

4.2.4 IGCC process with Hydrogen­selective Membrane and CO2

Capture .................................................................................................51

4.3 Lignite Fired Boiler Cycles .................................................................. 54

4.3.1 Air Driven Boiler without CO2 capture ...............................................54

4.3.2 Oxyfuel Boiler with cryogenic ASU and CO2 capture ........................58

4.3.3 Oxyfuel Boiler with Integrated OTM Reactor and CO2 capture .........62

5 Modelling of Essential Sub-Processes ....................................................... 69

5.1 A Generic Cooled Gas Turbine ............................................................ 69

5.1.1 General Information ............................................................................69

5.1.2 Thermodynamic Basics of Gas Turbine Performance

Calculations .........................................................................................71

5.1.3 Modelling of film cooling ...................................................................74

5.2 Coal Gasification .................................................................................. 81

5.3 Sulphur Removal .................................................................................. 82

5.4 CO­Shift Reaction and CO2 Separation Process .................................. 83

5.5 CO2 Compression ................................................................................. 83

6 Modelling of Integrated Membrane Reactors ......................................... 85

6.1 Introduction to the Modelling ............................................................... 85

6.2 Pressure drop ........................................................................................ 87

6.3 Heat Transfer ........................................................................................ 91

6.4 Mass Transfer ....................................................................................... 93

6.5 Parametric Studies on the Membrane Reactors.................................. 100

6.5.1 Results for the Oxygen Transport Membrane (OTM) ...................... 100

6.5.2 Results for the Hydrogen­selective Membrane Reactor ................... 112

7 Simulation of the Analysed Power Generation Cycles ......................... 119

7.1 Simulation of the IGCC Cycles .......................................................... 119

7.1.1 Design Point Comparison of all Investigated Configurations .......... 119

7.1.2 Variation of the Operating Conditions of the OTM Reactor ............ 126

7.2 Simulation of the Oxyfuel Boiler Cycles ........................................... 127

Table of Contents VII

7.2.1 Design Point Benchmarking of all Investigated Configurations ...... 127

7.2.2 Variation of the Operating Conditions of the OTM Reactor ............ 131

7.2.3 Variation of the Heat Input to Gas Turbine and Boiler .................... 133

8 Conclusions and Recommendations for Further Work ........................ 135

8.1 Summary of the Main Results ............................................................ 135

8.2 Conclusions on the IGCC configurations ........................................... 136

8.2.1 General Conclusions on the Cycle Layout ....................................... 136

8.2.2 Conclusions on the OTM Reactor as part of the IGCC .................... 136

8.2.3 Conclusions on the hydrogen­selective Reactor as part of the

IGCC ................................................................................................. 137

8.3 Conclusions for LFB Cycles .............................................................. 138

8.3.1 General Conclusions on the Cycle Layout ....................................... 138

8.3.2 Conclusions on the OTM Reactor as part of the LFB cycle ............ 139

8.4 Recommendations for further work ................................................... 139

Appendix ......................................................................................................... 143

A.1 Gas compositions of the different IGCC configurations .................... 143

A.2 Properties of certain positions of the gas turbine process of the

IGCC configurations .......................................................................... 145

A.3 Reference cases from ENCAP SP 3 of the lignite fired boiler

process with and without CO2 capture .............................................. 146

A.4 Emitted and captured carbon dioxide of the investigated power

generation processes ........................................................................... 147

A.5 Variation of the feed temperature of the OTM reactor ...................... 149

A.6 Power and Efficiency Balances for IGCC cycles with Integrated

ASU .................................................................................................... 157

References ....................................................................................................... 159

Nomenclature IX

Nomenclature

Abbreviations

AIR Air fired cycle

AR4 IPCC Fourth assessment report

ASU Air Separation Unit

ATR Autothermal Reformer

AZEP Advanced Zero Emission Power

CLC Chemical Looping Combustion

CoE Cost of Electricity

DAI Dangerous Anthropogenic Interference

DMPEG Dimethyl ether of polyethylene glycol

ENCAP Enhanced Capture of CO2

EOR Enhanced Oil Recovery

FCG Flue Gas Cooling

GHG Greenhouse Gas

HP High Pressure

HRSG Heat Recovery Steam Generator

HTC Heat Transfer Coefficient

IAE International Energy Agency

IGCC Integrated Gasification Combined Cycle

IPCC Intergovernmental Panel on Climate Change

IP Intermediate Pressure

LFB Lignite Fired Boiler

LHV Lower Heating Value

LP Low Pressure

MDEA Methyl diethanolamine

MIEC Mixed Ionic and Electronic Conducting

OTM Oxygen Transport Membrane

OXY Oxyfuel

TAR Third Assessment Report

TAT Temperature after Turbine (turbine exit temperature)

TIT Turbine Inlet Temperature

UNFCC United Nations Framework Convention on Climate Change

WSC Water/Steam Cycle

X Table of Contents

Latin letters

A area m2

b factor 1

c factor m

c specific heat capacity J mol­1 K­1

C concentration of atomic hydrogen mol m­3

cm pre­factor mol s­1 m­1

d diameter m

D diffusion coefficient m2 s­1

E activation energy kJ mol­1

j oxygen flux mol s­1 m­2

k hydrogen permeability mol m­1 s­1 Pa­0.5

k surface oxygen exchange coefficient cm s­1

k overall heat transfer coefficient W m­2

K Sieverts constant mol m­3 Pa­0.5

K model parameter 1

l length m

m mass kg

Ma Mach number 1

n number (of) 1

n polytropic exponent 1

N hydrogen flux mol s­1 m­2

Nu Nusselt number 1

p partial pressure or absolute pressure Pa

Pe Peclet number 1

Pr Prandtl number 1

Q Heat W

R universal gas constant J mol­1 K­1

Re Reynolds number 1

s factor 1

T Temperature K

U circumference m

V volume m3

w flow velocity m s­1

W shaft power W

X membrane thickness m

X flow length m

Nomenclature XI

Greek letters

convective heat transfer coefficient W m­2 K­1

thickness m

difference 1

isentropic exponent 1

dynamic viscosity m2 s­1

efficiency 1

number pi 1

thermal conductivity W m­1 K­1

density kg m­3

friction factor 1

loss factor 1

Subscripts

0 constant

a outer

A activation

blade blade

c critical

C Compressor

Cb combustor

cool cooling

Ex exit

gen generator

GT gas turbine

H atomic hydrogen

H2 molecular hydrogen

hyd hydraulic

i inner

irrev irreversible

m membrane

mech mechanical

mem membrane

O2 molecular oxygen

p polytropic

p constant pressure

rev reversible

S Sieverts

s isentropic

T Turbine

tech shaft

th thermal

Superscripts

n pressure exponent o ideal gas

12 Table of Contents

Summary XIII

Summary

In this work two different types of coal fired power generation cycles are

thermodynamically analysed and benchmarked to each other. First, different

configurations of an integrated gasification combined cycle (IGGC) are investigated.

Second, different configurations of a lignite fired boiler (LFB) cycle are studied. For

both types of power generation cycles, appropriate configurations without CO2 capture

are analysed and are used as reference cycles for the appropriate power generation

cycles with CO2 capture. Furthermore, membrane reactors are integrated into these

power generation cycles in addition to CO2 capture. One goal of the present work is to

show the thermodynamic potential of these power generation cycles with CO2 capture

and integrated membrane reactor in comparison to the appropriate cycle without

membrane reactor. Another goal is to investigate how the operating conditions of the

membrane reactor induces the overall power generation cycles in terms of power

output and net efficiency.

For both types of power generation cycles, IGCC and LFB cycles, pure oxygen is

required as an oxidant for the gasification or the combustion process. Conventionally

the oxygen is supplied by a cryogenic air separation unit (ASU). The specific

expenditure of energy for producing the oxygen in a cryogenic ASU is high. This

circumstance is the reason for the idea of integrating an oxygen transport membrane

(OTM) into such a power generation cycle. The OTM reactor is integrated into the gas

turbine cycle and separates oxygen from preheated air and thus the ASU becomes

redundant. For the IGCC, additionally a configuration with integrated hydrogen­

selective membrane reactor is analysed to compare its thermodynamic potential to the

one of the OTM reactor. The arrangement of the OTM reactor is similar for both

analysed power generation cycles. The OTM reactor is located after the first

combustion chamber, so that the feed stream is preheated by the combustor to 900°C

before the stream enters the OTM reactor. In case of the IGCC configuration, the

retentate stream leaving the OTM reactor is further heated in a second combustor. In

case of the LFB cycle it should be emphasised that by integrating the OTM reactor

into the cycle, the lignite fired boiler cycle is ‘converted’ to a combination of a coal

fired boiler and a natural gas fired combined cycle power plant.

The IGCC without CO2 capture achieves a net efficiency of 45.1%. If CO2 capture

is applied to the IGCC, the net efficiency drops by 10.0%­pts. In case of the OTM

reactor it is found that the net efficiency is determined by the operating conditions of

the OTM reactor. In case of challenging operating conditions, meaning a large

XIV Summary

difference in total pressure across the membrane, the IGCC with integrated OTM

reactor achieves a net efficiency of 35.5%. This shows that the thermodynamic

potential of the IGCC with integrated OTM reactor is limited because the net

efficiency can only be slightly higher, by 0.4%­pts. If the difference in total pressure

has to be smaller to limit the mechanical load on the membrane material, the net

efficiency reduces to 31­33%. In this case the cycle becomes less attractive compared

to the IGCC with CO2 capture. The thermodynamic potential of the IGCC with

integrated hydrogen­selective membrane reactor is higher, obtaining a net efficiency of

36.7%.

The air fired LFB cycle without CO2 capture attains a net efficiency of 48.9%. If

oxygen is used instead of air and the CO2 is captured, the net efficiency drops by

9.8%­pts. If the OTM reactor – as part of a gas turbine – is integrated into the LFB

cycle, the net efficiency might increase to 43.4%. Same as for the IGCC, the operating

conditions determine the performance of the overall power generation cycle. In case of

conservative operating conditions (a low pressure difference across the membrane),

the net efficiency drops to 36­40%.

Apart from the impact on the overall power generation cycle, the operating

conditions of the OTM reactor determine the size of the membrane reactor. Parametric

studies on the OTM reactor are carried out to show how the operating conditions

influence the membrane surface area. For the IGCC with integrated OTM reactor and

the investigated range in pressure and mass flow rate of sweep stream, the membrane

surface area differs from 100,000 to 700,000 m2. An increased feed temperature would

reduce the overall membrane surface area because the permeation flux through the

membrane increases with higher temperatures. The difference in oxygen partial

pressure across the membrane is the driving force for the mass transport through the

membrane. Both parameters, mass flow rate and pressure of the sweep stream

determine the oxygen partial pressure on the sweep side of the OTM reactor. On the

one hand, it can be concluded that the pressure of the sweep stream has a stronger

impact on the membrane surface area than the mass flow rate. If the goal is to reduce

the membrane surface area as much as possible, then the pressure of the sweep stream

has to be as low as technically feasible. On the other hand it should be kept in mind,

that the requirement of a low sweep pressure leads to an increased mechanical load for

the membrane material. Therefore the OTM reactor needs to be designed mechanically

as a pressure vessel, which is capable to withstand the difference in total pressure.

1 In t roduct ion 1

1 Introduction

1.1 Motivation

The Intergovernmental Panel on Climate Change (IPCC) draws the conclusion in

the fourth assessment report (AR4) based on observational evidence from all

continents and most oceans that there is a correlation between greenhouse gas

emissions, temperature increase and regional climate changes. The AR4 states, “many

natural systems are being affected by regional climate changes, particularly

temperature increases. A global assessment of data since 1970 has shown it is likely

that anthropogenic warming has had a discernible influence on many physical and

biological systems.” [1] The AR4 should provide guidance to decision­makers for

identifying levels and rates of climate change that may be associated with ‘dangerous

anthropogenic interference’(DAI) with the climate system, according article 2 of the

United Nations Convention on Climate Change (UNFCCC) [2]. Ultimately, the

determination of DAI cannot be based on scientific arguments alone, but involves

other judgements informed by the state of scientific knowledge [1].

Furthermore, the AR4 summarises the correlation between concentrations of

greenhouse gases (GHG) and change in local temperature by means of analysis of ice

cores. Ice cores from different holes in the Antarctic down to a depth of more than

3,000 m are analysed and cover a time period of 650,000 years [5­8]. The air which is

trapped in those ice cores is extracted and its composition is determined. The GHG

concentration of carbon dioxide (CO2), methane (CH4) and nitrous oxide (N2O) is

shown in figure 1­1. Apart from the GHG concentrations, the variation in deuterium

(heavy hydrogen) D is given in figure 1­1 because D is a proxy for local

temperature [5]. The time period of 650,000 years covers four ice age cycles with short

warm periods (interglacials, illustrated by grey shaded bars in figure 1­1) and longer

cold periods (glacials). In the last four interglacial periods the concentration of CO2 in

the atmosphere was always less than 300 ppmv, whereas in the year 2000 the CO2

concentration has risen to 370 ppmv. In other words, all three GHG’s concentrations

analysed from the ice cores show the highest value of the last 650,000 years for all

three GHG’s (CH4: 1750 ppbv, N2O: 315 ppbv) [4]. For the same period it can be

concluded that the temperature level correlates temporally with the concentration of

CO2 in the atmosphere. This indicates that the connection between the cause of global

warming and the consequences are correctly understood. The cause of global warming

is the increased amount of global GHG emissions, which has grown due to human ac­

2 1 .1 Mot iva t ion

Figure 1-1: Variations of deuterium (δD) in antarctic ice, which is a proxy for local

temperature, and the atmospheric concentrations of the greenhouse gases

carbon dioxide (CO2), methane (CH4), and nitrous oxide (N2O) in air trapped

within the ice cores and from recent atmospheric measurements. Data cover

650,000 years and the shaded bands indicate current and previous interglacial

warm periods [4].

tivities since pre­industrial times, with an increase of 70% between 1970 and 2004 [9].

The increase in global temperature, as the final consequence of global warming, is

caused by changes in the atmospheric concentrations of GHG’s and aerosols, land

cover and solar radiation because they alter the energy balance of the climate system

and are drivers of climate change [4]. The third assessment report from 2001 shows

that the mean global surface temperature has risen by 0.6 K since 1850 (representing

the starting point of the industrialisation) [3]. The AR4 as subsequent report from 2007

updates the result of the increase in global surface temperature to 0.76 K [4].

The AR4 gives best estimates for a quantitative correlation between CO2

concentrations in the atmosphere and temperature increase in the future. In case the

CO2­equivalent concentration1 can be limited to 450 ppmv, the best estimate of the

increase in global surface temperature is 2.1 K; its is ‘likely’ (probability of occurrence

1 CO2­equivalent concentration is the concentration of CO2 that would cause the same amount of radiative

forcing as a given mixture of CO2 and other forcing components [4].

1 In t roduct ion 3

is > 66%) that the increase will be between 1.4 and 3.1 K [4]. During an ice age cycle

the change in global surface temperature is 5­6 K [8]. Comparing these two changes in

global surface temperature, it gives an impression for the order of magnitude. The

increase in global surface temperature may be between 33 and 50% compared to the

decrease of an ice age cycle.

As mentioned before the goal of the AR4 is to provide a scientific guidance to

support a common understanding and finally an agreement on limits for global

warming. Due to the fact that the climate change is a global problem all countries

worldwide are affected. Therefore any agreement on limits and possible mitigation

measures that is found, must be supported by as many countries as possible. A first

step was achieved at the latest United Nations Climate Change Conference, which

took place from November 29 to December 10, 2010, in Cancun, Mexico. The

conference encompassed the sixteenth Conference of the Parties (COP 16), at which

the decision was adopted, amongst others, that “deep cuts in global greenhouse gas

emissions are required according to science, with a view to reducing global GHG

emissions so as to hold the increase in global average temperature below 2°C above

pre­industrial levels”. [10] The outcome of the United Nations Climate Change

Conference is an indication that the political will is increasing to quantify limits of

global warming. However, the outcome of the latest United Nations Climate Change

Conference does not explicitly addresses how the limitation in temperature increase

should be achieved and by what amount each country/party should reduce its GHG

emissions. Such an agreement in form of a follow­up to the

Figure 1-2: Global anthropogenic GHG emissions in 2004 (left hand side). GHG emissions

by sector in 2004 (right hand side). 1) Excluding refineries, coke ovens etc.,

which are included in industry [9 – modified].

4 1 .1 Mot iva t ion

Kyoto Protocol [11] from 1997 – also part of the United Nations Convention on

Climate Change – is still outstanding.

The idea of avoiding GHG emissions lead consequently to the question of who the

largest contributors of the worldwide GHG emissions are. The working group III

contributing to the AR4 summarises the global GHG emissions in 2004 [9]. The global

anthropogenic GHG emissions are given in the pie chart on the left hand side in figure

1­2, whereas the GHG emissions are divided into sectors in the pie chart on the right

hand side of figure 1­2. More than the half of the global anthropogenic GHG

emissions, nearly 57%, are caused by burning fossil fuels. Dividing the GHG

emissions to the sectors where they are emitted, it can be seen that approximately a

quarter of the global anthropogenic GHG emissions are caused by energy supply. The

breakdown of the global GHG emissions by sectors indicates clearly that it is

worthwhile to capture CO2 from fossil fuelled power generation cycles because energy

supply is the largest share of GHG emissions of all sectors. The Internal Energy

Agency (IEA) gives an overview in their latest World Energy Outlook 2010 (WEO­

2010) [12] that about 81% of the world primary energy demand is provided by fossil

fuels. The share of fossil fuels on the world primary energy demand is further divided

into coal, oil and gas. One third of the energy supplied by fossil fuels is provided by

coal [12].

The IEA is an intergovernmental organisation that acts as energy policy advisor to

28 member countries. The goal of the IEA is to promote energy security to its member

countries. In this framework the IEA publishes the World Energy Outlook on a yearly

Figure 1-3: Share of yearly energy-related CO2 emissions savings by policy measures in the

“450 scenario” presented by the IEA in the WEO-2010 [12].

1 In t roduct ion 5

basis. In their latest report, WEO­2010, the IEA has updated projections of energy

demand, production, trade and investment, fuel by fuel and region by region to 2035.

Furthermore it includes a new scenario that anticipates future actions to meet the

commitments, which governments have made by COP 15 / COP 16, to tackle climate

change. The WEO­2010 presents the “450 Scenario”, which sets out an energy

pathway consistent with the 2°C goal through limitation of the GHG concentrations

gases in the atmosphere to a CO2­equivalent concentration of around 450 ppmv. The

development of energy­related CO2 emissions for both the current policies scenario

and the 450 scenario is shown in figure 1­3. Today the total energy­related CO2

emissions are around 30 GT/a. Continuing with current policies scenario, the CO2

emissions would increase to about 42 GT/a. In case of the 450 scenario the energy­

related CO2 emissions need to decrease to 21 GT/a. Furthermore, in this scenario

nearly 70% of the savings of CO2 emissions is achieved by energy efficiency and

utilisation of renewable energies. Carbon Capture and Storage (CCS) is assumed to

contribute approximately 20% to the overall savings on CO2 emissions. The 450

scenario presented in the WEO­2010 shows that CCS is one of the main contributors

for achieving a drastic reduction of CO2 emissions in the future decades.

1.2 The ENCAP Project

The major part of the work of this thesis is carried out in the framework of the

European research project ENCAP2 [13]. The overall budget of the ENCAP project is

€22.1 million and is funded by the European Commission with €10.5 million. ENCAP

is part of the sixth framework for research and technological development. The

duration of the ENCAP project is 60 months from 2004 to 2009. The main objective of

the ENCAP project is to develop new pre­combustion CO2 capture technologies and

processes for power generation based on fossil fuels (mainly hard coal, lignite and

natural gas), which can be integrated into actual sustainable energy systems in order to

reduce CO2 emissions. ENCAP aims at technologies that meet the target of at least

90% CO2 capture rate and 50% CO2­capture­cost reduction (compared to ‘state­of­the­

art’ levels in 2004) [13]. The consortium of the ENCAP project is a group of 33 legal

entities comprising 6 large European fossil fuel end users, 11 leading technology

providers, and 16 research technology & development providers. The following

corporations and organisations are involved in the ENCAP project: Energi E2, Dong

2 ENCAP is the abbreviation of “ENhanced CAPture of CO2”

6 1 .3 Scope of the Work

Energi, Alstom, DLR, Chalmers University, RWE Power, Siemens, ISTFA, NTNU,

PPC, Mitsui Babcock, TNO, University of Paderborn, Statoil, Linde, IFP University of

Stuttgart, Vattenfall, Air Liquide, SINTEF, University of Twente, Lurgi, University of

Ulster, BOC.

Figure 1-4: Structure of the European research project ENCAP.

The activities within the ENCAP project are structured in 6 sub­projects (SP). Each

SP is further divided into different work packages. The structure of the ENCAP project

is graphically illustrated in figure 1­4. The present thesis is a result of the work carried

out in the framework of the of the sub­project 6 (SP 6), which deals with “Novel pre­

combustion and oxyfuel capture concepts”, in the ENCAP project. The Chair of

Thermodynamics and Energy Technologies at the University of Paderborn, who has

been partner of the ENCAP was in total funded for three years by the ENCAP project.

1.3 Scope of the Work

The scope of this work is to thermodynamically analyse different power generation

cycles with and without CO2 capture. For showing the thermodynamic potential of the

investigated configurations, the cycles are benchmarked to each other. In this context

an analysis of baseload conditions for the design­point of each cycle is carried out.

1 In t roduct ion 7

Aspects like part­load operation or off­design behaviour are briefly mentioned but they

are not covered by the benchmarking. In some of the configurations investigated in

addition to CO2 capture is applied membrane reactors are also integrated. In particular

for configurations with integrated membrane reactor, the following questions are

addressed:

What is the reduction in net efficiency of a power generation cycle if CO2 capture

is applied to the cycle?

How is the thermodynamic potential of a power generation cycle with CO2

capture affected if a membrane reactor is additionally integrated into such a cycle?

How does the operating conditions of the membrane reactor impact the net

efficiency of the overall power generation cycle?

2 Carbon Dioxide Capture Technologies 9

2 Carbon Dioxide Capture Technologies

2.1 Overview

The separation of carbon dioxide from power generation cycles can be achieved in

different ways. Technologies to capture CO2 from power generation cycles are

commonly classified by the location of the separation process relative to the

combustion process. In general, there are three different CO2 capture technologies:

Pre­combustion CO2 capture

Post­combustion CO2 capture

Oxyfuel cycles3

These technologies are commly referred to Carbon Capture and Storage (CCS),

although only in the present work, only the capturing of CO2 from the power

generation cycle is addressed. Some aspects of the storage of CO2 are described below.

The above mentioned CO2 capture technologies for power generation cycles are

schemati cally shown in figure 2­1. For all CO2 capture technologies a CO2­rich

stream is separated and further treated for storage or transport. The treatment of the se­

Figure 2-1: Overview of different CO2 capture technologies for power generation cycles

and industrial processes [14].

3 Sometimes referred to as integrated CO2 capture

10 2 .1 Overview

parated CO2 includes a purification (reducing the amount of impurities and inert gases)

and compression above the critical pressure of 73.77 bar [15], commonly to a pressure

of 100­150 bar. At high pressure the separated CO2 is prepared to be either directly

stored or transported via pipelines or trucks.

In the case of post­combustion CO2 capture, the power generation cycle is less

affected compared to the other capture technologies. The process is usually air­driven

and no additional treatment of the fuel is required; therefore an additional treatment of

the exhaust gas stream is required. For pre­combustion capture, the fuel being either

coal or natural gas is converted to a hydrogen­rich synthesis gas (syngas) and the

carbon is separated from the syngas stream. The syngas is used as fuel for the power

generation cycle. For the conversion of the fuel generally technically pure oxygen (95­

97% by mole) is required, which is conventionally supplied by a cryogenic Air

Separation Unit (ASU). In oxyfuel cycles the cycle layout remains nearly unchanged

compared to conventional power cycles but the oxygen is used instead of air as

oxidiser in the combustion process. Therefore the combustion products consist then

mainly of CO2 and H2O. The oxygen required is also supplied by a cryogenic ASU.

For these three different CO2 capture technologies some examples of established

cycles are briefly described. For the different power generation cycles with CO2

capture coal (hard coal and bituminous coal) and natural gas are considered in the

present work because together they have a share of about 49% on the primary energy

supply worldwide [12]. Furthermore, most of the publications focus on power

generation cycles, which are either coal or natural gas fired cycles.

Research on power generation cycles with CO2 capture started in the 1990s. A

large variety of different concepts of power generation cycles with CO2 capture can be

found in the literature. An early published (1998) and often cited publication from

Bolland and Mathieu [16] deals with two different options (oxyfuel and post­

combustion) CO2 capture from a natural gas fired combined cycle power plant.

Qualitative and quantitative comparison of coal fired and natural gas fired power

generation cycles have been continued in the last decade by various authors [17­24].

Summarising it can be said that, depending on the fuel and the capture technology, the

net efficiency of the power generation cycles decreases by 9­12% points if CO2

capture is applied. Besides the ‘invention’ of innovative configurations with CO2

capture, in most cases benchmarking is carried out for the design­point to show the

thermodynamic potential of the different power generation cycles with CO2 capture.

The different configurations are usually divided by the type of fuel they are using. The

latest publications deal with a more detailed analysis of existing concepts for power

2 Carbon Dioxide Capture Technologies 11

generation cycles with CO2 capture, rather than coming up with new configurations.

Recent publications on power generation cycles with CO2 capture address topics like

transient behaviour of the power plant [25], part­load (off­design) operation [26, 27],

reliability and operability analysis [28] as well as analysing and optimising existing

cycles in terms of performance (power output and efficiency) and costs [29, 30].

The goal of capturing CO2 from power generation cycles is to prevent its releasing

into the atmosphere; therefore after separation of CO2 from the power generation cycle

it needs to be stored. Due to the enormous amounts of CO2 potentially captured from

fossil fuelled power plants, the only possibility is to store it safely somewhere on a

long term basis. Currently three different methods of storing CO2 are discussed:

Geological storage [31­35]

Oceanic storage [36­38]

Enhanced Oil Recovery (EOR), where CO2 is injected into an oil reservoir (with

the goal to increase the delivery rate of oil) [39­41]

Different geological structures seem to be able to store CO2 safely in the long term.

Therefore different structures have been investigated worldwide for storing CO2, for

example in Australia the storage of CO2 in bituminous coal seams is investigated [31].

In the United States the storage of CO2 in caprock is studied [32]. Furthermore saline

aquifers are considered for a storage of CO2 at a depth of 1000­3000 m [33­35]. In

studies regarding the oceanic storage, direct disposal of liquid CO2 on the ocean floor

is considered. At oceanic depths below approximately 3000 m, liquid CO2 density is

higher than that of seawater and CO2 is expected to sink and form a pool on the ocean

floor. In addition to chemical reactions between CO2 and seawater to form a hydrate,

fluid displacement is also expected to occur within the ocean floor sediments [36].

Apart from modelling the oceanic storage of CO2 [36], experimental work has been

carried out in the Pacific Ocean, too [37, 38]. EOR is considered as possible

application to utilise the captured CO2. The “IEA GHG Weyburn CO2 Monitoring and

Storage Project” was launched in 2000 by the Petroleum Technology Research Centre

in Regina, Canada [39]. The Weyburn field is located in southeast corner of the

province of Saskatchewan in western Canada. The Weyburn unit is an oil field with a

size of 180 km2. Different studies draw conclusions about the long­term storage in

EOR based on three­dimensional models [40, 41]. The goal of all studies about storing

CO2 is to assess the long­term storage and predicting leakage rates versus time. The

time frame in such studies range from 100 to 1000 or even more years.

12 2 .2 Pre­combust ion CO2 Capture

In addition to technical questions of storing CO2, economic, social and political

issues are addressed in various studies. According [42] the costs for transport and

storage are estimated to range from 0 to 15 $/T CO2. In case large­scale application

storing of CO2 were realised, legislation either national or European regulation would

be required. For the time being such legislation does not exist and [43, 44] highlight

which gaps need to be addressed before large­scale storage could take place. For

example geological storage of CO2 has been conducted also in Germany in the

framework of an European research project CO2STORE4 [45]. In the CO2STORE

project, CO2 is stored at a depth of around 1600 m because the “Schweinrich

structure” was selected as the most suitable candidate in north­eastern Germany. For

the time being there is no national legislation in Germany in force to regulate the

geological storage of CO2. It shows that not only technical challenges in Germany but

also legal ones have to be overcome before large­scale CCS can be realised. Also

social and political issues have to addressed on the way of large­scale realisation of

CCS. Social and political questions are discussed in terms of the development of CCS

within the European Union until 2050 in [46].

Some power generation cycles with CO2 capture belonging to the three different

technologies (pre­, post­combustion and oxyfuel cycles) are briefly presented in the

following.

2.2 Pre-combustion CO2 Capture

In power generation cycles with pre­combustion CO2 capture the carbon of the fuel

is removed before the fuel is combusted for example in a gas turbine. Usually

technically pure oxygen is used as oxidiser instead of ambient air. Conventionally the

oxygen is generated by means of a cryogenic ASU. In case of coal (hard coal or

lignite) gasification of coal to a CO­ and H2­rich syngas takes place under the

presence of steam and oxygen. The syngas produced requires cleaning (de­dusting,

COS­hydrolysis and desulphurisation) before in a CO­shift reactor CO is converted to

CO2, which can be separated from the syngas stream. For reasons of combustion

stability and NOx emission levels, the H2­rich syngas is diluted with nitrogen before it

is combusted in a gas turbine of a combined cycle power plant. This configuration is

referred to as Integrated Gasification Combined Cycle (IGCC). Studies have shown

that IGCC is a promising technology for coal­fired power generation cycles with CCS

4 CO2STORE is part of fifth European framework programme.

2 Carbon Dioxide Capture Technologies 13

[47, 48]. An IGCC with and without CCS is explained in greater detail in section 4.2

because different IGCC configuration are thermodynamically assessed in the present

work.

In particular the aspect that an IGCC could be used for co­generation of hydrogen

and electricity makes it an interesting power generation cycle in future power

generation concepts [49]. It is currently planned by RWE Power AG to built an IGCC

with CCS in Germany, which should start operation early in 2015 [50]. In addition a

feasibility study has been performed for realising an IGCC with and without CCS in

Germany [51]. Other studies investigating innovative concepts for IGCC cycles

propose for example plasma gasification [52] or a combination of an IGCC with a

Graz cycle [53].

In the case of natural gas fired combined cycles the fuel is catalytically converted

in an AutoThermal Reformer (ATR). At first the natural gas is converted to CO and

H2, the CO is further converted to CO2 in a water­gas section. The CO2 can be

separated by chemical or physical absorption and the hydrogen­rich syngas is burned

in a gas turbine [54­56].

2.3 Oxyfuel Cycles (Integrated CO2 Capture)

Oxyfuel cycles are applicable to both natural gas fired combined cycles as well as

to the coal fired steam plants. The ‘only’ change is that ambient air is replaced by

technically pure oxygen. In general the plant layout might be slightly affected by the

switch from air to oxygen but remains principally unchanged. The exhaust gas

produced during the combustion process comprises mainly CO2 and H2O due to the

absence of nitrogen. For both type of fuel (coal and natural gas) the combustion

process operates close to stoichiometric conditions. The excess oxygen in the

combustion process is in the order of 1 to 5%. This is because the production of

oxygen in a cryogenic ASU requires high expenditure of energy. For controlling the

temperature in the combustion process a significant amount of the exhaust gas is sent

back to the combustion process. The exhaust which is recirculated can be considered

as a CO2­rich stream because the recirculation is done after cooling of the exhaust gas.

Most of the water in the exhaust gas is already condensed out as a result of the cooling

process; therefore, in case of combined cycles, the working fluid in the compressor is

mainly CO2. This shows that a conventional gas turbine cannot be used for such an

oxyfuel cycle but a new CO2 compressor needs to be designed. Also the ratio of

14 2 .4 Post ­combust ion CO2 Capture

compressor inlet mass flow and turbine may changed compared with a conventional

gas turbine.

Studies on natural gas fired oxyfuel cycles can be found in [57­59]. Several studies

dealing with oxyfuel combustion for hard coal, lignite or bituminous coal have been

conducted [60­63]. Another recent publication addresses the optimisation of

purification of the separated CO2 [64]. Lignite fired boiler cycles with and without

CCS are described in more detail in section 4.3 because various configurations with

and without CCS are thermodynamically assessed in the present work. Oxyfuel cycles

are currently only considered for coal fired power plants because the gas turbine

required for such an oxyfuel cycle does not currently exist; therefore gas turbine

oxyfuel cycles are of more interest to the scientific community.

Another type of oxyfuel cycle is Chemical Looping Combustion (CLC) where

commonly natural gas as fuel reacts with a metal oxide powder (particles size <

10 m). The CLC comprises two reactors in which the reactions takes place. The

reaction is unlike a conventional combustion process because the air and the fuel

remain in separate environments and have no direct contact with each other. In a first

reactor the reaction of the fuel and the metal oxides takes place (reduction reactor). In

a second step, the reduced metal oxide is circulated to the second reactor, the oxidation

reactor. The metal oxide carries the chemical energy in from of heat to the oxidation

reactor, reacts with oxygen in the air, and is regenerated to a metal oxide. The metal

oxide then circulates back to the reduction reactor to react with the fuel. These two

reactors would substitute the combustor of a gas turbine [65­67]. The latest research

activities focus on the investigation of new metal oxides for increasing the reactivity

[68­70] and to apply the CLC concept to combustion of syngas [71] or coal [72].

Lately a 1 MW CLC pilot plant started first test operation in December 2010 at the

Technical University of Darmstadt [73].

2.4 Post-combustion CO2 Capture

The attractiveness of post­combustion CO2 capture technologies is that the power

generation cycle itself it not much affected. Both the fuel and the oxidiser (ambient

air) remains unchanged and the produced exhaust gas is treated after leaving the

process in the CO2 capture unit. Post­combustion CO2 capture is considered for coal­

fired (and lignite­fired) steam power plants [74­76] as well as for natural gas fired

combined cycle power plants [77­80]. Due to the different carbon content in coal and

2 Carbon Dioxide Capture Technologies 15

in natural gas, the major difference for the two fuels is the concentration of CO2 in the

exhaust stream. For coal­fired power plants the concentration of CO2 is around 12­

14% by mole, whereas it is only approximately 4% by mole for natural gas fired

combined cycles.

Commonly, the separation of CO2 from the exhaust gas stream is achieved by

chemical absorption. An aqueous solvent such as Monoethanolamine (MEA) is used.

The absorption has to take place at a lower temperature than the ‘standard’ exhaust

temperature of around 80°C; therefore the exhaust gas needs additional cooling to a

temperature of around 40°C. Due to the properties of the aqueous solvent, the

regeneration of the solvent has to take place on a higher temperature level (120­

140°C). The required heat which is needed in the regeneration process is usually

provided by low pressure (LP) steam from the water/steam cycle. This is the major

drawback to post­combustion CO2 capture because the regeneration requires a large

amount of specific energy. The reduction in power output due to the extraction of large

quantities of LP steam results in a penalty in net efficiency of the overall power

generation cycle by around 9­12% points (see [17­24]).

For natural gas fired combined cycles an idea is to recirculate a fraction of the

exhaust gas to increase the CO2 concentration in the exhaust gas stream and to lower

the mass flow to be treated in the CO2 capture unit [78­80] and thus to decrease the

penalty in efficiency due to CO2 capture. The layout of CO2 capture and its integration

into the power plant has not changed much in the latest publications. A more detailed

modelling of the absorption unit has been conducted; for example dynamic modelling

of the absorption process [81, 82]. A newer separation technique is Chilled Ammonia

Process (CAP), where the absorptions takes place at an even lower temperature

(around 5°C) and ammonia (NH3) is used as solvent. The promising feature of CAP is

that the specific expenditure of energy maybe lower compared to MEA [83, 84]. Due

to the drawback of high specific expenditure of energy for the regeneration, the

research activities focus on new solvents to lower both the regeneration temperature

and the expenditure of energy for the regeneration process [85, 86]. Furthermore, some

work has been conducted to use ionic liquids [87] or activated carbon [88] as solvents

for post­combustion CO2 capture. Finally another study shows potential for

optimisation of the construction phase of a CO2 capture unit using chemical absorption

[89].

16 2 .5 Power Genera t ion Cycles wi th Integra ted Membrane Reactors

2.5 Power Generation Cycles with Integrated Membrane Reactors

In this work power generation cycles with integrated membrane reactor are

thermodynamically analysed. Therefore this section is introduced to briefly emphasise

some aspects on power generation cycles with integrated membrane reactors. At the

point in time when the present work was initiated, only a few publications were

available dealing with membrane reactors integrated in power generation cycles. In the

last five years the interest in membrane reactors as part of power generation cycles

with CO2 capture has increased. Habib et al. [90] and Bredesen et al. [91] give good

overviews of the latest developments and different applications of membrane reactors.

Another review on power generation cycles in particular with pre­combustion CO2

capture can be found in [92]. The earlier mentioned benchmarking studies [17­24]

comparing different power generation cycles with CO2 capture cover also partly

membrane­based cycles.

Reviews on polymeric membranes for separating CO2/N2 mixtures, which are not

covered by this work can be found in [93, 94]. All authors of the two reviews work in

the area of membrane development. Therefore the focus of both reviews is on the

properties of the different membrane materials; they also gather a lot of data of various

publications. The application of polymeric membranes to separate CO2/N2 mixtures

are described by [95, 96]. Franz and Scherer [95] analysed an IGCC process with

integrated membrane reactor for pre­combustion CO2 capture, whereas Kotowicz et al.

[96] have such a membrane investigated for post­combustion CO2 capture for a coal

fired steam power plant.

The AZEP5 cycle is an innovative concept, where a OTM reactor is integrated into

a gas turbine cycle [97­100]. Apart from doing simulations, a main goal of the AZEP­

project was to develop the membrane reactor as the most important component of the

concept and to conduct experimental testing of the membrane reactor [101].

Assumptions for the operating conditions regarding flow velocities inside the

membrane reactor have been adapted for the present work. A similar project is the

Oxycoal­AC project [102­104]. The Oxycoal­AC project is coordinated by the

‘Lehrstuhl für Wärme­ und Stoffübertragung’ at the RWTH Aachen University. The

goal of the project is to develop an OTM reactor for a coal fired oxyfuel boiler. It

includes the design of the membrane reactor [105] and to built a pilot reactor to proof

the feasibility of the concept.

5 Advanved Zero Emission Power (AZEP). The AZEP­project is part of fifth European framework program.

3 Membranes for Gas Separa t ion 17

3 Membranes for Gas Separation

3.1 Background Information

The principle of separation of gas streams by means of membranes has been known

for more than 100 years, though large­scale applications have only appeared in the past

50 years. Membrane processes for separation purposes covers a large variety of

applications, mainly in chemical process engineering; industrial applications are:

reverse osmosis, dialysis, electrodialysis, microfiltration, ultrafiltration, pervaporation,

liquid membranes and gas permeation. In the case of gas permeation the main

applications are [106­108]:

Separation of CO2 or H2 from methane or other hydrocarbons,

Adjustment of H2/CO ratio in synthesis gas,

Natural gas dehumidification,

Separation and recovery of H2 from industrial gases,

Separation of air into either nitrogen­ or oxygen­enriched streams,

Recovery of helium,

Recovery of CO2 or CH4 from biogas.

As previously described, during the last decade membranes for gas separation have

more and more found their way into power generation cycles with CO2 capture. In

particular in power generation cycles with either pre­combustion CO2 capture or

oxyfuel cycles with CO2 capture. Examples for power generation cycles with

integrated membrane reactors are provided in section 2.5.

In general, in a separation process occurring in a membrane reactor the feed

comprising a mixture of two or more components enters the reactor on the feed side

and is partially separated from the stream by means of a semi­permeable barrier: the

membrane. The component separated from the feed permeates through the membrane

and leaves the membrane reactor as permeate. Optionally, a sweep gas can be used on

the other side of the membrane. In this case the permeate is a mixture of sweep gas and

the separated component. The remaining part of the feed stream leaves the membrane

reactor as retentate. The separation process taking place in a membrane reactor is

generally shown in figure 3­1.

The membrane reactor may either be designed as coflow, counterflow or crossflow

apparatus. In case of a coflow flow reactor both streams, feed and sweep, enter the

reactor on the same side. In a counterflow reactor feed and sweep enter the reactor

from opposite sides. If no sweep gas were used, the reactor would be a crossflow appara­

18 3 .1 Background Information

Figure 3-1: Separation process in a coflow membrane reactor [106].

tus, where the permeate leaves the reactor immediately after permeating through the

membrane. The design of the membrane reactor depends on the application it is used

for. Examples for different applications and explanation of those membrane reactors

can be found in [106­108].

The driving force for the mass transport through the membrane is the difference in

partial pressure across the membrane of the component separated by the membrane

reactor. Generally the difference in partial pressure is determined and can be adjusted

by three parameters:

(i) Difference in concentration of the component separated in the membrane

reactor

(ii) An additional mass flow of a sweep gas

(iii) A difference in total pressure to increase the difference in partial pressure

across the membrane additionally

For the membrane reactors considered in this work both options (ii) and (iii) are

considered for the separation process. A sweep gas is used to lower the partial pressure

of the permeate on the sweep side of the membrane. The pressure difference across the

membrane is achieved by elevating the pressure of the feed stream. Additionally the

pressure of the sweep gas is also varied to change the difference in total pressure

across the membrane. A sensitivity analysis of these two parameters, pressure and

mass flow rate of the sweep stream, is carried out to investigate the impact on the

permeation flux through the membrane; this is described in detail in section 6.5. In

general, the difference in total pressure could also be achieved by applying a vacuum

pressure on the sweep side of the membrane.

Depending on the operating conditions of the membrane reactor, the membrane

may need to be able to withstand a difference in total pressure across the membrane.

3 Membranes for Gas Separa t ion 19

Figure 3-2: General structure of membranes: (a) symmetric membrane; (b) asymmetric

membrane, (c) composite membrane [109].

Commonly, inorganic membranes for gas separation show a structure shown in

figure 3­2. Single phase membranes are symmetric membranes: they consist of one

material. Palladium (Pd) membranes for separation of hydrogen, and mixed ionic and

electronic conducting (MIEC) membranes are examples of symmetric membranes.

Dual phase membranes are asymmetric membranes comprising a thin top layer and a

support structure that provides the mechanical strength. The selective properties of

such a membrane is determined by the top layer. Ideally the support structure

contributes as low as possible a resistance to the flow; therefore the support is typically

made of a porous structure having relatively large pores (commonly 2­20 m).

Another type of membrane is a Pd­based membrane in form of a composite membrane,

where the support is a metal structure. Since the support is an integrated part of the

membrane structure, its properties impact significantly the membrane performance

[109].

The geometry of the membrane itself as part of a membrane reactor is usually kept

to a simple shape for manufacturing reasons. Common membrane shapes are

illustrated in figure 3­3. Of course the simplest geometry is a flat sheet membrane. The

disadvantage of a flat membrane is a low surface/volume­ratio. In case of large

required membrane surface area, the overall reactor size becomes large. The

surface/volume­ratio is much higher for tubular or monolithic membranes. Hence

these two geometries are commonly considered for membrane reactors. The

membranes may either be asymmetric or symmetric. A tubular shape is assumed in the

present work (see chapter 6).

The different membrane geometries are incorporated into compact modules. A

membrane reactor comprises a certain number of modules to achieve the desired

membrane surface area. Both types of membrane geometries, tubular or monolithic,

resemble a shell­and­tube heat exchanger as shown in figure 3­4. The feed stream

might flow through the tubes or channels, and would permeate through the membrane

to the shell side of the module. The permeate might leave the module at different

location (as indicated in figure 3­4). In case a sweep gas is applied to the reactor, the

module may either be designed as coflow or counterflow apparatus. A commercial mo­

20 3 .1 Background Information

(a) (b) (c)

Figure 3-3: Different membrane geometries: (a) flat sheet; (b) tubular; (c) monolithic

[106].

dule might be 1 m long and up to 0.3 m in diameter [106].

When designing such a membrane module a variety of issues needs to be taken into

account. For modules in commercial­scale, a large challenge is to make the feed

stream gas­tight to the sweep stream. The chosen material of the membrane and the

casing of the reactor determine the possibility of producing a gas­tight joint between

the different materials. For instance brazing a ceramic membrane to a casing made of

stainless steel is successfully used for an oxygen transport membrane reactor [102]. Of

course the type of joint has to cover the whole operating temperature range of the

membrane reactor. Another practical issue is different thermal expansion of materials

of the membrane reactor. In case that different materials are joined together the design

has to ensure that the membrane cn withstand the high mechanical load due to thermal

expansion. The maximum thermal gradient of the whole membrane reactor might need

to be limited for transient operation such as start­up and shutdown of the reactor. An

operational aspect is the pressure drop for both streams. Assuming that the feed stream

(a) (b)

Figure 3-4: Common membrane modules: (a) tubular; (b) monolithic [106].

3 Membranes for Gas Separa t ion 21

is determined by the process into which the membrane reactor is integrated, most

likely the reactor will be designed such that the pressure drop is small for the feed

stream. Due to a much smaller sweep stream the pressure drop on the sweep side is

likely to be less important. Commonly, the feed stream flows inside the tubes or

channels. This is also assumed in this work (see chapter 6).

3.2 Membranes for Oxygen and Hydrogen Separation

3.2.1 Oxygen Transport Membranes

Type of Membranes

In general, an Oxygen Transport Membrane (OTM), also called an Ion Transport

Membrane (ITM), is used to separate oxygen from an oxygen containing mixture, such

as air. The ceramic membrane material allows the diffusion of oxygen anions and

potentially produces infinite selectivity for oxygen. The diffusion of oxygen anions

might be driven either by an oxygen partial pressure gradient across the membrane or

by an electrical potential gradient across the membrane [110].

The two main types of oxygen separation systems based on ceramic membranes are

pure oxygen­ions conducting membranes and mixed ionic­electronic conducting

(MIEC) membranes [111]. The two different types of ceramic membranes are

schematically illustrated in figure 3­7. In solid electrolytes the oxygen is transported

through the membrane in ionic form. A simultaneous flux of electrons in opposite

direction using an external circuit is required to fulfil electrical neutrality of the

membrane (see figure 3­7­a). Due to the capability of MIEC membranes the transport

of both oxygen ions and electrons takes place inside the membrane; therefore the opera­

(a) (b) (c)

Figure 3-5: Ceramic membrane types based on conduction mechanisms: (a) solid

electrolytes; MIEC membranes: (b) single phase membrane; (c) dual phase

membrane [111].

22 3 .2 Membranes for Oxygen and Hydrogen Separat ion

tion does not require electrodes and external circuit. Oxygen anions permeate through

a MIEC membrane from the high oxygen chemical potential side to the low oxygen

potential side, if, firstly, a MIEC membrane is placed under an oxygen chemical

potential gradient. Secondly, the MIEC membrane shows its conductivity in

temperature range of 750 to 1000°C [111, 112]. Therefore 750­1000°C is the range in

operating temperature for most MIEC membranes. MIEC membranes can either be a

single phase or dual phase membrane (see figure 3­7: b and c). Single phase

membranes consist of one material being capable of conducting both oxygen ions and

electrons. Dual phase membranes comprise two materials – an oxygen­ion conductor

such as doped zirconia and an electron conductor such as a noble metal [110]. The

different type of MIEC membranes are also summarised by [113].

Membrane Materials

In general MIEC compounds are dense ceramic membranes. The most common

types of MIEC materials exhibiting both high ion and electron conducting properties

are perovskites. These materials have high stability at elevated temperatures due to

their orthorhombic structure [114]. Perovskites are originally named from a mineral

oxide CaTiO3. The basic structure of this mineral was first thought to be cubic but it

was later found to be orthorhombic, and the name perovskite has been retained for this

type of structure [112].

The general structure of an ideal perovskite consists of the formula ABO3, where B

is a transition metal cation such as: iron (Fe), cobalt (Co), chromium (Cr), titanium

(Ti), manganese (Mn), nickel (Ni) or copper (Cu). The A­site ion may either be alkali

metals: lithium (Li) or rubidium (Rb), alkaline earth metals: barium (Ba), magnesium

(Mg), strontium (Sr), calcium (Ca) or sodium (Na). The general structure of a

perovskite material is shown in figure 3­6. In perovskites the A­site ions and the

oxygen ions form a shared cubic close packing in which octahedral vacancies the B

cations are located [102, 112, 114]. Therefore, on the one hand, a high number of

oxygen vacancies result in a large mobility of the oxygen ions and hence to a high

permeation flux of oxygen ions through the membrane. On the other hand, a high

number of oxygen vacancies in the perovskite structure at the same time cause

different stresses in the lattice. In addition high temperature and difference in oxygen

partial pressure also lead to different stresses within the lattice of the perovskite

structure. This stress within in the perovskite structure reduces the long­term stability

[115].

3 Membranes for Gas Separa t ion 23

Figure 3-6: Crystal structure of a perovskites ABO3 composite oxide [102].

An ideal perovskite consisting of only ABO3 is unlikely to be able to conduct

oxygen ions. For oxygen ions to diffuse into the perovskite, imperfections must occur

within the crystal structure. The imperfections or crystal defects provide oxygen

vacancies, which the diffusing ions can occupy while travelling through the

membrane. The ionic conduction can be understood as a hopping from one oxygen

vacancy to the next [114, 116]. Describing the mass transport of oxygen through a

dense mixed conducting membrane it can be said that the oxygen permeation occurs

via the transport of oxygen ions through the crystal structure. This kind of transport of

oxygen ions through the membrane is termed bulk diffusion.

When choosing a material for oxygen separation a large variety of properties have

to be considered. The most important properties of an oxygen transport membrane are:

oxygen­ion conductivity, electron conductivity, creep resistance, thermal expansion,

chemical stability, chemical expansion, CO2 tolerance, and steam tolerance [110].

These different properties can be influenced by doping the membrane material. Doping

is a principal method of tailoring the physical properties of the mixed conducting

material by means of the formation of nonintegral stoichiometric phases or solid

solutions by homogeneous doping with appropriate elements [117]. For perovskite­

structured oxides ABO3, it is often seen that a lower valence dopant B’ is introduced

on the B site to produce ABꞌ1­xBꞌxO3­6, thus forming defects. The symbol expresses

the amount of the ion vacancies (defects), which provide the pathway for ions [112,

118]. A wide range of compounds based on a perovskite structure can also be

produced with substitutions occurring for either the A atom, the B atom or both to

form a structure AxAꞌ1­xByBꞌ1­yO3­ [119].

Until now an enormous number of MIEC materials have been studied and are

available in literature (see table 3­1 and [112]). Two of the first publications on MIEC

6 A prime (ꞌ) denotes a negative excess charge [112].

24 3 .2 Membranes for Oxygen and Hydrogen Separat ion

as membrane material were issued in the 1980’s [120, 121]. Most of the materials are

of perovskite structure, including the most important families of Sr(Co,Fe)O3,

La(Co,Fe)O3, LaGaO3, and other non­perovskite structures. The most important

families are briefly described in the following. Liu et al. [112] gives an detailed

overview of the above mentioned materials. The following brief description gives an

impression about the large variety of different MIEC materials, which have been

studied in the last decades.

The first typical perovskite oxides are SrCoO3 and SrFeO3 (SCFO), which have

been studied intensively (see table 3­1 and [112]). SCFO perovskites achieve

remarkable high oxygen permeation rates. However, SCFO perovskites are usually

considered to be thermodynamically and structurally unstable at lower temperatures

and low partial pressure of oxygen. Therefore the effect of barium was investigated by

doping SrCo0.8Fe0.2O3­. The partial substitution of strontium with barium ions results

in perovskite oxide in the form of Ba1­xSrxCo0.8Fe0.2O3­ with x = 0.3­0.5. It was found

that the phase stability was greatly improved while the conductivity was not decreased.

It was further investigated to substitute the remaining strontium ion with titanium to

form a perovskite structure of BaTi0.2Co0.5Fe0.3O3­. Doping with zirconium instead of

titanium BaCo0.4Fe0.4Zr0.2O3­ was also found to achieve high permeation fluxes. [112]

The second family of perovskite oxide is La(Co,Fe)O3 (LCFO), for which also a

large number of different materials have been investigated (see table 3­1 and [112]).

LCFO­based MIEC materials exhibit significant ionic conductivity with prevailing

electronic conductivity. Although the oxygen permeation flux is lower compared with

SCFO­based materials, some problems suffered in those perovskites were minimised

in LCFO materials. In LCFO oxides, La can be partially substituted by divalent metal

cations such as barium, strontium or calcium. For cobalt substitution in LaCoO3­

gallium, chromium, iron, lead, and nickel have been proposed. For LaCo1­xCrxO3 (with

x = 0.1­0.4) it was found that oxygen permeation flux, electrical conductivity and

thermal expansion all decrease with increasing chromium concentration. Cobalt may

also be simultaneously substituted by nickel and iron. For the oxide system LaCo1­x­

yFexNiyO3 (with x = 0.1­0.2 and y = 0.1­0.3) the introduction of nickel leads to an

increase of electrical conductivity and decreasing thermal expansion. Furthermore, it

was found that the double­site­substitution perovskites in the form of La1­xSrxCo1­

yFeyO3­ (LSCF) give high oxygen permeability and also show a good stability under

air atmosphere, and were thus recognised as promising materials for air separation.

Another composite is La0.2Sr0.8Co0.2Fe0.8O3­. [112]

3 Membranes for Gas Separa t ion 25

The third group of typical perovskite oxides are LaGaO3 (LGO), which is used as

basis for different variations studied intensively (see table 3­1 and [112]). LGO­based

MIEC materials achieve a lower relative electronic conductivity but a higher ionic

conductivity, and have therefore attracted much attention in solid oxide fuel cells. The

cations of titanium, chromium, iron, cobalt and nickel have been used as dopants for

improving the physiochemical and transport properties. It was found that LaGa1­xNixO3­

Table 3-1: Different perovskite materials proposed for MIEC membranes.

Perovskite family: Sr(Co,Fe)O3 – (SCFO)

Composition Short name Reference

Ba0.5Sr0.5Co0.8Fe0.2O3­ BSCF [123­126]

Ba0.5Sr0.5Zn0.2Fe0.8O3­ BSZF [127]

SrCo0.8Fe0.2O3­ SCF [128, 129]

SrFe0.33Co0.67O3­ SFC [130]

SrFe0.67Co0.33O3­ SFC [130]

SrCo0.8Fe0.2O3­ doped with Al2O3 SCFA [131]

Sr0.95Co0.8Fe0.2O3­ SCF [132]

Perovskite family: La(Co,Fe)O3 – (LCFO)

Ba0.8La0.2Co0.8Fe0.2O3­ BLCF [128]

La0.6Sr0.4Co0.2Fe0.8O3­ LSCF [129, 133­135]

Sr0.8La0.2Co0.8Fe0.2O3­ SLCF [128]

La0.2Ba0.8Co0.8Fe0.2 O3­ LBCF [136]

La1­xSrxFe1­yGayO3­ LSFG [137]

(La0.75Sr0.25)0.95Cr0.5Mn0.5 O3­ LSCM [138]

La0.2Co0.8SrO3­ doped with CeO2 LCS [139]

La0.6Sr0.4CoO3 LSC [140]

La1­xSrxFeO3­ LSF [141]

Perovskite family: LaGaO3 (LGO)

LaGa0.65Ni0.20Mg0.15O3­ LGNM [122]

La0.85Ce0.1Ga0.3Fe0.65Al0.05O3­ LCGFA [142]

La0.5Pr0.5Ga0.65Mg0.15Ni0.2O3­ LPGMN [143]

La0.9Sr0.1Ga0.65Mg0.15Ni0.2O3­ LSGMN

La0.8Sr0.2Ga0.8Fe0.2O3­ LSGF [144]

La0.8Sr0.2Ga0.6Fe0.4O3­ LSGF [144]

La0.7Sr0.3Ga0.6Fe0.4O3­ LSGF [140]

LaGa1­xNixO3­ (with x = 0.2­0.6) LGN [145]

LaGa0.3Co0.6Mg0.1O3­ LGCM [146]

26 3 .2 Membranes for Oxygen and Hydrogen Separat ion

(with x = 0.2­0.5) possesses high oxygen permeation fluxes as well as low thermal

expansion coefficients. Furthermore, the substitution of gallium simultaneously by

nickel and magnesium yielding the perovskite structure LaGa0.65Ni0.20Mg0.15O3­. It

was found that the oxygen permeation could be increased by minimising the amount of

nickel, necessary to provide sufficient electronic conductivity, and adding magnesium

for increasing the oxygen deficiency [122]. The doping with strontium for the

lanthanum site and with magnesium for the gallium site was found to increase the

electrical conductivity. As a result La0.8Sr0.2Ga0.83Mg0.17O2.815 (LSGM) is an optimum

composition exhibiting stable ionic conductivity. [112]

Arbitrarily chosen examples of the three mentioned families of perovskites are

shown in table 3­1. The brief description of some of the different variations should

reveal that the development of new and improved MIEC materials is an active topic in

materials research. The evaluation and benchmarking of different MIEC materials has

not been the main focus of the present work. Therefore a ‘typical’ MIEC material was

chosen, the calculation of the mass transport through such a membrane was evaluated

by means of experimental data from [122] and the permeation fluxes were adapted to

the required membrane surface areas. This is described in chapter 6 in detail, where

also the modelling of the membrane reactor is depicted.

Regarding the operating conditions of an OTM reactor, the necessary difference in

oxygen partial pressure across the membrane is adjusted in different ways. In most of

the studies mentioned, the work experiments are carried out at atmospheric pressure,

on both sides of the membrane. The oxygen partial pressure is lowered by means of a

sweep gas on the permeate side of the membrane. In most cases either pure helium or a

mixture of helium and steam is used as sweep gas [131, 132, 136]. Only in a few cases

the tests are conducted with a pressure difference across the membrane. For instance in

[147] the feed pressure was chosen with 16 bar.

Compared to the amount of investigated MIEC material available in the literature,

the number of articles dealing with the impact of species such as CO2 on the

membrane performance is relatively small. According to Arnold et al. [125] pure CO2

as sweep gas immediately stops the oxygen permeation for a BSCF membrane,

whereas such membranes are capable of sustaining the oxygen permeation up to 10

mol­% CO2 in the feed stream. Caro [148] proposes barium­ and strontium­free MIEC

materials for applications in power generation cycles with CO2 capture. Caro [148]

recommends materials such as Fe­doped LaNi oxides.

3 Membranes for Gas Separa t ion 27

Mass Transport through a MIEC Membrane

The permeation of oxygen through a MIEC perovskite membranes is a complex

process that involves a number of different steps, which are schematically shown in

figure 3­7. The transport of oxygen can be divided into different steps connected to a

certain location of the cross­section of the membrane as follows [149]:

From the bulk air, the oxygen is firstly transported through a boundary layer to the

membrane surface (A)

At the membrane surface the oxygen dissociates (B) into the lattice of the

membrane material and diffuses to the other side of the membrane (C)

On the permeate side gaseous oxygen is formed at the membrane surface (D)

Finally, the oxygen diffuses from the membrane surface through a boundary layer

to the permeate bulk (E)

Figure 3-7: Simplified representation of oxygen transport through a MIEC membrane

[149].

General consensus has been reached that the overall oxygen permeation flux is

either controlled by bulk diffusion through the perovskite (C) or by surface exchange

on the permeate side of the membrane (D), provided that no mass transfer limitations

occur in the boundary layers. In case the oxygen permeation flux is determined by bulk

diffusion (C), it can be described by the Wagner equation, see [150] for a detailed

discussion7. High oxygen diffusivity (D in cm2 s­1) is important for achieving a high

oxygen permeation flux through thick membranes while a high surface oxygen

exchange coefficient (k in cm s­1) is critical for thin membranes. The thickness at

which the ratio of D to k equals 1 is called the critical thickness Lc (Lc = D/k). Below

the critical thickness the mass transport through the membrane is limited by surface

7 The theory of solid state electrochemistry can be found in [150].

28 3 .2 Membranes for Oxygen and Hydrogen Separat ion

exchange; above the critical thickness bulk diffusion determines the mass transport

through the membrane [151]. The critical thickness depends on the MIEC material and

on the operating temperature of the membrane. For MIEC membranes the critical

thickness may range from 1 to 500 m [109].

The Wagner equation is named after C. Wagner, who reported this relationship

already in the nineteen­thirties [152]. Recent publications use the Wagner equation to

describe the mass transport through MIEC membranes – in case of bulk diffusion

determining the mass transport [153­155] trough the membrane. The Wagner equation

can be expressed as

2

2

,

,

log

A

m

E

O FeedR Tm

m O Sweep

pcj e

X p. (3­1)

In equation (3­1) j denotes the oxygen permeation flux, cm a pre­factor, Xm the

membrane thickness and EA the required activation energy. The Wagner equation is

used in section 6.5.1 for calculating oxygen permeation flux through the MIEC

membrane.

3.2.2 Hydrogen-selective Membranes

Introductory Remarks

For the separation of hydrogen by means of membranes, dense bulk palladium (Pd)

or palladium­based membranes are widely used. With regards to power generation

cycles, hydrogen­selective membranes are commonly used for the enhancement of

water­gas­shift reaction [156­158], where CO is converted to CO2 by adding steam. In

general three different types of membranes for separation of hydrogen can be

distinguished:

Bulk palladium membranes

Palladium­alloy composite membranes with dense support

Palladium­alloy composite membranes with porous support

A detailed overview of hydrogen­selective membrane can be found in [109, 159,

160]. The history and applications of palladium­based membranes is provided by

Paglieri in [159], whereas Bredesen et al. [109, 161] describe these membranes

regarding their capabilities and applications for power generation cycles with CO2

capture. Rothenberger et al. [160] have gathered data from literature of permeability,

permeance and flux of thin film palladium membranes.

3 Membranes for Gas Separa t ion 29

Compared with conventional hydrogen separation technologies, such as pressure

swing adsorption, palladium membranes must possess the following capabilities [162]:

Constant high hydrogen flux and permselectivity

Chemical resistance to common gas stream components such as steam, carbon and

sulphur compounds

Long lifetime (> 10,000 h)

Durable (to withstand thermal cycling)

Cost effective and straightforward to fabricate

Able to be assembled into compact modules with membranes that are easy to seal

and replace

Membrane Materials

Bulk palladium membranes are capable of achieving high hydrogen permeation

flux, but beside that the above requirements also need to be fulfilled. In particular, if

such membranes have to withstand large pressure differences at elevated temperatures,

the use of palladium alloys can help to satisfy physical and chemical stability

requirements. Therefore, similar to oxygen MIEC materials for oxygen transport

membranes, a large variety of ternary palladium alloys have been investigated using

elements such as ruthenium, yttrium, aluminium, iron and rhodium [163­165]. It is

again referred to [109, 159] where several palladium alloys are summarised and

aspects regarding the fabrication of palladium­based membranes is addressed in a

comprehensive manner. A widely used palladium alloy is Pd77­Ag23 [166]. Similar to

oxygen transport membranes, a large variety of different palladium­based alloys exist

to optimise the properties of the membrane material for the required features for

specific applications.

Apart from the aspect of altering the properties by substitution of palladium by

different elements, another reason for using different material is the reduction of costs.

Palladium is an expensive material and by using either a palladium alloy or a support

made of a different material such as metals from the refractory group V [167], the

amount of palladium can be reduced. By the use of a supportive structure, which is

made from a different material, the thickness of the palladium membrane can

significantly be lowered. Some considerations regarding costs of palladium

membranes can be found in [109].

30 3 .2 Membranes for Oxygen and Hydrogen Separat ion

Mass Transport through a Palladium-based Membrane

The transport of hydrogen through a dense palladium­based membrane can be

divided into a series of different steps [109, 168], described from the high partial

pressure side to the low partial pressure side:

Diffusion of molecular hydrogen to the metal surface of the membrane (of the

feed side)

Dissociative adsorption of hydrogen on the surface

Transition of atomic hydrogen from the surface into the bulk metal

Atomic diffusion of hydrogen through the bulk metal

Transition from the bulk metal to the surface on the low partial pressure side

(permeate side)

Regeneration of hydrogen molecules on the permeate side of the membrane

Desorption of the molecular hydrogen from the surface

Diffusion of the molecular hydrogen away from the surface to the bulk gas

The flux of hydrogen through palladium is the product of the diffusion coefficient

Dm (Dm in m2 s­1) and the concentration gradient CH (CH in mol m­3) with the flux

of hydrogen atoms NH being twice that of hydrogen molecules [169]:

2

m

2

HH H m

CN N D

X. (3­2)

For thick membranes (Xm > 100 m), the limiting resistance is assumed to be the

transport of hydrogen atoms through the palladium. Under these conditions, the

surface reaction is considered to be very fast and the dissolved hydrogen atoms at the

surface of the palladium are in equilibrium with the hydrogen gas on either side of the

membrane. The concentration of hydrogen atoms in the palladium can be related to the

hydrogen partial pressure via the Sieverts equation. The exponent of 0.5 reflects the

dissociation of the gaseous hydrogen molecule into two hydrogen atoms that diffuse

into the metal, where an ideal solution of hydrogen atoms in palladium is formed and

KS represents the Sieverts constant [169]:

2

0.5H S HC K p . (3­3)

3 Membranes for Gas Separa t ion 31

Combining equation (3­2) and (3­3) yields the Richardson’s equation [170]:

2 22

0.5 0.5, ,

m2

H Feed H Sweepm SH

p pD KN

X. (3­4)

The first product in equation (3­4) corresponds to the hydrogen permeability k,

which is the half of the diffusion coefficient Dm and the Sieverts constant KS:

1

2 m Sk D K . (3­5)

Using the hydrogen permeability k, equation (3­5) can be expressed as:

2 22

0.5 0.5, ,

m

H Feed H SweepH

p pN k

X. (3­6)

Morreale et al. [169] have shown that the exponent in equation (3­6) to describe

the hydrogen permeation flux maybe larger than 0.5. In case that the mass transport

through the palladium­based membrane is not only determined by bulk diffusion but

by surface reactions, the exponent may range from 0.5 to 1.0. Therefore the exponent

in equation (3­6) is generally written as n:

2 22

, ,

m

n nH Feed H Sweep

H

p pN k

X. (3­7)

The temperature dependency of the hydrogen permeability can be expressed by

means of an Arrhenius­type relation [169]:

2 2 2

0, ,

m

( )

A

m

E

R T n nH H Feed H Sweep

kN e p p

X. (3­8)

Since a thick membrane is considered in the present work, thus assuming that the

mass transport is determined by bulk diffusion through the membrane, equation (3­8)

is used to describe the mass transport through the hydrogen­selective membrane in

section 6.5.2.

4 Analysed Power Genera t ion Processes 33

4 Analysed Power Generation Processes

4.1 Common Framework for Modelling of Power Generation Cycles

A common framework of the modelling of power generation processes has been

developed within the scope of subproject 6 in ENCAP. The common framework has

been reported as project internal deliverable D6.1.1 [171] by Bolland et al. in 2004. In

the following, this framework will be referred to as “ENCAP framework SP 6”. The

main goal of the common framework is to provide a set of boundary conditions of

several components which are used in various power generation processes. These

boundary conditions are used to facilitate a unified comparison of those processes. For

instance, the model of the generic cooled gas turbine, which is described in section 5.1,

is also defined by the common framework. Other boundary conditions from D6.1.1

[171] are applied in the simulation presented in this work are described in the

following paragraphs.

Since most of the current work has been undertaken within the framework of the

ENCAP project, a further ENCAP internal deliverable has been used as reference for

boundary conditions used for simulations in this work. In SP 1, guidelines for various

technology concepts are reported as ENCAP project internal deliverable D1.2.2 [172]

by Biede et al. in 2008, which is an updated version of the original deliverable D1.2.2

(submitted in 2004). The properties of both the bituminous coal used as fuel for the

IGCC processes and the lignite used as fuel for lignite fired boiler processes are also

defined in D1.2.2 [172].

The conditions of the ambient air, which is assumed in all simulations in this work,

are shown in table 4­1. For the analysed IGCC processes, the bituminous coal

“Douglas Premium 2” is used as fuel. Table 4­2 shows the chemical composition in form

Table 4-1: Composition and ISO conditions of ambient air used for simulations. The

ambient air composition has been defined by the ENCAP framework SP 6 [171];

the ISO conditions are in accordance with [173].

Parameter Components / vol­%

Nitrogen Oxygen Water Argon Carbon dioxide

Volume fraction 77.30 20.74 1.01 0.92 0.03

Pressure 1.01325 bar

Temperature 15°C

Relative humidity 60%

34 4 .1 Common Framework for Model l ing of Power Genera t ion Cycles

Table 4-2: Composition and lower calorific value of bituminous coal, which is used as fuel

for all analysed IGCC processes (described in section 4.2). The bituminous coal

is assumed to be the coal “Douglas Premium 2” as defined by SP 1 [172].

Fuel data Units Douglas Premium 2 (as received)

Carbon mass­% 66.52

Hydrogen mass­% 3.78

Oxygen mass­% 5.47

Nitrogen mass­% 1.56

Sulphur mass­% 0.52

Ash mass­% 14.15

Moisture mass­% 8.00

Lower calorific value MJ/kg 25.174

Table 4-3: Composition and lower calorific value of raw lignite and dried lignite, which are

used as fuel for all analysed lignite fired process (described in section 4.3). The

raw lignite is assumed to be a German blend as defined by SP 1 in [172].

Fuel data Raw lignite (as received)

Predried lignite

Carbon / mass­% 27.30 52.80

Hydrogen / mass­% 2.00 3.87

Oxygen / mass­% 10.30 19.92

Nitrogen / mass­% 0.40 0.77

Sulphur / mass­% 0.60 1.16

Ash / mass­% 4.90 9.48

Moisture / mass­% 54.50 12.00

Lower calorific value / MJ/kg 9.01 19.7

of the ultimate analysis based on the “as­received” coal and its lower calorific value8.

For the ultimate analysis, the percentage of carbon, hydrogen, oxygen, sulphur, and

nitrogen are determined. “As­received” indicates that the ultimate analysis includes

moisture and coal ash.

The ultimate analysis of the lignite which is used as fuel for the lignite fired

processes is given in table 4­3. The raw lignite has a low calorific value because it

contains more than 54% moisture. In contrast to hard or bituminous coal, the moisture

of lignite is not surface moisture, but mostly capillary moisture [175]; therefore the

8 According to DIN EN ISO 6976 [174] the lower calorific value is equivalent to the lower heating

value.

4 Analysed Power Genera t ion Processes 35

raw lignite requires pre­drying before it is combusted in the boiler. By means of pre­

drying, the moisture content is reduced by more than 40 percentage points down to

12% by mass, see table 4­3. The required steam is extracted from the intermediate

pressure steam turbine.

Different technologies are applied for lignite drying such as mechanical or thermal

pre­drying, and atmospheric or pressurised fluidised­bed pre­drying [176]. The various

technologies differ in their specific energy expenditure for the drying process. In this

work, atmospheric fine grain fluidised­bed technology is considered for pre­drying of

the raw lignite since it requires the lowest specific energy consumption. This

technology was developed by RWE Power AG (formerly: RWE Rheinbraun AG)

[177]. Atmospheric fluidised­bed technology is close to implementation on a large

scale in power generation processes and has been intensively studied by RWE Power

AG and others [176­180]. Buschsieweke [176] gives an overview of the different

technologies and investigated the drying of lignite by pressurised fluidised­bed

technology.

Table 4-4: Condition of the live steam and reheat of the steam cycle as part of a combined

cycle of an IGCC process.

Property Live steam Reheat

Pressure / bar 125 30

Temperature / °C 560 560

All analysed processes include a steam cycle. The steam condition in the steam

cycle as part of a combined cycle are shown in table 4­4. The configuration of the

steam cycle of the combined cycle has been defined by [171]. The heat recovery steam

generator (HRSG) considered is a triple­pressure HRSG with reheat. The pressure

levels after the feed water are:

High pressure: 125 bar

Intermediate pressure: 30 bar

Low pressure: 4.5 bar

A triple­pressure HRSG consists of a large number of single heat exchangers – an

economiser, evaporator and superheater. In each heat exchanger a pressure drop of 3%

in relation to the inlet pressure on the primary side9 is assumed. In addition, for pipes

9 The steam side is considered as the primary side in all heat exchangers in an HRSG. The exhaust gas

side is consequently termed the secondary side.

36 4 .2 In tegra ted Gasi f icat ion Combined Cycles ( IGCC)

and valves – from the superheater to the appropriate steam turbine – the following

pressure drops are considered for the different pressure levels:

High pressure: 7%

Intermediate pressure: 10%

Low pressure: 12%

The conditions of the live steam and the reheat of the supercritical steam cycle of

the lignite fired boiler processes are given in table 4­5. Since the reference case for the

lignite fired boiler process is taken from [181], all parameters are taken from the

process flow sheet of that deliverable. The process layout is described in detail in

section 4.3.

Table 4-5: Condition of the live steam and the reheat of the steam cycle as part of the lignite

fired boiler process.

Property Unit Live steam Reheat

Pressure bar 280 65

Temperature °C 600 620

The boundary conditions mentioned in this section are applied to all simulations in

the present work. Further details are given in chapter 4, where the different processes

are described; and in chapter 5 where the essential subprocesses are explained in

detail.

4.2 Integrated Gasification Combined Cycles (IGCC)

4.2.1 IGCC process without CO2 Capture

The IGCC process without capture of CO2 is used as a reference process for the

different IGCC configurations, which have been analysed here. The schematic process

layout of the IGCC process without CO2 capture is shown in figure 4­1. In general, an

IGCC process consists of four main subprocesses: gasification, syngas conditioning,

cryogenic ASU (Air Separation Unit), and a combined cycle. These subprocesses

interact with each other by means of heat and mass transfer. It is therefore necessary

that they are optimally integrated to avoid exergy losses due to this heat and mass

transfer. Bituminous coal is grained before it is transported via a lock­hopper to the

gasifier. The coal is gasified in the presence of technically pure oxygen (95 mol­%) to

achieve a high gasification temperature and thus a high carbon conversion ratio.

Unconverted carbon leaves the gasifier in the form of slag and fly ash, and represents a

4 Analysed Power Genera t ion Processes 37

loss in efficiency to the overall process. Besides oxygen and coal, some intermediate

pressure (IP) steam is fed to the gasifier. The steam acts as oxidiser to increase the

carbon conversion ratio and also influence the composition of the raw gas generated.

The IP steam is extracted from the steam turbine from the water/steam cycle.

After gasification, the raw gas is quenched with recirculated and cleaned synthesis

gas (syngas). Due to quenching the temperature is reduced from 1300°C to 900°C.

Further cooling of the syngas down to about 450°C is carried out in the syngas cooler,

see figure 4­1, by means of IP and HP steam generation. The IP and HP feed water is

used for this purpose, and after steam generation is returned to the HRSG. The syngas

leaves the cooler and, after heat exchange with combustible syngas, is cleaned in

various steps. First, fly ash is removed in two sub sequential Venturi­washer. The fly

ash removal is followed by COS­Hydrolysis, where carbon oxysulphide and steam are

converted to carbon dioxide and hydrogen sulphide. The steam is also taken from the

combined cycle. Before desulphurisation the syngas needs to be cooled to 45°C

because the desulphurisation uses chemical absorption to the separate H2S from the

syngas. Between desulphurisation and humidification heat recovery takes places. The

desulphurised syngas is heated to 120°C before humidification takes place. Beneficial

effects of humidification are that the mass flow rate is increased, which consequently

increases the power output of the gas turbine, and also that the added steam results in

lower NOx emissions of the gas turbine. After humidification and internal heat

recovery, the syngas is diluted with nitrogen, the nitrogen being a by­product of the

ASU. Since the nitrogen is delivered at atmospheric pressure, additional compression

is necessary before the nitrogen can be mixed with the syngas. The mixing ratio of

syngas to nitrogen is approximately 0.95 kg N2 / kg syngas. Due to the large amount of

nitrogen, the compression requires a considerable amount of electrical power. The

power output of the overall IGCC process is lowered by about 3.3 percentage points.

Carbon monoxide and hydrogen are the components which determine the calorific

value of the syngas. The mixing with nitrogen reduces the fraction of CO and H2 from

72 mol­% to 44 mol­%. About 75% of the calorific value is provided by the remaining

carbon monoxide. The other 25% is provided by the hydrogen. The lower calorific

value of the combusted syngas is 4.94 MJ/kg. The syngas is combusted with air in the

gas turbine. The pressure ratio of the compressor is assumed to be 17. The

combustorexit temperature is defined to be 1425°C, but before expansion in the

turbine the hot gas is completely mixed with the cooling air which has been extracted

after the compressor. The model of the gas turbine is described in detail in section 5.1.

The exit temperature of the gas turbine is approximately 584°C before the exhaust gas

enters the HRSG. The exhaust gas leaves the HRSG at approximately 87°C. The live

38 4 .2 In tegra ted Gasi f icat ion Combined Cycles ( IGCC)

Fig

ure

4-1

: Si

mpl

ifie

d fl

ow s

heet

of

the

IGC

C r

efer

ence

pro

cess

wit

hout

CO

2 ca

ptur

e an

d no

int

egra

tion

of

the

ASU

(IG

CC

-RE

F).

4 Analysed Power Genera t ion Processes 39

steam conditions are 560°C and 110 bar, see table 4­4. The condenser pressure is

defined to be 48 mbar. After expansion in the HP steam turbine, the steam is reheated

to 560°C. As mentioned before, water and steam are exchanged in various sup­

processes of the IGCC process. The largest mass exchange is the HP and IP feed water

which is used as coolant in the syngas cooler of the gasifier, although it is fed back to

the HRSG as saturated steam. Furthermore, IP steam is extracted for the gasifier, the

humidification of the syngas and for regeneration purposes in the desulphurisation

unit.

The IGCC process described does not possess CO2 capture capability. This process

is used as a references process against which other IGCC configurations can be

compared and in the following is referred to as IGCC­REF. In the following further

IGCC configurations are presented and their features are also described in comparison

to this reference process. Simulation results of the IGCC­REF process and the other

configurations are presented in chapter 7.1.

Figure 4­2 shows the schematic layout of the second IGCC configuration investigated.

It is an IGCC process also without capture of CO2 but with the cryogenic ASU fully

integrated into the gas turbine cycle. In the following, the process is called IGCC­

REF­ASU. Full integration of the ASU into the gas turbine cycle in this context means

that 100% of the required air mass stream is extracted from the gas turbine

compressor. The air is extracted at 15 bar with a temperature of approximately 385°C.

Before the air stream is fed to the ASU it is cooled in the HRSG. Therefore the heat

exchange with the water/steam cycle takes place in three additional heat exchanger.

Afterwards the air enters the ASU at 100°C. Hence, the integration of the ASU has

various impacts, not only with respect to the process layout but also to investment

costs and the operational behaviour of the components. On the one hand the external

air compressor is no longer required, but on the other various hardware changes are

necessary due to the integration of the ASU. First the compressor of the gas turbine no

longer matches the size of the gas turbine because the inlet mass flow rate of the

compressor is increased by about 22% (compared to the IGCC­REF cycle). Secondly,

as mentioned above, further heat exchangers are needed to recover the heat from the

extracted air stream.

From an operational point of view it should be mentioned that for start­up, shut

down and even in case of unforeseen failures an additional oxygen supply would be

required. Different concepts could be considered to cover this aspect. Either an

additional air compressor could be installed which would lead to a high redundancy

and would increase the investment costs significantly. Another possibility might be to

40 4 .2 In tegra ted Gasi f icat ion Combined Cycles ( IGCC)

Fig

ure

4-2

: Si

mpl

ifie

d fl

ow s

heet

of

the

IGC

C r

efer

ence

pro

cess

wit

hout

CO

2 c

aptu

re b

ut 1

00 %

int

egra

tion

of

the

ASU

(IG

CC

-RE

F-A

SU

).

4 Analysed Power Genera t ion Processes 41

use storage tanks to store oxygen. If the ASU produces slightly more oxygen than

required the tanks could be charged during regular operation of the power plant.

Apart from the changes described to the gas turbine the remaining components are

not modified. The heat input in the gasifier is kept constant (coal mass flow rate:

43.21 kg/s). Thus, the raw gas production is not affected. After cleaning and

conditioning the syngas, it is humidified with the same amount of IP steam as for the

IGCC­REF cycle (extracted from the HRSG). The humidification is followed by

internal heat exchange with syngas leaving the syngas cooler of the gasifier. Before the

syngas is combusted it is diluted with nitrogen. Again, the amount of nitrogen equates

to the IGCC­REF process. Although the ASU is fully integrated into the gas turbine,

which lowers the specific energy for the oxygen production, the nitrogen that is mixed

with the syngas is delivered from the ASU at identical conditions (a temperature of

15°C and atmospheric pressure). Therefore a nitrogen compressor is also required in

this configuration. The mass flow of the fuel and its composition (CO: 31 mol­%, H2:

13 mol­%) stays as the reference cycle. Although the combined cycle is significantly

affected, the power output of the steam turbine increases only slightly due to the

additional heat of the extracted air for the ASU. The operating condition of the gas

turbine is consistent with the reference cycle: the pressure ratio of the compressor is

17, the combustor exit temperature is defined to be 1425°C. Thus, the exhaust

temperature is the same as for the reference cycle. Consequently the live steam and

reheat conditions remain unchanged, see table 4­4. The condenser pressure is set to

48 mbar. The pressure drop in the heat exchangers of the HRSG is defined as

percentage of the inlet pressure according to the introductory section of this chapter

(see page 33).

The integration of the additional heat extracted from the air by the ASU leads to the

requirement of three further heat exchanger that are integrated, depending on the

temperature level, at different locations within the HRSG. The efficient utilisation of

this heat requires a high number of heat exchangers to minimise the mean temperature

differences. In general, low­temperature heat can be utilised up to a certain amount. In

the case of the integrated ASU, the additional heat, which is integrated into the

water/steam cycle leads to an increase of the exhaust gas temperature of about 50 K.

This configuration has been investigated to understand how the reduced

expenditure of energy for the ASU impacts the net efficiency of the overall cycle. In

comparison to the reference process it is expected that the net efficiency will increase

but the investment costs will be affected adversely due to the additional components.

42 4 .2 In tegra ted Gasi f icat ion Combined Cycles ( IGCC)

In the following, various IGCC configurations with CO2 capture are described. The

two configurations, IGCC­REF and IGCC­REF­ASU, presented in this section are

used as reference cycles. Later in the discussion of the simulation results (section 7.1),

the processes utilising CO2 capture are compared with those without separation of CO2

to analyse which sub­processes are mainly affected by the additional steps of the

separation process. Furthermore, crucial parameters will be identified which cause the

largest expenditures of energy and, thus, reduce the efficiency mostly.

4.2.2 IGCC process with Cryogenic Air Separation Unit and CO2 Capture

In this section the IGCC process with cryogenic ASU and CO2 capture (IGCC­

CAP) is presented. A simplified flow sheet of the IGCC­CAP process is given in

figure 4­3. The separation of the CO2 is carried out by means of physical absorption.

Physical absorption is a well­known separation technology not only for separation of

CO2 but also for H2S.

As far as the gasification and cleaning of the raw gas are concerned, the process

remains unchanged if CO2 capture is applied to the process or not. Since the syngas

consists mainly of carbon monoxide and hydrogen after desulphurisation, an additional

process step is required before the separation of carbon dioxide can take place. After

desulphurisation and internal heat recovery, a CO­shift reaction is required. The

following reaction is carried out in a two stage process;

2 2 2 CO H O CO H (4­1)

The CO­shift reaction is exothermic (hR = 44.477 kJ/mol) [95]. The water which

is required for the reaction itself is supplied in the form of slightly superheated IP

steam extracted from the HRSG. Since the reaction is exothermic and additional heat

is delivered by the steam, the generated heat is used to produce some IP steam.

Therefore IP feed water taken from the HRSG is evaporated and sent back to the

HRSG as saturated steam.

During the CO­shift reaction, 93.5% of the carbon monoxide is converted to carbon

dioxide. The syngas is further cooled to 40°C before the CO2 separation takes place.

Due to the low temperature, the water content is reduced from 15 mol­% to nearly zero

(< 0.2 mol­%). The separation of carbon dioxide is facilitated by physical absorption.

Approximately 98% of the incoming CO2 is captured in the separation process, so that

the syngas leaving the separation unit consists of H2: 85, N2: 8 and CO: 4 (mol­%).

After the CO2 separation, the process steps are the same as that of the IGCC process with­

4 Analysed Power Genera t ion Processes 43

Fig

ure

4-3

: Si

mpl

ifie

d fl

ow s

heet

of

the

IGC

C p

roce

ss w

ith

CO

2 c

aptu

re b

ut n

o in

tegr

atio

n of

the

ASU

(IG

CC

-CA

P).

44 4 .2 In tegra ted Gasi f icat ion Combined Cycles ( IGCC)

out CO2 capture. By means of internal heat exchange with raw gas leaving the syngas

cooler, the combustible syngas is reheated to 350°C before dilution with nitrogen takes

place. Unlike the IGCC­REF process, after mixing with nitrogen, the fuel gas contains

about 45 mol­% hydrogen. The difference in the combusted syngas is caused by the

separation of carbon in the form of CO2. For the IGCC­CAP process about 95% of the

heating value is delivered by the hydrogen. The calorific value of the combusted

syngas is 6.93 MJ/kg. The configuration of the combined cycle is identical to that of

the IGCC­REF process. The same model for the generic cooled gas turbine is used, see

section 5.1. The key operating parameters of the gas turbine are the same as for the

IGCC­REF cycle (compressor pressure ratio: 17; combustor exit temperature:

1425°C). Consequently, also the live steam conditions remain unchanged, see table

4­4. The condenser pressure is defined to be 48 mbar.

The difference in the exhaust temperature is caused by interaction of the HRSG with

the CO­shift reaction. This affects the steam mass flow rate produced in the LP and IP

evaporators of the HRSG. Therefore the stack temperature changes slightly compared

to the IGCC­REF process. Similar to the reference process (IGCC­REF), the IGCC­

CAP process has been investigated in two different configurations. These

configurations differ only in the way that the cryogenic ASU is integrated into the gas

turbine cycle. In the case of the first configuration (IGCC­CAP), all air required for the

ASU is delivered by an external air compressor, see figure 4­3. The second

arrangement, where the ASU is fully integrated, 100% of the air is extracted from the

compressor of the gas turbine. The IGCC process with CO2 capture and fully

integrated ASU is named IGCC­CAP­ASU. The schematic flow sheet of the IGCC­

CAP­ASU process is given in figure 4­4.

In the IGCC process itself, the gasification, the syngas treatment (including CO­

shift reaction and CO2 capture) and the combined cycle remain unchanged when the

ASU is integrated into the gas turbine process. The main difference between the

configurations with and without integration of the ASU is that the design of the gas

turbine will change significantly because the mass flow rate in the compressor will

vary non­proportionally compared with the mass flow rate of the turbine. The mass

flow rate through the compressor is approximately 140 kg/s larger than the exhaust

gasof the turbine since this amount of air is required by the ASU. The size of both

turbomachines – compressor and turbine – will therefore not fit to a conventional

layout of a standard gas turbine. Furthermore, additional heat exchangers are required

to integrate the heat produced in the steam cycle to preheat some feed water. The

advantage of the integration of the ASU is a lower specific energy consumption, how­

4 Analysed Power Genera t ion Processes 45

Fig

ure

4-4

: Si

mpl

ifie

d fl

ow s

heet

of

the

IGC

C p

roce

ss w

ith

CO

2 c

aptu

re a

nd w

ith

100

% i

nteg

rati

on o

f th

e A

SU (

IGC

C-C

AP

-ASU

).

46 4 .2 In tegra ted Gasi f icat ion Combined Cycles ( IGCC)

ever this provides a larger amount of low temperature heat which can only be

inefficiently utilised in the steam cycle. This leads to a higher stack temperature of the

exhaust gas because not all of the low temperature heat can be used. In addition, the

large number of heat exchangers have an adverse effect on the capital costs.

These two configurations, IGCC­CAP and IGCC­CAP­ASU, are used as reference

cycles with CO2 capture. The following two configurations with incorporated

membranes and CO2 capture, which are described in the next two sections, are

compared to those processes with a more mature CO2 capture technology. If novel or

innovative processes such as membrane­based processes are analysed, it is reasonable

to compare them not only with a reference cycle without CO2 capture but also to

evaluate them against more mature processes. This kind of comparison may better

illustrate where the novel concept has advantages or drawbacks with respect to the

more mature CO2 capture technology.

4.2.3 IGCC process with Integrated Oxygen Transport Membrane (OTM)

and CO2 Capture

The IGCC process with integrated Oxygen Transport Membrane (OTM) and CO2

capture is schematically presented in figure 4­5; hereafter called IGCC­OTM. The

membrane reactor is located between two combustion chambers which are again

between the compressor and turbine of the gas­turbine. The main goal of the

membrane reactor is to separate oxygen from pre­heated air. The oxygen is

subsequently used as an oxidant in the entrained flow gasifier. Steam extracted from

the water/steam cycle is used as sweep gas in the membrane reactor.

The integration of an OTM reactor into a gas turbine cycle is similar to the AZEP

concept. In the AZEP concept the OTM reactor is also integrated into the gas turbine,

but the sweep side it is fed with a mixture of methane, water and carbon dioxide. The

feed side of the reactor is fed with air from the gas turbine compressor. The OTM

reactor is comprised of different sections. In the first section, methane is first

combusted with the oxygen, which is transferred through the membrane. On the one

hand the combustion causes heating of the whole reactor. On the other hand the

transferred oxygen is consumed which reduces the oxygen partial pressure on the

sweep side of the membrane reactor. Both the higher operating temperature and larger

pressure ratio of the oxygen partial pressure increase the oxygen permeation rate

through the membrane. For further description of the AZEP concept see [97­101]. The

main reason for the introduction of the OTM reactor in the IGCC­OTM process is the

4 Analysed Power Genera t ion Processes 47

substitution of the cryogenic ASU. The oxygen which is required for the gasification is

produced by separation of oxygen out of the preheated air stream; therefore the

cryogenic ASU is not required. Since the ASU is replaced by the OTM reactor, there is

the potential to reduce energy consumption and thus to improve efficiency.

In comparison to the IGCC­CAP process the gasification and syngas cleaning and

treatment sections remain unchanged. Therefore, the composition of the combusted

syngas does not change ( H2: 45, N2: 52 and CO: 2 mol­%). After dilution with

nitrogen, the synthesis gas is distributed to both combustion chambers. Air from the

compressor is pre­heated to 900°C before entering the OTM reactor. The oxygen

concentration is reduced by around 5 mol­% during the first combustion. The high

temperature of the air is required because the OTM reactor operates only above 750°C.

The material of the reactor is a MIEC10 membrane. This ceramic material needs high

temperatures to be ionically and electronically conductive. Since the permeation rate

increases with temperature, 900°C has been chosen as the inlet temperature of the feed

stream (pre­heated air). Only a part of the oxygen permeates through the membrane

because a certain amount of oxygen is required for the second combustion chamber.

The air stream after the membrane reactor still contains approximately 10­11 mol­% of

oxygen.

Intermediate pressure steam is used as sweep gas on the permeate side of the OTM

reactor. Sweep gas is used on the permeate side of the OTM reactor to lower the

partial pressure of the oxygen which permeates through the membrane. Steam is used

from the water/steam cycle because it does not chemically react with oxygen and is

neutral to the membrane material, meaning it does not impact the oxygen permeation

rate. Furthermore, the mixture of steam and oxygen can easily be separated after the

OTM reactor by means of cooling and condensation. Before the steam enters the OTM

reactor it is internally heated by the steam/oxygen stream leaving the reactor. The

heating is necessary to reduce the heat exchange from the feed stream to the sweep

stream. After internal heat exchange the steam/oxygen mixture has a temperature of

approximately 500°C. The heat of the mixture is utilised in the HRSG. Due to a large

mass flow and a high temperature several heat exchangers are required to use the heat

as efficiently as possible in the HRSG.

The retentate stream leaving the OTM reactor is fed to the second combustion

chamber of the gas turbine. In the second combustor the oxygen­depleted air is heated

to the same hot gas temperature of 1425°C as in the other IGCC processes. The reasons

10 Mixed Ionic Electronic Conducting

48 4 .2 In tegra ted Gasi f icat ion Combined Cycles ( IGCC)

Fig

ure

4-5

: Si

mpl

ifie

d fl

ow s

heet

of

the

IGC

C p

roce

ss w

ith

CO

2 c

aptu

re a

nd i

nteg

rate

d O

TM

rea

ctor

(IG

CC

-OT

M).

4 Analysed Power Genera t ion Processes 49

reasons why the second combustor is required are, firstly, that the power output of the

gas turbine would be dramatically less if there were no additional combustor.

Secondly, the exhaust temperature of the gas turbine would be much lower, so that the

live steam temperature would need to be reduced. This would also lower the efficiency

of the steam cycle. Similar to the other IGCC configurations, the exhaust gas of the

gas turbine flows through a triple pressure HRSG before it is fed to the stack. The

configuration of the HRSG changes slightly because of the additional heat exchangers

that are required to recover the heat transferred in the OTM reactor from the feed to

the sweep side of the reactor.

Since the OTM reactor consists of a large number of small tubes, a pressure loss

occurs on both sides of the OTM reactor (see description of the OTM reactor in

chapter 6). The pressure loss on the retentate side of the OTM reactor impacts the

power output of the gas turbine cycle adversely because it reduces the pressure ratio in

the turbine. The heat which is transferred in the OTM reactor lowers the temperature

of the retentate stream on the feed side of the reactor; therefore, more fuel is needed to

achieve the same hot gas temperature. This lowers the thermal efficiency of the gas

turbine. Since different configurations of the OTM reactor have been analysed, the

outlet temperature of the retentate stream varies from 850°C to 880°C. The hot gas

temperature remains constant compared to the other IGCC configurations at 1425°C.

Therefore also the exit temperature of the turbine changes only slightly (less than 5 K).

The performance of the OTM reactor impacts on the power output and efficiency of

the overall IGCC process. For example, the mass flow and the pressure level of the

steam which is extracted from the steam turbine determines the reduction of the steam

turbine power output. Furthermore, the ratio of oxygen partial pressure determines the

required surface of the oxygen transport membrane and hence the pressure loss in the

membrane reactor. Due to the way that the OTM reactor is integrated into the gas

turbine cycle, the conditions on the feed side of the reactor have been kept constant.

Different configurations of OTM reactor regarding the conditions (mass flow and

pressure) of the sweep stream are investigated. Details of the OTM reactor can be

found in section 6.5.1. The results and the impact on the IGCC­OTM process are

presented and discussed in section 7.1.

The gasification process itself does not change in comparison to the IGCC­CAP

process. The syngas treatment ends with desulphurisation followed by internal heat

exchange and then the CO­shift reaction. Subsequently, water is separated from the

syngas stream in such a way that the syngas stream is fully saturated before CO2

capture takes place. The separation of carbon dioxide is achieved by means of physical

50 4 .2 In tegra ted Gasi f icat ion Combined Cycles ( IGCC)

absorption. As mentioned earlier, the separated CO2 is dried by a multistage,

intercooled compression. The syngas leaving the CO2 capture process possesses a high

content of hydrogen (H2: 85, N2: 8, CO: 4 mol­%). Before the syngas is distributed to

both combustion chambers it is first heated by internal heat exchange with raw syngas

leaving the gasifier, and second diluted with nitrogen. Due to the dilution, the amount

of hydrogen is reduced to 45 mol­%. The composition of the fuel gas is identical to

that from the IGCC­CAP process. The same holds true for the operating parameter of

the gas turbine (pressure ratio and hot gas temperature, see chapter 5.1, page 69 et

sqq.) and the configuration of the bottoming steam cycle (live steam conditions, reheat

and condenser pressure, see table 4­4).

In this work, the IGCC­OTM process has been analysed for base load condition,

but it should be mentioned that the following aspects need to be considered if such a

configuration were realised. Although the OTM reactor is designed in such a way that

the cryogenic ASU is eliminated, but for transient operation (start­up, shut­down and

unforeseen events) of the overall process, it will be necessary to provide an additional

oxygen source. Either oxygen tanks need to be available which are able to deliver the

required oxygen to the gasifier until the gas turbine and especially the OTM reactor

operate at base load or a small cryogenic ASU could be provided, which could

produce the required oxygen until full oxygen production of the OTM reactor is

achieved. A small additional ASU would be beneficial since the nitrogen, which is

required for dilution of the syngas could be produced by this additional ASU.

Otherwise, the required nitrogen would also need to be provided by tanks. This would

in addition affect the operating costs of the overall process adversely.

Apart from the OTM reactor as part of the gas turbine cycle, the complexity of the

IGCC­OTM process is increased compared to previous configurations. The amount of

steam which is extracted from the water/steam cycle and in particular the increased

number of the extraction locations leads to a higher complexity of the HRSG. In

addition to the exchange with the gasifier, the CO­shift reactor and the

desulphurisation unit, IP or LP steam is extracted to be used as sweep gas in the OTM

reactor. The heat of the steam/oxygen mixture leaving the reactor is also utilised in the

HRSG. Therefore the utilisation of all of the additional heat is a compromise between

efficient usage and complexity of the heat exchanger in the HRSG. Furthermore, more

low temperature heat is available, which cannot be utilised in an optimal way. Hence

the temperature of the exhaust gas after the HRSG increases by around 50°C compared

to the IGCC­CAP process.

4 Analysed Power Genera t ion Processes 51

4.2.4 IGCC process with Hydrogen-selective Membrane and CO2 Capture

Another membrane­based IGCC presented in this work is an IGCC process with an

integrated hydrogen­selective membrane. Such a concept is schematically shown in

figure 4­6. In the following this concept is named the IGCC­H2 cycle. The advantage

of the IGCC­H2 cycle with the integrated hydrogen­selective membrane is that no CO2

capture unit is required because the hydrogen is separated in the membrane reactor and

the remaining syngas – mainly carbon dioxide and water – can be separated by means

of condensing. The syngas treatment is the same as that for the IGCC­CAP and IGCC­

OTM processes until desulphurisation and internal heat exchange with raw syngas

leaving the gasifier. Therefore the composition after CO­shift reaction is identical

(CO2: 36, H2: 50; N2: 5, H2O: 4; CO: 3 mol­%) for all three processes. The first

difference in the syngas treatment occurs after the first heating to 350°C. A second

internal heat exchange with syngas leaving the membrane reactor takes place, so that

the syngas is further heated up to operating temperature of the membrane reactor at

600°C. Then the syngas enters the membrane reactor where, due to its high selectivity,

only hydrogen is separated from the syngas. On the sweep side of the membrane

reactor pressurised nitrogen is used as sweep gas. The nitrogen is a by­product of the

cryogenic ASU which produces the oxygen required for the gasifier. In contrast to the

OTM reactor nitrogen instead of steam is used as sweep gas because nitrogen is also

neutral to the membrane material and, most importantly, the separated hydrogen needs

dilution with nitrogen before it can be combusted in a conventional gas turbine.

The remaining syngas which leaves the membrane reactor on the feed side still

contains small amounts of combustible products. Some carbon monoxide remains in

the syngas because the conversion ratio of the CO­shift reaction is 93%. In addition it

is assumed that only 99% of the hydrogen is separated in the hydrogen­selective

membrane reactor. Thus, supplementary firing is required to convert those combustible

components to water and carbon dioxide. Since the syngas contains no oxygen, oxygen

needs also to be supplied to this supplementary burner. Only as much oxygen as is

required for complete stoichiometric combustion is provided. The oxygen is also

supplied by the cryogenic ASU. The burned syngas leaving the supplementary burner

has a temperature of about 1100°C. Due to the combustion in the supplementary

burner the stream leaving the supplementary burner contains mainly carbon dioxide

and water (CO2: 62, H2O: 29; N2: 8, Ar: 1 mol­%). Therefore the stream should not be

designated as syngas any longer but rather as the CO2­rich stream. The heat of the

CO2­rich stream is recovered in three different ways. First, an internal heat exchange

with the syngas which is fed to a membrane reactor takes place. After that some of the

52 4 .2 In tegra ted Gasi f icat ion Combined Cycles ( IGCC)

Fig

ure

4-6

: Si

mpl

ifie

d fl

ow s

heet

of

the

IGC

C p

roce

ss w

ith

CO

2 c

aptu

re a

nd s

elec

tive

hyd

roge

n m

embr

ane

rea

ctor

(IG

CC

-H2)

.

4 Analysed Power Genera t ion Processes 53

heat is utilised by heating up the mixture of hydrogen and nitrogen before further

dilution with additional nitrogen. Finally, the low temperature heat is utilised by means

of additional heat exchangers in the water/steam cycle.

After cooling the CO2­rich stream to 173°C in the HRSG and further to 22°C by

means of cooling water, the CO2­rich stream is compressed to 110 bar. The

compression takes place in four intercooled stages and consequently water is

condensed out after each stage. Since the CO2­rich stream is the remaining stream of

the former syngas, the amount of nitrogen is higher in comparison to the other IGCC

processes using physical absorption to separate the CO2. After compression the

concentration of CO2 is around 88 mol­%; nitrogen and argon complete the stream

with 11 and 1 mol­% respectively. This is also the composition of the product stream

of the separated CO2. No further treatment of the stream is considered in this work.

The sweep stream leaving the hydrogen­selective membrane reactor contains about

60 mol­% hydrogen. The remainder is pure nitrogen. Since the mixture is at a lower

pressure than the operating pressure of the gas turbine it needs to be compressed

before it can be used as fuel in the gas turbine. The lower pressure is necessary to

ensure and increase the difference of the hydrogen partial pressure between feed and

sweep stream across the membrane reactor. The difference in hydrogen partial

pressure is the driving force for the transport mechanism through the membrane. The

feed pressure – operating pressure of the gasifier minus the pressure loss occurring

during the syngas treatment – is assumed to be 25 bar. The pressure of the sweep

stream is set to 5 bar. The mixture of hydrogen and nitrogen is compressed to 25 bar

before it is further diluted with additional nitrogen. The amount of hydrogen is reduced

to 45 mol­% before the fuel gas is combusted in the conventional gas turbine. The

dilution is required to compensate the change of combustion behaviour of the

hydrogen­rich fuel. The combustion behaviour changes for hydrogen combustion –

compared to hydrocarbons – because, both, the chemical reactivity and the flame

speed of hydrogen are much higher. Furthermore, a higher adiabatic flame temperature

results in higher NOx emissions. Considering these implications of hydrogen

combustion and assuming that the gas turbine is based on today’s GT technology, the

amount of hydrogen should be kept below 50 mol­% to ensure stable combustion.

The configuration of the gas turbine is identical to the other investigated IGCC

processes. Besides the fact that the fuel gas composition changes (H2: 45; N2: 55 mol­

%), the configuration of the gas turbine remains unchanged: the pressure ratio of the

compressor is 17 and the combustor exit temperature is 1425°C, see section 5.1. The

calorific value of the fuel gas is, at 6.78 MJ/kg, slightly lower than that for the IGCC­

54 4 .3 Ligni te Fi red Boi ler Cycles

OTM process. Furthermore, the configuration of the steam cycle remains constant:

same live steam conditions (see table 4­4) and a condenser pressure of 48 mbar.

The advantage of the IGCC­H2 process of not requiring a CO2 capture unit

unfortunately comes with the drawback that an supplementary burner is needed to

combust the remaining combustible components of the former syngas stream. The high

temperature due to this combustion is not usable efficiently. The utilisation of the heat

produced can only be a compromise between complexity (numbers of additional heat

exchangers) and acceptable temperature differences to reduce the loss of exergy in

each heat transfer. The effect of this aspect on the whole IGCC process will be further

discussed in section 7.1 and in chapter 8.

4.3 Lignite Fired Boiler Cycles

Various configurations of a lignite fired steam power plant have been investigated.

The abbreviation “LFB” (Lignite Fired Boiler) cycle is introduced to referring to the

different configurations of the lignite fired boiler cycles. First an air driven lignite fired

boiler process without CO2 capture serves as a reference case for the following

investigations. The reference case is used as a basis for a semi­closed oxyfuel boiler

process, where oxygen (95 mol­%) is used as the oxidant, supplied by a cryogenic air

separation unit (ASU). The oxyfuel boiler process serves again as a base case for an

oxyfuel boiler process whereby the ASU is substituted by an oxygen transport

membrane (OTM) reactor to separate the required oxygen for the boiler from

preheated air. In this case the OTM reactor is integrated in a gas turbine cycle. The

reference case and the oxyfuel boiler with ASU have been investigated within ENCAP

and are part of the ENCAP deliverables D3.3.3.2 [181] and D3.3.2 [182], respectively.

4.3.1 Air Driven Boiler without CO2 capture

A simplified flow sheet of the air driven, lignite fired boiler process is shown in

figure 4­7. In the following this cycle will be referred to as LFB­AIR cycle. The

lignite, which is burned in a supercritical once through boiler, is dried by means of

intermediate pressure steam extracted from the steam cycle, before it is fed together

with preheated ambient air to the boiler. In the supercritical boiler, feed water is heated

from about 300°C to 600°C at an inlet pressure of 310 bar.

Furthermore, steam coming from the high pressure steam turbine is reheated up to

620°C at an inlet pressure of 64 bar. The exhaust gas leaves the once through boiler

4 Analysed Power Genera t ion Processes 55

with a temperature of 380°C and is further cooled to 170°C in the primary air pre­

heater. In order to achieve a low stack temperature of the exhaust gas flow, boiler feed

water is heated from 90°C to 130°C, which results in a stack temperature of 110°C for

the exhaust gas flow.

The live steam enters the single high pressure steam turbine at 600°C, at a pressure

of 280 bar. After the intermediate pressure steam is reheated to 620°C, the steam

enters one double­flow intermediate pressure steam turbine. The expansion to the

condenser pressure of 48 mbar takes place in two double­flow low pressure steam

turbines. After complete condensation, the condensate pump elevates the pressure up

to 32 bar. In the first six feed water heaters, a pressure drop of about 15 bar occurs, so

that the de­aerator operates at 17 bar. Heat for the de­aerator is supplied by

intermediate pressure steam. The boiler feed water leaves the de­aerator at nearly

saturation temperature before it is compressed in the feed pump to 334 bar. The feed

pump is driven by a steam turbine. A small steam turbine, which operates with

intermediate pressure steam delivers about 30 MW to drive the feed pump. The last

four feed water heaters raise the water temperature from 210°C to 300°C, at which

temperature the feed water enters the once through boiler. The pressure drop in the last

four feed water heaters and the boiler is more than 50 bar, so that the high pressure

steam enters the turbine with 280 bar.

For the ambient air, ISO conditions are assumed (see table 4­1). The raw lignite coal

has a lower calorific value of 9.01 MJ/kg, see table 4­3. After drying with intermediate

pressure steam, the lower calorific value of the lignite is increased to19.7 MJ/kg (see

table 4­3). The four steam turbines are considered to operate in single­shaft

configuration. The heat of the cooling water is dissipated by means of a natural­

draught wet cooling tower. All four steam turbines (1x HP, 1x IP, 2x LP) are stage­

wise modelled. For each stage an isentropic efficiency is assumed. The HP part

consists of 2 stages, the IP part is modelled with 4 stages and finally the LP part has 3

stages. The isentropic efficiencies for the those stages are given in table 4­6.

In the following, more details on the four steam turbines of LFB­AIR cycle are

presented. From table 4­7 to table 4­10 the mass flow rates of incoming, outgoing, and

the extracted streams of each turbine are given. The high pressure live steam mass

flow is 716.6 kg/s, which is produced in the boiler. The outgoing steam of the low

pressure turbine that is fed to the condenser is only 409.5 kg/s. In total nearly 43% of

the live steam mass flow is extracted from all steam turbines for the purpose of

preheating feed water, driving the feed water pump and drying the raw lignite.

The largest amount of steam is extracted from the intermediate pressure steam turbine,

56 4 .3 Ligni te Fi red Boi ler Cycles

Fig

ure

4-7

: Si

mpl

ifie

d fl

ow s

heet

of

a li

gnit

e fi

red

boil

er r

efer

ence

pro

cess

(L

FB

-AIR

). T

he p

roce

ss i

s ai

r dr

iven

; no

CO

2 c

aptu

re u

tili

sed.

4 Analysed Power Genera t ion Processes 57

about 21% of the live steam mass flow. This is because the intermediate pressure

steam is used to drive the feed water pump, to provide heat for the de­aerator and the

drying of the raw lignite and for two feed water heaters.

Since no CO2 capture takes place in this reference case, the cycle emits specifically

810 g CO2/kW(e). Details on the exhaust gas stream regarding CO2 emissions can be

found in the appendix in table 4­3. Desulphurisation is not included in the simulation.

The main results of the reference case are presented in section 7.2.

Table 4-6: Isentropic efficiencies for each stage of the HP, IP and LP steam turbines.

Isentropic efficiency HP IP LP

First stage 95.0 95.5 94.0

Second stage 95.0 95.5 90.0

Third stage ­­­ 97.5 85.0

Fourth stage ­­­ 96.0 ­­­

Table 4-7: LFB-AIR cycle: HP steam turbine data. The steam conditions at the inlet, the

outlet and the extractions regarding pressure, temperature and mass flows are

given. In addition, the gross power is shown.

Port m / kg/s T / °C P / bar

IN 716.6 599.2 277.0

OUT 683.9 362.5 65.0

First extraction 32.7 394.6 82.0

Total power output MW 284.37

Table 4-8: LFB-AIR cycle: IP steam turbine data. The steam conditions at the inlet, the

outlet and the extractions regarding pressure, temperature and mass flows are

given. In addition, the gross power is shown.

Port m / kg/s T / °C p / bar

IN 630.4 622.5 62.0

OUT 476.3 231.2 4.1

First extraction 37.7 525.9 35.0

Second extraction 30.3 423.8 18.0

Third extraction 86.1 394.6 9.0

Total power output MW 445.05

58 4 .3 Ligni te Fi red Boi ler Cycles

Table 4-9: LFB-AIR cycle: First LP steam turbine data. The steam conditions at the inlet,

the outlet and the extractions regarding pressure, temperature and mass flows

are given. In addition, the gross power is shown.

Port m / kg/s T / °C p / bar

IN 150.7 231.2 4.1

OUT 129.5 32.2 0.05

First extraction 2.7 144.6 1.7

Second extraction 18.5 91.8 0.75

Total power output MW 85.99

Table 4-10: LFB-AIR cycle: Second LP steam turbine data. The steam conditions at the inlet,

the outlet and the extractions regarding pressure, temperature and mass flows

are given. In addition, the gross power is shown.

Port m / kg/s T / °C p / bar

IN 314.4 231.2 4.1

OUT 280.0 32.2 0.05

First extraction 2.7 144.6 1.7

Second Extraction 16.8 72.7 0.35

Third extraction 14.9 54.0 0.15

Total power output MW 190.51

4.3.2 Oxyfuel Boiler with cryogenic ASU and CO2 capture

In the following, the oxyfuel boiler cycle with cryogenic ASU and CO2 capture is

presented. A simplified flow sheet of the oxyfuel lignite fired boiler cycle is shown in

figure 4­8. The oxyfuel lignite fired boiler cycle with CO2 capture is in the following

referred to as LFB­OXY. The lignite, which is burned in a supercritical once­through­

boiler, is dried by means of intermediate pressure steam extracted from the steam

cycle, before it enters the boiler. In this case, technically pure oxygen (95 mol­%)

rather than ambient air is used as oxidant for the combustion process. The oxygen is

produced by a cryogenic ASU. Before the oxygen is fed to the boiler it is also

preheated to 300°C. In addition to the oxygen, part of the flue gas is recycled back to

the boiler in order to control the boiler temperature. If the boiler is air driven, the large

portion of nitrogen in the ambient air acts as inert gas for the combustion process. Due

to the missing nitrogen in the case of the oxyfuel boiler, this role is fulfilled by the flue

gas. Of course, the setup of the water/steam side of the oxyfuel boiler does not change

compared to the air driven boiler. The supercritical feed water enters the once through

4 Analysed Power Genera t ion Processes 59

boiler at 310 bar and 300°C and the superheated live steam leaves the boiler at 600°C.

In addition, steam coming from the high pressure steam turbine is superheated up to

620°C at an inlet pressure of 64 bar. The exhaust gas, which consists mainly of carbon

dioxide and water (CO2: 58, H2O: 33; N2: 5, O2: 2, Ar: 2 mol­%), leaves the once

through boiler at a temperature of 480°C.

Afterwards, the exhaust gas is further cooled to 170°C in the primary air preheater,

before around 15% of the flue gas is recycled, the total flue gas stream requires

cleaning. First, the flue gas is sent through electrostatic particle precipitator (ESP)

followed by flue gas cleaning. In this oxyfuel configuration no special CO2 capture

unit is required but rather the cleaned flue gas is cooled down further close to ambient

temperature. Due to cooling of the flue gas water condenses so that the CO2

concentration increases from 58 to 89% by mole. The rest of the flue gas is nitrogen,

oxygen and argon at 5, 3 and 1 mol­%, respectively, and some remaining water at

2 mol­%.

Before compression of the CO2­rich exhaust stream, inert gases need to be removed

because a large amount of inert gas would increase the power required for CO2

compression. Depending on the purpose of the ‘product’ CO2, it may be necessary to

have a high purity of the CO2. Due to the removal of the inert gases and the multistage

compression of the CO2, where the remaining water is further condensed, the purity of

the compressed carbon dioxide is 97% by mole. The pressure of CO2 after compression

is as for the other cycles 110 bar.

To facilitate comparison, details on the four steam turbines of the lignite fired

oxyfuel boiler are given in the following. From table 4­11 to table 4­14 the mass flow

rates of incoming, outgoing, and the extracted streams of each turbine are given. The

high pressure live steam is 742.7 kg/s, which is produced in the boiler. The outgoing

steam of the low pressure turbine that is fed to the condenser is only 417.1 kg/s. The

steam is extracted from all steam turbines for the purpose of preheating feed water, dry­

Table 4-11: LFB-OXY cycle: HP steam turbine data. The steam conditions at the inlet, the

outlet and the extractions regarding pressure, temperature and mass flows are

given. In addition, the gross power is shown.

Port m / kg/s T / °C P / bar

IN 742.7 599.2 277.0

OUT 708.9 362.5 65.0

First extraction 33.8 394.6 82.0

Total power output MW 294.75

60 4 .3 Ligni te Fi red Boi ler Cycles

Fig

ure

4-8

: Si

mpl

ifie

d fl

ow s

heet

of

a li

gnit

e fi

red

boil

er o

xyfu

el b

oile

r pr

oces

s w

ith

CO

2 c

aptu

re (

LF

B-O

XY

).

4 Analysed Power Genera t ion Processes 61

Table 4-12: LFB-OXY cycle: IP steam turbine data. The steam conditions at the inlet, the

outlet and the extractions regarding pressure, temperature and mass flows are

given. In addition, the gross power is shown.

Port m / kg/s T / °C p / bar

IN 647.8 622.5 62.0

OUT 495.2 231.2 4.1

First extraction 39.0 525.9 35.0

Second extraction 26.8 423.8 18.0

Third extraction 86.8 327.0 9.0

Total power output MW 459.15

Table 4-13: LFB-OXY cycle: first LP steam turbine data. The steam conditions at the inlet,

the outlet and the extractions regarding pressure, temperature and mass flows

are given. In addition, the gross power is shown.

Port m / kg/s T / °C p / bar

IN 154.8 231.2 4.1

OUT 131.3 32.2 0.05

First extraction 4.8 144.6 1.7

Second extraction 18.7 91.8 0.75

Total power output MW 87.52

Table 4-14: LFB-OXY cycle: second LP steam turbine data. The steam conditions at the

inlet, the outlet and the extractions regarding pressure, temperature and mass

flows are given. In addition, the gross power is shown.

Port m / kg/s T / °C p / bar

IN 323.0 231.2 4.1

OUT 285.8 32.2 0.05

First extraction 4.8 144.6 1.7

Second Extraction 17.2 72.7 0.35

Third extraction 15.2 54.0 0.15

Total power output MW 194.78

ing of the raw lignite and for driving the feed water pump.

Since the configuration of the water/steam cycle changes only slightly compared to

the air driven boiler process, the extraction streams relative to the inlet stream do not

change much. As for the air driven boiler the largest amount of steam is extracted from

62 4 .3 Ligni te Fi red Boi ler Cycles

the intermediate pressure steam turbine, about 44%. The absolute numbers of the

steam flows differ from the LFB­AIR cycle because the amount of exhaust gas

changes. Due to this, the amount of steam produced varies from one cycle to the other.

Since CO2 capture takes place in the LFB­OXY cycle specifically 98 g CO2/kW(e) is

emitted by the cycle. Details on the exhaust gas stream regarding CO2 emissions are

given in the appendix in Table A4­3.

4.3.3 Oxyfuel Boiler with Integrated OTM Reactor and CO2 capture

Similar to the IGCC configurations, for the lignite fired boiler process, a membrane

based configuration has been investigated in this work. As for the IGCC processes,

technically pure oxygen (95 mol­%) is required for the combustion of the lignite fired

oxyfuel boiler process. In this case, the OTM reactor provides the oxygen substituting

for the ASU in the previous cycle. The configuration of the lignite fired boiler with

integrated OTM reactor is shown schematically in figure 4­9. In the following the

cycles is referred to as LFB­OTM cycle.

The way the OTM reactor is introduced into the existing lignite fired boiler leads to

a combination of a coal fired steam power plant and a natural gas fired combined cycle

power plant because the OTM is integrated into a gas turbine process. The gas turbine

is required because the air stream needs to be preheated before it can be fed as a feed

stream to the OTM reactor. The OTM reactor, which consists of a mixed ionic and

electronic conducting membrane (MIECM), can operate only at temperatures above

750°C. At lower temperatures the mechanism of exchanging oxygen ions and electrons

would not function. The air preheating is achieved by combustion of natural gas in

compressed air. The oxygen­depleted air leaving the OTM reactor is expanded in an

uncooled gas turbine.

The gas turbine is not added to the existing water/steam cycle but rather the process

layout changes significantly. Since the boiler requires around 165 kg/s of oxygen, the

inlet mass flow of the compressor in the order of 1500 kg/s is large compared to a

conventional heavy duty gas turbine. The large mass flow results in a large amount of

heat which needs to be utilised in the water/steam cycle. Therefore the feed water

preheating chain becomes unnecessary. Most of the feed water from the condenser to

the inlet of the supercritical once through boiler is heated by the exhaust gas of the

turbine and the steam/oxygen mixture leaving on the sweep side of the OTM reactor.

The remaining part of the feed water is heated by the exhaust gas stream from the

boiler.

4 Analysed Power Genera t ion Processes 63

Fig

ure

4-9

: Si

mpl

ifie

d fl

ow s

heet

of

a li

gnit

e fi

red

boil

er w

ith

an i

nteg

rate

d O

TM

rea

ctor

(L

FB

-OT

M)

and

CO

2 c

aptu

re.

64 4 .3 Ligni te Fi red Boi ler Cycles

The operation of the OTM reactor is very similar to that of the IGCC­OTM

process. The feed side of the OTM reactor operates on a gas turbine pressure level of

20 bar; the sweep side at 5 bar. The total pressure differences between both sides of

the membrane increases the oxygen partial pressure ratio, which is the driving force

for the mass transport through the membrane. The feed stream enters the OTM reactor

with a temperature of 900°C. LP steam is extracted from the steam turbine and is used

as a sweep gas on the sweep side to lower the oxygen partial pressure on this side of

the OTM reactor. Before the extracted steam is fed to the reactor it is preheated to

reduce the heat exchange from the hot feed stream to the cooler sweep stream. The

preheating of the steam is carried out by internal heat exchange with the steam/oxygen

mixture leaving the OTM reactor. The steam is heated from 250°C to 850°C before it

enters the OTM reactor. The amount of steam of the oxygen/steam mixture is about

60 mol­% after the OTM reactor. By cooling the mixture, the amount of steam is

reduced to 5 mol­% because the rest of the steam condenses. Before the oxygen is fed

to the boiler its temperature is 25°C.

The configuration of the once through boiler and the steam turbine remains

unchanged compared to the other boiler cycles. The separated (and cooled) oxygen is

then sent to the boiler. The combustion process is controlled with respect to the

maximum temperature by means of recirculated flue gas. The lignite is dried by use of

IP steam. As for the other configurations, the preheated feed water enters the once

through boiler at 300°C. The superheated live steam leaves the once­through boiler at

600°C and 280 bar. After expansion in the HP steam turbine, the steam is again

reheated to 620°C at an outlet pressure of 60 bar in the once­through boiler. Due to the

change of the feed water preheating chain, the number of extraction points in each

steam turbine is significantly reduced. Details on the steam extraction and the inlet and

outlet conditions are given for each steam turbine in table 4­15 to table 4­18. After

condensation at a pressure of 48 mbar and pressurising in the condenser pump, the

feed water is sent to the deaerator. Subsequently the feed water enters the preheating

chain again.

The separated mixture of water and CO2 is treated in the same way as for the

oxyfuel boiler process with cryogenic ASU. The CO2­rich stream is multistage­wise

compressed with intercooling and condensing of water between the stages. The

compression end pressure is 110 bar. After compression the purity of the CO2 is

96 mol­%.

The combination of the natural gas fired gas turbine and the lignite fired oxyfuel

boiler leads to the creation of a “mixed” cycle; a combination of a coal fired steam

4 Analysed Power Genera t ion Processes 65

power plant and natural a gas combined cycle. Due to this configuration it is

questionable which reference process (without CO2 capture) this process should be

compared with. Therefore the total heat input to the cycle can be divided into one part

for the gas turbine and another for the boiler. If the total heat were applied to the GT,

the process would refer to a pure combined cycle. If the heat load were simply brought

into the boiler, the process could be considered as a pure steam power plant. This way

enables the possibility to refer to a reference process as a combination of those

processes. For the configuration investigated in the present work, the GT heat load is

roughly 25% of the total heat input. The variation of the GT heat load and the

comparison to the reference process is described in chapter 7.2.

Another aspect caused by the integration of the gas turbine into the lignite fired

oxyfuel boiler is that the CO2 capture rate is reduced. The emissions produced in the

gas turbine are not captured. The reasons for this are, firstly, separation by means of

post­combustion CO2 capture of those emissions would require an additional capture

unit because the oxyfuel boiler process has no capture unit (the separation takes place

by condensation of the exhaust gas). Secondly, an additional CO2 capture would

adversely affect the net efficiency of the overall process. Again, in case of post­

combustion CO2 capture, steam would be required for the regeneration process of the

solvent. Furthermore, the low CO2 concentration would lead to a large amount of

solvent and subsequently require, specifically, a large amount of steam. For those

reasons it has been decided that the GT exhaust is not treated in any way but sent

directly to the stack. Thus, the overall CO2 capture rate decreases to around 65% ­

depending on the heat load of the GT.

Due to the size chosen for the lignite fired oxyfuel boiler process ( 1000 MW), the

required mass of air through the gas turbine is in the order of 1200 kg/s. Such a large

compressor inlet mass flow requires at least two heavy duty gas turbine based on

today’s largest available GT (such as the Alstom GT26 [183], Siemens SGT5000

[184] or Mitsubishi J­class [185]). The fact that at least two such heavy duty gas

turbines would be required gives an impression how large the combination of gas

turbines with integrated OTM reactor would be if such a cycle were realised.

Furthermore, it needs to be investigated if a conventional gas turbine could be used

without hardware modifications because no cooling flow (or only small quantities)

would be required due to the moderate temperature of 900°C at the combustor exit. In

addition to that, the gas stream leaving the combustor is reduced by the separation of

the oxygen in OTM reactor. Hence the mass flow in compressor and turbine might not

66 4 .3 Ligni te Fi red Boi ler Cycles

match, which could make hardware modifications necessary. This aspect is beyond the

scope of this work.

To facilitate comparison, details on the four steam turbines of the LFB­OTM cycle

are given in the following. From table 4­15 to table 4­18 the mass flow rates of

incoming, outgoing, and the extracted streams of each turbine are given. The high

pressure live steam is 730.1 kg/s, which is produced in the boiler. The outgoing steam

of the low pressure turbine that is fed to the condenser is 525.7 kg/s. In contrast to the

other boiler cycles only 28% of the live steam mass flow is extracted from the steam

turbines (for the other two cycles about 43% is extracted). The reason for the low

amount of extracted steam is the change of the cycle layout. The integrated gas turbine

and the OTM reactor ‘generate’ additional heat, which is utilised for feed water

preheating. Therefore much less steam is extracted from the steam turbines.

Consequently the overall power output of the steam turbines increases significantly by

more than 120 MW. The extraction form the IP steam turbines remains similar to the

other two configuration because the IP steam is used to drive the feed water steam

turbine, for drying the raw lignite and as sweep gas for the OTM reactor.

Table 4-15: LFB-OTM cycle: HP steam turbine data. The steam conditions at the inlet, the

outlet and the extractions regarding pressure, temperature and mass flows are

given. In addition, the gross power is shown.

Port m / kg/s T / °C p / bar

IN 730.1 599.1 249.7

OUT 730.1 378.4 65.0

Total power output MW 277.69

Table 4-16: LFB-OTM cycle: IP steam turbine data. The steam conditions at the inlet, the

outlet and the extractions regarding pressure, temperature and mass flows are

given. In addition, the gross power is shown.

Port m / kg/s T / °C p / bar

IN 738.8 622.5 62.0

OUT 525.7 231.2 4.1

First extraction 11.3 423.8 18.0

Second extraction 56.8 327.0 9.0

Third extraction 145.0 256.3 5.0

Total power output MW 557.71

4 Analysed Power Genera t ion Processes 67

Table 4-17: LFB-OTM cycle: first LP steam turbine data. The steam conditions at the inlet,

the outlet and the extractions regarding pressure, temperature and mass flows

are given. In addition, the gross power is shown.

Port m / kg/s T / °C p / bar

IN 170.3 231.2 4.1

OUT 170.3 32.2 0.05

Total power output MW 105.46

Table 4-18: LFB-OTM cycle: second LP steam turbine data. The steam conditions at the

inlet, the outlet and the extractions regarding pressure, temperature and mass

flows are given. In addition, the gross power is shown.

Port m / kg/s T / °C p / bar

IN 355.4 231.2 4.1

OUT 355.4 32.2 0.05

Total power output MW 223.08

In comparison to the LFB­OXY cycle the specific CO2 emissions are significantly

higher with 263 g CO2/kW(e) because, as previously mentioned, the exhaust gas

stream of the integrated gas turbine is not treated. Details on the exhaust gas stream

regarding CO2 emissions are given in the appendix in Table A4­3.

5 Model l ing of Essent ial Sub­Processes 69

5 Modelling of Essential Sub-Processes

In this chapter the most important sub­processes, of which the processes presented

in the previous chapter consist, are described in detail. The sub­processes are either

part of both processes, the IGCC processes and the lignite boiler processes, or are only

part of one of those processes.

5.1 A Generic Cooled Gas Turbine

5.1.1 General Information

The model of a cooled generic gas turbine has been developed within the

framework of the ENCAP project [13]. The model has also been published by

Jonssons et al. [187] in 2005. A schematic design of the cooled gas turbine is given in

figure 5­1. The gas turbine consist of an axial, adiabatic compressor and turbine and a

combustion chamber. In addition, pressure losses for the intake and exhaust have been

considered at the inlet of the compressor and the outlet of the expander. On the right

hand side of figure 5­1, the open gas turbine, also called Brayton cycle, for a typical

heavy duty gas turbine (pressure ratio of 16.8) is shown in a T,s­diagram.

(a) (b)

Figure 5-1: (a) A schematic layout of the generic cooled gas turbine; (b) cycle of the

generic cooled gas turbine in a T,s-diagram.

70 5 .1 A Gener ic Cooled Gas Turbine

The turbine inlet temperature (TIT) is commonly used to describe the temperature

level of existing gas turbines. In general, the performance of a gas turbine is

determined by the hot gas temperature. The hot gas temperature is that at the exit of

the combustion chamber. Gas turbine manufacturers do not communicate hot gas

temperatures of their existing turbines for commercial reasons. Furthermore the hot

gas temperature reflects the technology level of the gas turbine. Therefore a different

parameter is used to describe the temperature level of a gas turbine. Commonly, the

TIT is applied to compare different gas turbine among each other. Various methods

exist to determine the turbine inlet temperature. The ISO standard 2314 [173] is a

widely­used method to back calculate the TIT. For a cooled gas turbine – where hot

gas and cooling air are mixed at various positions in the turbine – the TIT is the

corresponding inlet temperature which would occur if the same turbine was uncooled

and achieved an identical power output. In other words, the TIT is a notional

temperature which would appear if the hot gas stream were mixed with the total

amount of cooling air before the expansion of the combustion gas took place. The

model of the generic cooled gas turbine is in accordance with the ISO standard 2314

[173].

The model of the generic cooled gas turbine is not only used in the context of

IGCC processes but also for different gas turbine processes such as oxyfuel gas turbine

processes within sub­project 6 of the ENCAP­project. The goal of modelling a generic

cooled gas turbine is two fold. Firstly, it is to consider the effect of cooling according

to today’s cooling technologies of heavy duty gas turbines, and secondly, since

different industrial partners were involved in the development of the gas turbine

modelling, the aim was to avoid using one manufacturer’s technology for the

modelling. Thermodynamically speaking the effect of cooling can be represented by

an additional pressure loss and a reduction of the overall polytropic efficiency of the

turbine. In cases involving different working fluids (such as those involving non­

standard fuels and oxyfuel combustions gases), while a detailed model of a cooled

turbine would need to be adjusted for different gas properties such as heat transfer

coefficients (HTC), this simpler model can be directly applied across all such cases.

Before the cooling model is described, some basic of thermodynamic performance

calculations of a gas turbine are presented. Afterwards the way the cooling is

considered in the performance calculation is described. Finally some parameter

variations have been conducted to show the impact of cooling on the operational

behaviour of the gas turbine over its whole operating range.

5 Model l ing of Essent ial Sub­Processes 71

5.1.2 Thermodynamic Basics of Gas Turbine Performance Calculations

Thermal efficiency

Gas turbine processes generally work according the Brayton cycle. An idealised

gas turbine process would consist of two isentropic – compression and expansion –

and two isobaric changes of state. In general, for any heat engine, the thermal

efficiency is defined as the ratio of output and input. Since the input to the heat engine

is a certain heat input and the output is the generated power, the thermal efficiency can

be expressed as

net

thin

W

Q. (5­1)

In equation (5­1) the absolute value of the power output is used because the general

convention of algebraic signs in thermodynamic calculation says that generated power

has a negative algebraic sign.

More specifically, the generated power of a gas turbine is the sum of power

generated by the turbine and the required power to drive the compressor:

net T CW W W . (5­2)

Merging equation (5­1) and (5­2), the thermal efficiency for a gas turbine yields:

T Cth

fuel

W W

m LHV. (5­3)

For calculating the net efficiency of the power generation cycle, the thermal

efficiency is reduced because of additional small losses occurring due to friction in

bearings of the rotor. These losses are represented by the mechanical efficiency mech.

Further losses occur in the generator where the mechanical power (rotational energy)

is converted to electricity (electrical energy). Applying both efficiencies the net

efficiency of power generation cycle can be written as:

net th mech gen . (5­4)

Specific technical work

The specific technical work equals the isentropic change in enthalpy from inlet to

outlet of a turbine or a compressor. Assuming that the working medium can be

72 5 .1 A Gener ic Cooled Gas Turbine

handled as ideal gas with a constant isobaric heat capacity, the difference in enthalpy

can be expressed by means of the temperature difference

,( )1

oo

s p out s inoh c T R T T

. (5­5)

For an isentropic change of state the temperature difference which occurs during

compression or expansion, temperature and pressure are connected via the isentropic

exponent as shown in the following equation

1

,

o

oout s out

in in

T p

T p

. (5­6)

The index “o” indicates that the working fluid is assumed as ideal gas. Equation

(5­5) and (5­6) can be used to calculate the specific technical work for an isentropic

expansion in a turbine

1

, , 11

o

ooout

tech rev T in oin

pw R T

p

. (5­7)

The specific work for the compressor is calculated in a similar manner. The

technical work of a gas turbine is defined as the shaft power divided by the compressor

mass flow. The technical work can be expressed as follows

, , , ,

,,

T in tech T C in tech Ctech GT

C in

m w m ww

m

. (5­8)

The technical work and reversible technical work are connected via isentropic or

polytropic efficiencies, respectively. This is described in detail in the next section.

Isentropic and polytropic efficiencies

For both compression and expansion in a compressor or turbine, respectively,

isentropic or polytropic efficiencies link irreversible and reversible technical work to

each other. By means of those efficiencies (and the reversible technical work) it is

possible to calculate the irreversible technical work either stage wise of a component

or the overall component itself. The polytropic and isentropic differ in their definition,

which is significant for multistage compression or expansion, because they compare dif­

5 Model l ing of Essent ial Sub­Processes 73

Figure 5-2: Comparison of a real change of state from point 1 to point 2 and an isentropic

change of sate from point 1 to point 2s. Schematically shown in a T,s-diagram.

ferent points which each other. The isentropic efficiency compares a irreversible

change of state (from 1 to 2 in figure 5­2) to an isentropic one from the same starting

point (from 1 to 2s in figure 5­2). Therefore the isentropic efficiency for an adiabatic

compression is defined as follows:

, , ,

,, ,

tech rev C out s ins C

tech irrev C out in

w h h

w h h

(5­9)

The analogical relation is valid for an adiabatic expansion in a turbine. Assuming

additionally that the working fluid behaves as ideal gases with a constant isobaric heat

capacity, enthalpy differences can also be expressed as temperature differences:

, ,

,, , , ,

tech irrev T out in out ins T

tech rev T out s in out s in

w h h T T

w h h T T

(5­10)

Assuming an adiabatic change of state for an ideal gas with a constant isobaric heat

capacity, the polytropic efficiency can also be expressed by means of isentropic and

polytropic exponents o and n, respectively [186]:

,1

1

o

p T o

n

n

(5­11)

This definition of the polytropic efficiency can be inserted into equation (5­10) for

calculating the isentropic efficiency, in which the specific technical work is expressed

in form of equation (5­7) – for the irreversible case that n o. By doing so the

following relationship between isentropic and polytropic efficiency can be written as

follows

74 5 .1 A Gener ic Cooled Gas Turbine

,1

,1

1

1

o

p To

o

o

out

ins T

out

in

p

p

p

p

. (5­12)

For calculating the isentropic exponent o the mean logarithmic temperature

ln

out in

in

out

T TT

T

T

(5­13)

and the mean logarithmic pressure

ln

out in

in

out

p pp

T

T

(5­14)

are required to be calculated between inlet and outlet condition.

A last step is required to consider the impact of cooling for modelling a gas turbine.

A simplified model to take the impact of cooling air into account is described in the

next section.

5.1.3 Modelling of film cooling

As mentioned earlier the reason for introducing cooling to the hot gas parts of a gas

turbine is that the material temperature limits the maximal allowable temperature of

the gas turbine process. Due to cooling it is possible that the highest temperature at the

turbine inlet is more than 300 K above the actual allowable material temperature. The

hot gas parts are protected by means of cooling against these high temperatures.

For gas turbines a widely­used cooling technology is film cooling where a gas

covers the surface area of the hot gas parts. In general, air is used as cooling medium

and is extracted from the compressor, bypasses the combustion chamber and is then

fed to the turbine or to hot parts of the combustion chamber. The cooling air enters the

hot gas path through a large number of small holes so that it forms a protective film for

the surfaces which are in direct contact with the hot gas stream. One of the latest

development is to use steam instead of air as cooling medium. The steam is supplied

from the heat recovery steam generator (HRSG). Some gas turbine manufacturer use

5 Model l ing of Essent ial Sub­Processes 75

this technology for their heavy duty gas turbines. Steam has the advantage that it has a

higher specific heat capacity. This leads either to a reduction in cooling gas flow or a

higher gas temperature while retaining the same material temperatures. A risk with

steam cooling is that the small cooling holes can be blocked which would interrupt the

cooling. From an operational point of view it might be difficult to provide steam for all

different types of operation, e. g. during start­up of the gas turbine.

As schematically shown in figure 5­1 the model assumes that the total cooling air

mass flow is extracted at the compressor exit, bypasses the combustion chamber

completely and is mixed with the hot gas stream leaving the combustion chamber. An

energy balance around the mixing process yields the mixed turbine inlet temperature

TIT, which can be expressed as follows:

p, p,

p,

Cb Ex Cb Ex Cb Ex cool cool cool

out out

m c T m c TTIT

m c

. (5­15)

Nowadays the mixed turbine inlet temperature ranges from 1200 to 1500°C.

Besides the limitation in maximum allowable metal temperature, another limiting

factor are nitrogen oxide emissions (NOx­emissions). Without any mitigation measures

it is assumed that a limit in acceptable NOx­emissions is reached at approximately

1600°C.

The model of the cooled gas turbine in the present work defines the ratio of cooling

air mass flow and hot gas mass flow shown in equation (5­15) by means of the

exchanged heat between cooling air and hot gas by means of introducing a maximum

allowable blade temperature. The blade temperature represents the metal temperature

which must not be exceeded. The ratio of cooling air mass flow and hot gas mass flow

is defined as:

p,CbExcool CbEx blade

CbEx p,cool blade cool

scm T T

bm c T T

. (5­16)

The parameter b and s in equation (5­16) can be adapted to represent a certain

technology level of the cooling system. Within the scope of the ENCAP­project these

empirical parameters have been set to b = 0.1884 and s = 1.0, respectively, which

represent a technology level of F­class heavy duty gas turbines such as Alstom’s

GT26B [183] or Siemens’ SGT5­4000F [184]. The isobaric specific heat capacities

used in the equation above are calculated by means of differences in enthalpy from the

appropriate inlet temperature and the assumed blade temperature. For the cooling air

stream the isobaric specific heat capacity is defined as

76 5 .1 A Gener ic Cooled Gas Turbine

blade cool,

blade coolp cool

h hc

T T

, (5­17)

and for the hot gas stream the isobaric heat capacity is accordingly calculated as

CbEx blade,

CbEx bladep CbEx

h hc

T T

. (5­18)

Applying equation (5­16) results in a model of the cooled gas turbine but

additionally to the impact on the mixed turbine inlet temperature, the cooling affects

the flow condition of the hot gas stream. The mixing of cooling air and the hot gas

stream will disturb the hot gas stream which again can be expressed as an ‘generated’

pressure loss for the hot gas stream. From a thermodynamic point of view a pressure

loss can be interpreted as a reduction in efficiency of the turbine or of the stage of a

turbine [187]. Therefore the idea is to calculate the additional pressure loss caused by

the mixing and convert this pressure loss into a reduction in polytropic efficiency for

the cooled turbine. This is described in the following.

In case of film cooling the cooling air enters the hot gas path via a large number of

small holes which are distributed across the hot gas parts – mainly blades and vanes in

the turbine – so that an even distribution of cooling air on the surfaces is achieved. The

flow pattern of the hot gas stream is disturbed because the cooling air needs to be

accelerated when it enters the hot gas path and, depending on the geometry of the

cooling holes, the cooling air changes its flow direction. These two effects are

responsible for the drop in pressure of the hot gas stream that can be expressed as

follows [188]

2cool cool

in CbEx CbEx

p m mMa K

p m m. (5­19)

The level of the pressure loss is proportional to the amount of the cooling mass

flow. The mixing loss factor accounts for the direction in which the cooling fluid is

injected into the hot gas stream. If the injection was perpendicular to the hot gas flow,

the mixing loss factor would be unity. The isentropic exponent , the Mach­number

Ma and the mixing loss factor can be summarised by only one model parameter K.

Assuming an isentropic exponent of around 1.3, a Mach­number in the range of 0.6­

0.8 with a conservative estimation of the mixing loss factor of 0.3 to 0.6, the simplified

model parameter K would range from 0.15 to 0.5 [187].

The additional pressure loss is as follows converted to a reduction in polytropic

efficiency p

5 Model l ing of Essent ial Sub­Processes 77

out

in

out

in

ln

ln

p p

p

p

pp

p p

. (5­20)

The change in polytropic efficiency is then subtracted from the polytropic

efficiency of the uncooled turbine to calculate the polytropic efficiency of the cooled

turbine. Finally, all parameters that have been used for calculating the polytropic

efficiency of the cooled turbine and then consequently the net efficiency of the gas

turbine are summarised in table 5­1.

Table 5-1: All used parameter of the generic model of the cooled gas turbine.

Description Symbol Value

Combustor pressure loss pCb 3%

Blade surface temperature Tblade 860°C

Polytropic efficiency of the uncooled turbine

hT, unc. 87.94%

Factor b 0.1884

Factor K 0.237

Exponent s 1

The point with a pressure ratio in the compressor of 17 and a combustor exit

temperature of 1425°C was used as a reference point. For the reference point a net

efficiency of the gas turbine of 38.5% has been agreed within the ENCAP­project. The

parameters b, K, s and the polytropic efficiency were adjusted accordingly to achieve

the net efficiency of the reference point (1425°C / PR = 17). For the range in hot gas

temperature and pressure ratio investigated, all parameters have been kept constant.

Because the model of the generic gas turbine was used for air­driven gas turbine as

well as for oxyfuel turbines, the pressure ratio has been varied from 10 to 40. For the

combustor exit temperature a range from 1200°C to 1600°C has been covered.

The net efficiency of the cooled gas turbine versus the specific shaft power11 is shown

in figure 5­3. The net efficiency is mainly determined by the pressure ratio of the

compressor. In total the net efficiency ranges from 33.0% to 42.5% for the investigated

pressure ratios. At smaller pressure ratios (< 20), the dependency of net efficiency on

the combustor exit temperature is small. For those pressure ratios the maximum net

efficiency is achieved between a combustor exit temperatures of 1150°C and 1400°C.

11 The specific shaft power is the net power output of the gas turbine divided by the compressor mass flow.

78 5 .1 A Gener ic Cooled Gas Turbine

Figure 5-3: Gas turbine net efficiency versus specific shaft power of the generic cooled gas

turbine for various pressure ratios (10 to 40) and combustor exit temperatures

(1000°C to 1600°C).

Figure 5-4: Gas turbine net efficiency versus combustor exit temperature for the generic

cooled gas turbine for various pressure ratios (10 to 40).

5 Model l ing of Essent ial Sub­Processes 79

For higher pressure ratios, the maximum in net efficiency is shifted towards higher

combustor exit temperatures. The exact maximum in net efficiency can be seen in

figure 5­4 where the net efficiency is plotted versus the combustor exit temperature.

For the largest pressure ratio of 40, the maximum net efficiency is reached at a

combustor exit temperature as high as 1500°C.

When analysing the results of the net efficiency it should be considered that the gas

turbine operates in combined cycle configuration. The net efficiency is a reasonable

measure to verify the modelling of the cooled generic gas turbine but in the context of

the analysed power generation processes, the main focus is on the combined cycle

efficiency. This is the reason for choosing the reference point with a pressure ratio of

17 and combustor exit temperature of 1425°C. The relatively small pressure ratio leads

to a gas turbine exit temperature that is high enough to provide a sufficient amount of

exhaust energy to the bottoming water/steam cycle. Therefore the maximum of the

combined cycle efficiency does not correspond to the maximum of the simple cycle

efficiency because the gas turbine exit temperature has to be taken into account when

evaluating the combined cycle efficiency.

As a result of the model of the generic cooled gas turbine, the cooling air mass flow

and the polytropic efficiency of the turbine are presented in figure 5­5 and figure 5­6,

respectively. The parameters shown in table 5­1 are used as boundary conditions for

the calculation of the results in figure 5­5 and figure 5­6. The total required cooling air

mass flow ranges from 5 to nearly 40% of the compressor mass flow for combustor

exit temperatures between 1000 and 1600°C; the pressure has been varied from 10 to

40. As expected, the required cooling mass flow increases continuously with

combustor exit temperatures. For small pressure ratios, the increase in cooling air mass

flow is nearly linear. The higher the pressure ratio, the larger the fraction of cooling air

because the additional pressure drop caused by the mixing of cooling air and the hot

gas stream varies linearly with the inlet pressure, see equation (5­19). Comparing the

different pressure ratios for a certain combustor exit temperature, the impact of the

pressure ratio on the cooling air mass flow continuously increases with higher

combustor exit temperature.

The polytropic efficiency of the turbine shows a contradictory picture compared to

the cooling air mass flow, compare figure 5­5 and figure 5­6. With increasing

combustor exit temperature, the polytropic efficiency decreases continuously. For

pressure ratios between 10 and 20, the polytropic efficiency shows a similar

distribution. For higher pressure ratios the polytropic efficiency decreases more

strongly with higher combustor exit temperatures. The distribution of cooling air mass

80 5 .1 A Gener ic Cooled Gas Turbine

Figure 5-5: Cooling air mass flow rate versus the combustor exit temperature; for various

pressure ratios (10 to 40). Cooling mass flow is expressed in percentage from

the compressor intake mass flow.

Figure 5-6: Polytropic efficiency of the generic cooled turbine versus the combustor exit

temperature; for various pressure ratios (10 to 40).

5 Model l ing of Essent ial Sub­Processes 81

flow and polytropic efficiency fit to each other because the higher the mass flow of

cooling becomes, the larger the difference between the polytropic efficiency of the

uncooled and the cooled turbine must be. It can be seen from equation (5­20) that a

larger pressure drop leads to bigger drop in the polytropic efficiency.

The model of the generic cooled gas turbine is used for the investigated IGCC

processes described in chapter 4.2. All investigated IGCC processes assume that the

gas turbine operates at the reference point with a pressure ratio of 17 and a combustor

exit temperature of 1425°C.

5.2 Coal Gasification

In general gasification processes at an elevated pressure can be classified by the

type of reactor: (i) moving bed, (ii) fluidised bed, and (iii) entrained flow reactor.

Another way of categorisation is [189]:

allothermal or autothermal gasification

oxygen­ or air­blown gasification

cooled or adiabatic gasifier

pressure and temperature level of the gasification

Figure 5­7 shows the different types of gasifiers. The major differences are the

operating conditions, residence time of the fuel in the gasification process and the gasifi­

(a) (b) (c)

Figure 5-7: Different reactor types used in gasification processes:(a) moving bed gasifier,

(b) fluidized bed gasifier, (c) entrained flow gasifier [191].

82 5 .3 Sulphur Removal

cation temperature. The residence time of the coal in the gasification process varies

from a few seconds in entrained flow gasifiers up to one hour in a moving bed reactor.

A high temperature in the gasification process is required to get a high carbon

conversion rate. A low carbon conversion rate represents a loss of thermal efficiency

because the non­converted carbon leaving the gasification can no longer be utilised

[190]. In addition, a high gasification temperature favours a high carbon conversion

rate. On this account, oxygen­blown gasifiers are favourable for IGCC processes. In

the present work an entrained flow gasifier is considered for all IGCC cycles.

With regards to the chemical reactions occurring, gasification of coal is a very

complex process. The main chemical reactions determining the composition of the

produced syngas contain the heterogeneous solid­gas reactions [192]:

Partial combustion C + ½ O2 CO (exothermic)

Combustion C + O2 CO2 (exothermic)

Gasification with hydrogen C + 2H2 CH4 (exothermic)

Boudouard reaction C + CO2 2 CO (endothermic)

Gasification with steam C + H2O CO + H2 (endothermic)

and the homogeneous gas­gas reactions [192]:

Water­gas shift reaction CO + H2O CO2 + H2 (exothermic)

Methanation CO + 3H2 CH4 + H2O (exothermic)

When modelling a gasification process using these main reactions chemical

equilibrium is assumed without considering chemical kinetics. Although equilibrium is

theoretically only reached after infinite time, the time of reaction in an entrained flow

gasifier is so short that equilibrium can be assumed [189]. In this work the gasification

process is numerically modelled using Aspen Plus. In Aspen Plus the gasification is

split into two modules RYIELD and RGIBBS taken from the Aspen Plus library. In

the first reactor (RYIELD) the coal mass flow is calculative divided into its different

elements: C, H2, O2, S, H2O, Cl2 and ash. In the second reactor (RGIBBS) the

composition of the raw synthesis gas is calculated under the presence of steam and

oxygen [189].

5.3 Sulphur Removal

In the IGCC cycles desulphurisation takes place after de­dusting of the synthesis

gas. The whole desulphurisation process comprises different steps (in chronological

order) [189]:

5 Model l ing of Essent ial Sub­Processes 83

COS­hydrolysis: conversion of COS to H2S

Absorption: separation of H2S from synthesis gas (absorption in the

solvent)

Regeneration: desorption from the solvent

Claus process: catalytic conversion of H2S to elemental sulphur

Tail gas process: catalytic conversion of exhaust gas of the Claus process

The separation is achieved by means of chemical absorption. Here Methyl

diethanolamine (MDEA) is assumed as solvent. Due the chemical stability of the

solvent, the synthesis gas needs to be cooled before desulphurisation to around 40°C.

The process has not been modelled in great detail but rather the specific expenditure of

energy for the sulphur removal has been adapted from [189] and a separation rate of

90% of H2S is assumed.

5.4 CO-Shift Reaction and CO2 Separation Process

If CO2 capture is applied to an IGCC cycle consequently also a CO­shift reaction is

required, see section 4.2.2. In the CO­shift reaction CO is converted to CO2 by adding

steam (and thus heat) to the synthesis gas. The CO­shift maybe prior or after to the

desulphurisation. If the CO­shift reaction is prior to the desulphurisation it is then

referred to as sour­shift configuration. In case it takes place after desulphurisation it is

called sweet­shift configuration. According [193] the CO­shift reaction is carried out

in two stages. The “Selexol” process uses the physical solvent Dimethyl ether of

polyethylene glycol (DMPEG) [194]. In case of CO2 capture the Selexol unit removes

CO2 as well as sulphur components. The Selexol is selectively regenerated to produce

separate CO2 and sulphur components streams [193, 195]. In a simplified model of the

CO­shift reaction an conversion rate of 93% is assumed and a capture of 98% for the

CO2 leaving the shift reaction is assumed.

5.5 CO2 Compression

After separation of the CO2 it needs to be compressed for transport and storage. In

accordance with [172] for all configurations investigated, the pressure after

compression has been defined as 110 bar. This assumption was been used for all

calculations throughout the whole ENCAP project. Within the sub­project 6 it was

further decided to use a given specific expenditure of energy for the compression of

the separated CO2 from 1 atm to 110 bar. In the “ENCAP framework SP 6” [171], a

84 5 .5 CO2 Compression

value of 0.365 MJ/kg CO2 for the specific expenditure of energy was defined. The

advantages of using such a simplified model for the compression is, that firstly for all

calculations throughout different working groups the same assumption is used.

Secondly, the stream which is compressed is not pure CO2 but it is rather a mixture of

CO2 and H2O and small amounts of non­condensable gases such as N2, O2 and Ar. Due

to the behaviour of such a mixture an accurate calculation of the compression process

is challenging for existing equations of state because this mixture does not behave as

an ideal gas. Therefore such a simple assumption is a good compromise between

accuracy and effort for the simulation. Finally it helps the quantitative comparison of

simulation results from different working groups (using different simulation tools).

6 Model l ing of In tegra ted Membrane Reactors 85

6 Modelling of Integrated Membrane Reactors

6.1 Introduction to the Modelling

The theory of both investigated types of membrane reactors, OTM and hydrogen­

selective membranes, have been described in chapter 3. In this chapter the focus is on

process related parameters which are important if such a membrane reactor is part of

the power generation cycle. Therefore, the following aspects have been considered:

The permeate flux for given conditions on both sides of the reactor

The pressure along the reactor for both streams

The overall heat transfer coefficient of the reactor

The size of the membrane surface area

Both types of reactors, the OTM and the hydrogen­selective membrane reactor, are

modelled in the same way. The reactors are assumed to be counter­flow apparatuses.

The principle layout is schematically shown in figure 6­1. The stream from which the

respective component is separated, is named the feed stream. After separation from

oxygen or hydrogen, respectively, the stream leaves the membrane reactor as the

retentate stream. On the other side of the membrane reactor, the sweep stream enters

the membrane reactor. The mixture of the sweep stream and the mass flow transferred

through the membrane (oxygen or hydrogen, respectively) leaves the membrane

reactor as permeate stream.

Figure 6-1: Schematic layout of the model of the membrane reactor.

The mass flows, pressures and temperatures of all streams impact the operating

conditions of the membrane reactor. Therefore the reactor should not only be

considered for separation of the permeate stream from the feed stream but also as a

heat exchanger and a pressure vessel. The temperature and the mass flow rate of the

sweep stream determine the amount of heat which is transferred inside the membrane

reactor. Furthermore the difference in total pressure across the membrane causes a

86 6 .1 In t roduct ion to the Modell ing

certain mechanical load on the membrane material. The combination of operating

temperature and pressure difference result in a limit regarding the mechanical stresses

for the membrane material. This fact results in contradictory requirements for the

membrane reactor: on the one hand, high temperature and a large pressure difference

across the membrane promote a high permeation; but on the other hand, these

conditions generate high thermal and high mechanical stresses for the membrane

material. The challenge is to find operating conditions for which these contradictory

requirements can be met satisfactorily.

Figure 6-2: Discrete elements into which the membrane reactor is divided. For each

element heat and mass transfer is calculated in an iterative way.

The membrane is, in general, modelled as a counter­flow apparatus where not only

mass transfer takes place, but also heat is transferred from the feed stream to the sweep

stream due to different inlet temperatures of both inlet streams. Both, heat and mass

transfer are determined by the temperature, amongst other parameters, of both streams

at a certain location on the membrane reactor. Therefore the membrane reactor is

divided into discrete elements and for each element the calculation of heat and mass

transfer is conducted, see figure 6­2. For numerical reasons the OTM reactor is

modelled with 75 elements, whereas the hydrogen­selective membrane reactor is

modelled with 128 elements.

The mass transport through the membrane is determined by temperature, pressure

and the partial pressure of each component that is transferred through the membrane,

of both streams in the reactor. At the same time the development of the temperature

along the length of the membrane reactor is affected by the mass transfer through the

membrane. In a first step the temperature profile along the reactor length assuming a

counterflow heat exchanger without mass transport is calculated. The calculated

temperature distribution for both streams is used for a first computation of the

permeation flux for each element along the reactor length. By utilising the new mass

flows through each element, the temperature distribution for both streams is updated.

Afterwards by means of the updated temperature distribution, the permeation flux for

each element is re­calculated. These two steps are repeated iteratively until

6 Model l ing of In tegra ted Membrane Reactors 87

convergence is reached. For each element and both streams the total pressure, the

partial pressure of the permeate, the temperature and the isobaric heat capacities are

computed. The required thermophysical properties are calculated by means of the

Redlich­Kwong­Soave equation of state [196].

For the heat transfer the overall heat transfer coefficient depends on the flow

condition of each stream in the membrane reactor. Assuming a given geometry of the

membrane reactor, the velocity for each stream can be calculated. With those

assumptions the pressure drop can be estimated which is explained in more detail in

the following.

6.2 Pressure drop

The pressure drop for each stream in the membrane reactor depends on the

geometry of the reactor and the flow velocity of the appropriate stream, as described in

chapter 3.1. Due to the fact that this work does not focus on the design of a specific

reactor, a generic layout of the membrane reactor has been assumed which is

schematically shown in figure 6­3. The feed stream flows inside the tubes while the

sweep stream flows on the shell side of the tubes. Because of the large size of those

reactors, effects at the inlet or outlet of the reactor are neglected in this pressure drop

model.

The cross­section for the sweep stream, indicated with A in figure 6­3, can be

calculated using the following relation for a given outer diameter da of the tubes:

da

di

da

A

Figure 6-3: Assumed geometry and arrangement of the tubes of the membrane reactor.

88 6 .2 Pressure drop

2

2

4

aa

dA d . (6­1)

To determine the Reynolds number the hydraulic diameter of the cross­section for

the sweep stream needs to be calculated. The hydraulic diameter dh can be expressed

as:

4 4

h

a

A Ad

U d (6­2)

The pressure loss will be determined by the actual conditions of the streams in the

membrane reactor. The goal of this calculation is to estimate the pressure loss for

design point conditions and to evaluate the level of pressure loss in such a membrane

reactor. The pressure loss at the design point is the most important operation point

because at off­design conditions the pressure drop will decrease due to smaller mass

flows and therefore its impact becomes less important on the overall performance of

the power generation process.

As indicated by figure 6­3 the membrane reactor comprises a certain number of

tubes having a certain inner and outer diameter and a defined length. Given these four

parameters the pressure drop of each stream can be calculated for given flow

conditions. This is described in the following using some assumptions taken from

literature. The data used for this calculation are taken from the AZEP­project [97­101].

The utilisation of the membrane reactor in the AZEP­project is similar to the

configuration described in this work. Therefore it is reasonable to make use of the

assumptions considered in the AZEP­project.

The AZEP project assumed a specific configuration of a SGT­800 from Siemens

(formerly Alstom GTX100). For this gas turbine the conditions after the compressor

have been taken from the gas turbine library from GateCycleTM [196]. The feed stream

of the membrane reactor is assumed to be identical to the compressor exit conditions.

For the SGT­800 in the design point the conditions after the compressor are assumed

to be: 110 kg/s, 20.27 bar and 450°C. With those assumptions the volumetric flow rate

can be calculated by means of the specific volume of the fluid. The specific volume

has been calculated using Lemmon et al. [198] in RefpropTM [199]. The mass flow of

110 kg/s results in a volumetric flow rate of 7.06 m3/s. Estimation of the size of the

AZEP membrane reactor (according [200]) yields in an overall volume of 50.264 m3.

Furthermore, assuming a surface to volume­ratio of 750 m2/m3, the overall membrane

surface area can be calculated to 37,698 m2. With these overall flow conditions the

6 Model l ing of In tegra ted Membrane Reactors 89

required cross­sectional area A can be written as the ratio of the volumetric flow rate V and the mean flow velocity w:

V

Aw

. (6­3)

At the same time, assuming the geometry illustrated in figure 6­3, that area is also

defined by the total number of tubes and their inner diameter

2

4id

A n

. (6­4)

Merging equations (6­3) and (6­4), the total number of tubes can be expressed as

2

4

i

Vn

w d

. (6­5)

The membrane surface area is determined by a combination of the total number of

tubes and the geometry, meaning the ratio between inner diameter and length.

Depending on these two parameters the membrane surface area can be written as:

Mem iA n d l (6­6)

or as:

Mem

i

Al

n d . (6­7)

Equations (6­5) to (6­7) show how the parameters diameter, length and number of

tubes are related to each other and result in a certain membrane surface area.

Furthermore, the assumptions regarding the reactor geometry are used in the following

to determine the pressure drop along the membrane reactor. The calculation of the

pressure drop is conducted according the VDI Heat Atlas [201]. In general, the

pressure drop inside a circular tube is defined as

2

2i

i

l wp

d

, (6­8)

where is the friction factor, l the length of the tube, the density of the fluid and wi

its mean flow velocity; di stands for the inner diameter of the tube. In addition the

friction factor depends on the Reynolds number which is defined as follows

90 6 .2 Pressure drop

Re i iw d

. (6­9)

The density and the dynamic viscosity have also been calculated by using

Lemmon et al. [198] in RefpropTM [199], for the inlet condition of the membrane

reactor of 20.27 bar and 450°C. Then, the friction factor is taken from the VDI Heat

Atlas for the appropriate range of 3,000 < Re < 100,000. In this range the friction factor

is described by

1

40.3165 Re

, (6­10)

Having the friction factor calculated, the pressure drop can be determined

according to equation (6­8). The relative pressure drop versus the inner tube diameter

is shown in figure 6­4. The relative pressure drop refers to the inlet pressure of the

sweep stream of 20.27 bar. In the range of 0.5 mm and 2.0 mm, the pressure drop

decreases from around 3.3% to 2.3%. The distribution of the pressure leads to the

conclusion that the increase of the tube diameter, which reduces the pressure drop, has

a stronger impact on the pressure than the tube length, which increases the pressure

drop.

Figure 6-4: The relative pressure drop and the tube length vs the inner diameter of the

tubes.

6 Model l ing of In tegra ted Membrane Reactors 91

The required tube length for each inner diameter is given on the ordinate on the

right hand side in figure 6­4. For the same range of inner diameter the tubes length

varies from around to 6 to 15 meters. The calculation of the pressure drop

demonstrates that a relative pressure drop of approximately 2.5% is reasonable and is

in the range of the pressure drop of 3% according [99] as desired design criteria.

6.3 Heat Transfer

The goal of the membrane reactor is to separate the permeate from the feed stream,

the reactor itself operates to some extent also as a heat exchanger due to the different

temperature levels on both sides of the membrane reactor; therefore the heat transfer

from the feed stream to the sweep stream cannot be avoided and needs to be calculated

because the temperatures of both streams have an impact on the mass transfer. As

shown in chapter 3.2, the ability of the membrane material to transport ions and

electrons depends strongly on its temperature.

The feed stream enters the membrane reactor at a certain temperature; 900°C in

case of the OTM reactor and 600°C for the hydrogen­selective membrane reactor. In

any case the temperature level is higher than that of the sweep stream. Considering the

membrane reactor as a heat exchanger, the capability of transferring heat is determined

by the product of surface area A and the overall heat transfer coefficient k [202]:

1

1 1

Feed Feed Mem Mem Sweep Sweep

k A

A A A

. (6­11)

For the feed stream which flows inside of the tubes, the heat transfer coefficient

Feed is determined by the Nusselt number Nu, the thermal conductivity Feed and the

inner diameter di:

Feed

iFeed

Nu

d

. (6­12)

The same assumptions as used in the previous section for determining the pressure

drop have been used to calculate the Nusselt number in equation (6­12). For the inlet

conditions of the feed stream the Reynolds number according equation (6­9) yields a

value of higher than 3000, which represents a turbulent flow. For a turbulent flow in a

circular tube the VDI Heat Atlas [201] gives the following relationship to calculate the

Nusselt number

92 6 .3 Heat Transfer

2

3

(Re 1000)Pr8 1

1 12.7 Pr 18

idNu

l

, (6­13)

where the friction factor is defined as

2(1.82lg(Re) 1.64) . (6­14)

According equation (6­12) the convectice heat transfer coefficient yields a value of

320 W/(m2 K), assuming the thermal conductivity of the membrane material with

1 W/(m K). Baehr [202] gives a value of 1.03 W/(m K) for a typical ceramic material.

The conditions of the sweep stream, which flows on the shell side of the tubes, are

used to calculate the heat transfer coefficient. The grey shaded area in figure 6­3

shows the cross section where the sweep flows through. For this geometry a hydraulic

diameter is calculated which is required for determine the Reynolds number. In case of

the OTM reactor (for the IGCC­OTM process and for the oxyfuel boiler process, see

figure 4­5 and figure 4­9) steam is used as sweep gas. The conditions of the steam

used as sweep stream are a temperature of 450°C at a pressure of 1.5 bar. For the

hydrogen­selective membrane reactor nitrogen is employed as sweep gas, see figure 4­

6. In this section the heat transfer coefficient is calculated exemplarily for the OTM

reactor. The calculation for the hydrogen­selective membrane reactor is conducted

accordingly.

Due to the fact that the flow of the sweep stream does not flow through a tube but

on the shell side of the tubes, a hydraulic diameter is used to calculate the convective

heat transfer coefficient for the sweep stream.

Sweep

hydSweep

Nu

d

. (6­15)

The calculation of the heat transfer coefficient is done in the same manner as for

the feed side, compare equation (6­12) and (6­15). The hydraulic diameter according

to [201] is defined as

4

hydA

dU

. (6­16)

The same holds true for calculation of the Reynolds number, where the hydraulic

diameter also has to be employed,

6 Model l ing of In tegra ted Membrane Reactors 93

Rehydw d

. (6­17)

For calculation of the Reynolds number a mean flow velocity of 5 m/s has been

assumed. The thermophysical properties of the steam (inlet conditions: 450°C and

1.5 bar) have been calculated with Wagner and Pruß [203] in RefpropTM [199]. Due to

low density of the steam, the Reynolds number is much lower than that of the feed

stream. The flow of the sweep stream is therefore considered to be laminar. For such a

flow the Nusselt number is given by the following equation [201]

1 23 3

3.657 0.0499tanh

tanh 2.264 1.7

Nu XX

X X

. (6­18)

In equation (6­18) X+ represents the flow length, which can be calculated as

follows:

l

XdPe

, (6­19)

in which the Péclet number Pe is the product of Reynolds number and Prandtl number

Re PrPe . (6­20)

For the assumptions mentioned above the Nusselt number yields a value of 3.567,

which again results in a convective heat transfer coefficient of 250 W/(m2 K) for the

sweep stream. Using the same surface area of the feed stream as for calculation of the

pressure drop – in section 6.2 – of 37,698 m2 and assuming a inner and outer diameter

of 1 mm and 1.5 mm, respectively, the overall heat transfer coefficient according

equation (6­11) yields a value of 152 W/(m2 K). Since the heat transfer from the feed

stream to sweep stream impacts the temperatures of both streams, it also has an

indirectly influence on the mass transfer because that depends as well on the operating

temperature of the membrane reactor.

6.4 Mass Transfer

In section 3.2 of the previous chapter, the theory of the mass transport mechanism

was presented for different membrane types. Here, the method for calculating the

permeate flux through each type of membrane will be presented. In particular the

behaviour of oxygen transport and hydrogen­selective membranes, which are used in

94 6 .4 Mass Transfer

this work, will be considered. For both types of membrane, the permeation flux is

calculated specifically per unit area, so that for a given overall membrane surface area

the total mole flow can be determined.

Oxygen Transport Membrane Reactor

As previously mentioned the mass transfer through a oxygen transport membrane is

determined by several parameters. First, at least for a certain thickness of the

membrane, the mass transport is mainly determined by the diffusion of electrons and

ions through the material [151]. Secondly, the operating conditions regarding pressure

and temperature influence the oxygen permeation through the membrane. For the

oxygen transport membranes considered in this work the mass transfer can be

expressed by the Wagner equation (see equation (3­1)). Experimental data from

Shaula et al. [122] are used to evaluate the calculation of the oxygen permeation

obtained from the Wagner equation. Shaula et al. [122] investigates

LaGa0.65Ni0.20Mg0.15O3­ (based on the LaGaO3 perovskite structure), which represents

a common membrane material. It is not the goal to benchmark different membrane

materials, but rather to investigate the impact of different operating conditions for the

membrane reactor on the required membrane surface area. Therefore it seems

reasonable to use a ‘typical’ membrane material and to focus on the relative changes

for the chosen membrane material. Of course, if different membrane materials were

investigated, the absolute oxygen flux would differ, but the relative changes would

show a similar picture. This aspect is discussed in more detail in the section 6.5.

The calculation of the oxygen flux according the original form of the Wagner

equation (see also section 3.2.1) is not capable of matching the experimental data from

Shaula et al. [122]. To overcome this mismatch two correction factors are introduced

to the Wagner equation. An additional term, c, for the membrane thickness and an

exponential factor, b, are introduced to the Wagner equation. Applying these two

additional factors the adapted Wagner equation becomes

2

2

,

,

log

A

m

E b

O FeedR Tm

m O Sweep

pcj e

X c p. (6­21)

The values of the parameters applied in equation (6­21) are summarised in table

6­1. Both the oxygen flux obtained from the adapted Wagner equation as well as the

experimental data from Shaula et al. [122] are presented in figure 6­5. The logarithm

of the oxygen flux is plotted against the logarithm of the ratio of oxygen partial

6 Model l ing of In tegra ted Membrane Reactors 95

pressure of feed and permeate side of the membrane. The values calculated by the

adapted Wagner equation are represented by the blue solid lines, whereas the

experimental data from Shaula et al. [122] are given by the red points. The curves are

shown for three different operating temperatures: 850, 900 and 950°C. The deviations

between calculation and experimental data are acceptably small for all three

temperatures. The driving force for the mass transport through the membrane is the

ratio of oxygen partial pressure. Additionally, the permeation is favoured by high tem­

Table 6-1: Parameters used in the adapted Wagner equation to calculate the oxygen flux

through the OTM membrane.

Parameter Unit Value

Pre­factor cm mol s­1 m­1 0.1502

Activation energy EA kJ mol­1 114.32

Membrane thickness Xm mm 1.0

Factor b 1 0.7873

Factor c mm 0.2987

Figure 6-5: Logarithm of the oxygen permeation flux versus the logarithm of the ratio in

oxygen partial pressure. The red points represent experimental data from

Shaula et al. [122]; the solid lines are calculated by the adapted Wagner

equation (equation (6-21)).

96 6 .4 Mass Transfer

peratures. Both effects can be seen in figure 6­5, the oxygen flux increases with

operating temperature as well as with a larger ratio in oxygen partial pressure. Of

course, the aim is to achieve an oxygen permeation flux as high as possible to reduce

the overall membrane surface area. In this work the operating conditions of the

membrane reactor and the total required oxygen mass flow are used as input for the

calculations of the membrane reactor. The operating conditions (mass flows,

temperature and pressure on both sides of the membrane reactor) determine the oxygen

permeation flux and the total oxygen mass flow determines eventually the required

overall surface area of the membrane reactor.

The operating temperature is mainly determined by the temperature of the feed

stream entering the membrane reactor. Similarly to gas turbines, the maximum

allowable temperature is determined by limitation of the material properties of the

membrane. This upper temperature limit and the minimum required temperature to

operate a MIEC membrane (> 750°C) determines the range of operating temperature

of such a membrane reactor. Due to these circumstances the feed temperature is

assumed to be constant at 900°C for all configurations investigated. A feed

temperature of 900°C seems reasonable, or at least, a reasonable compromise between

achieving high permeation fluxes through the membrane and a technically feasible

temperature of such membrane reactors.

The combination of mass flows and the total pressure on both sides of the

membrane reactor determines the ratio in oxygen partial pressure in each location

along the membrane reactor. The higher the difference in total pressure on both sides

of the membrane, the larger the ratio in oxygen partial pressure. Moreover, the larger

the mass flow of the sweep stream, the smaller the oxygen partial pressure on the

permeate side of the membrane reactor (mixture of sweep and permeate streams).

These considerations reveal that several combinations of total pressure, mass flows

and temperature levels of both streams (feed and sweep streams) could be considered

to investigate the impact on the overall membrane surface area. The conditions of the

feed streams have been kept constant, whereas both the pressure and the mass flow of

the sweep stream have been varied in the parametric studies. The results of the

parametric studies are presented in section 6.5.1.

Hydrogen-selective Membrane Reactor

As described in chapter 3.2.2, the mass transport mechanism can be expressed by

the Sieverts equation, see equation (3­7). For thick membranes (thickness > 100 m)

6 Model l ing of In tegra ted Membrane Reactors 97

the main resistance for the mass transport is assumed to be the transport of hydrogen

atoms through the palladium. Under these conditions, the surface reaction is

considered to be very fast and the dissolved hydrogen atoms are in equilibrium with

the hydrogen gas on either sides of the membrane [159].

As previously described, in general, hydrogen­selective membranes are made of

palladium and in contrast to oxygen transport membranes, there is less variety in

membrane material. Bulk palladium and palladium­based membranes are considered

to be the most promising material to achieve both a high selectivity and a high

hydrogen permeation. The Sieverts equation to express the hydrogen permeation

through the membrane is evaluated by means of experimental data from Morreale et al.

Morreale et al. [169] fabriacted several membrane disks with a thickness of 1 mm and

measured the hydrogen permeation through the membrane at temperatures between

350 and 900°C for a large range in hydrogen partial pressure differences.

Figure 6-6: Hydrogen permeation flux versus difference in hydrogen partial pressure. The

points represent experimental data from Morreale et al. [169]; the solid lines

are calculated by the Sieverts equation (equation (3-7)). The different symbols

stand for different membrane samples which Morreale et al. investigated.

98 6 .4 Mass Transfer

Both the experimental data and the calculation of the hydrogen permeation flux are

illustrated in figure 6­6. The hydrogen permeation flux is proportional to the gradient

of the hydrogen partial pressure across the membrane. With increasing temperatures

the rate of change of the hydrogen permeation flux increases. The different symbols

represent different membrane samples, which Morreale et al. [169] fabricated and

investigated. The Sieverts equation, which is used to calculate the solid lines in figure

6­6, has been already previously presented in chapter 3.2.2, see equation (3­6). The

pressure exponent of n = 0.5 in equation (3­6) reflects the dissociation of the gaseous

hydrogen molecules into two hydrogen atoms that diffuse into the metal. As

mentioned in chapter 3.2.2 a pressure exponent of n = 0.5 means that the mass

transport through the membrane is determined by bulk diffusion, for which it is

reasonbale to assume due to a membrane thickness of 1 mm. All parameters used in

the Sieverts equation are listed in table 6­2.

The temperature of the feed stream entering the membrane reactor determines the

operating temperature of the reactor. The hydrogen­selective membrane operates at

temperatures above 300°C. Of course, the higher the operating temperature, the higher

the hydrogen permeation flux. The upper temperature is restricted by the limit of the

maximum allowable material temperature of the membrane material. The lower and

the upper limit of the operating temperature define the theoretical possible temperature

range from around 300 to 900°C. In case of the hydrogen­selective membrane reactor,

the possible achievable temperature of the syngas stream (feed stream) constrains the

operating temperature of the membrane reactor. The membrane reactor is located in

the syngas stream (between gasifier and combustor of the gas turbine), see figure 4­6.

Therefore it is assumed that the feed stream is heated to 600°C before it enters the

membrane reactor. This is a compromise of being high enough to achieve a reasonable

high hydrogen permeation flux, but still being feasible in terms of internal heat

exchange of the syngas stream from gasifier to the combustor of the gas turbine.

Table 6-2: Parameters used in the Sieverts equation – equation (3-6) – used for the

calculation of the hydrogen flux through the membrane.

Parameter Value

Pre­factor k0 1.92 10­7

Activation energy EA kJ mol­1 13.81

Membrane thickness Xm mm 1.0

Pressure exponent 0.5

6 Model l ing of In tegra ted Membrane Reactors 99

With respect to the mass flows and composition resulting in a certain hydrogen

partial pressure along the membrane reactor, the condition of the feed stream (syngas

stream) is defined by the gasification process. Therefore the conditions (mass flow,

temperature and pressure) of the feed stream are assumed to be constant. In contrast,

the conditions of the sweep stream is varied. The sweep stream is pure nitrogen

provided as a byproduct from the ASU. The conditions of the sweep stream are varied

with regards to its mass flow and total pressure. A low hydrogen partial pressure can

be achieved by, either or both, a low total pressure of the sweep and a large mass flow

of nitrogen which reduces the mole fraction of hydrogen. This variations lead to

different hydrogen partial pressure on the sweep side of the membrane reactor and thus

to a change of the difference in hydrogen partial pressure across the membrane.

Therefore several combinations of total pressure, mass flows of the sweep stream are

investigated to establish the impact on the overall membrane surface area. The results

of the parametric studies are presented in section 6.5.2.

The approach for investigating the performance of the hydrogen­selective

membrane is the same as that one for the OTM reactor. The focus in this work is on

the operating conditions of the membrane reactor rather benchmarking the hydrogen

permeation for different membrane materials or various configurations of the

membrane geometry or the design of the membrane reactor. Therefore the operating

conditions have been varied to investigate how the hydrogen permeation flux differs

with changed operating conditions. This approach leads to the fact that the changes in

hydrogen permeation flux are analysed rather than the absolute level. The goal is to

analyse how the membrane surface area changes relatively for different operating

conditions.

It should be emphasised that a major difference between the OTM and the

hydrogen­selective membrane reactor is the fraction of permeate and feed stream. In

case of the hydrogen­selective membrane reactor, the hydrogen is separated from the

syngas stream. Therefore the goal is to separate the hydrogen completely from the feed

stream. The amount of hydrogen which is not separated from the feed stream can be

interpreted as a loss in heat for the overall power generation process. The requirement

to separate all hydrogen is more challenging for simulation of the separation process.

Therefore less variations are conducted for the hydrogen­selective membrane reactor

compared to the OTM reactor. In comparison to that the OTM reactor separates a

smaller fraction of oxygen from the feed stream. Therefore most of the parametric

studies are performed for the OTM reactor. The parametric studies are described in the

following section.

100 6 .5 Parametric Studies on the Membrane Reactors

6.5 Parametric Studies on the Membrane Reactors

6.5.1 Results for the Oxygen Transport Membrane (OTM)

In both analysed cycles, the IGCC­OTM and the LFB­OTM, the configuration of

the OTM reactor is identical. The OTM reactor is integrated into a gas turbine cycle.

Compressed air leaving the compressor at around 450°C is further preheated to 900°C

before the oxygen­rich gas stream (a mixture of preheated air and a small fraction of

combustion gases) is fed to the OTM reactor.

The configuration of the IGCC­OTM process is described in section 4.2.3, see

figure 4­5. The lignite fired oxyfuel boiler with OTM is depicted in 4.3.3, see figure 4­

9. A small difference is that in case of the IGCC­OTM cycle, the fuel is syngas

produced by the gasification (mainly hydrogen and nitrogen), whereas for the lignite

fired oxyfuel boiler, the fuel is assumed to be natural gas. The impact on the

composition of the feed stream, in particular, on the oxygen concentration is

negligible. Furthermore, the retentate stream leaving the OTM reactor is treated in

different ways for the two power cycles. In case of the IGCC­OTM cycle, the retentate

stream is further heated before the exhaust gas expands in the turbine. The

combination of pressure ratio and turbine inlet temperature is optimised for combined

cycle configuration, see section 4.2.3.

According [171] the live steam temperature of the water/steam cycle is defined as

565°C. Therefore the exit temperature of the gas turbine has to be in the order of 580­

590°C. In case of the lignite oxyfuel boiler with OTM, the retentate stream is directly

expanded after leaving the membrane reactor. Here, the exhaust gas of the turbine is

used to preheat feed water from the water/steam cycle. Therefore the exhaust

temperature should be lower to reduce the temperature difference between feed water

and exhaust gas. Due to this circumstance the exhaust temperature of the turbine is

around 300°C. The further processing of the retentate stream does not impact the

conditions of the feed stream. For both configurations, IGCC­OTM and LFB­OTM,

the inlet conditions are considered to be identical. Althogh there will be a small

difference in the molar compostion because different fuels are used for both

configurations.

However, the calculation of the oxygen flux through the OTM reactor is not

impacted because only the oxygen partial pressure on both sides of the membrane

reactor is considered. Therefore the parametric studies are conducted for the OTM

reactor as part of the IGCC­OTM cycle. The conditions of the feed stream for the

IGCC­OTM cycle are as follows:

6 Model l ing of In tegra ted Membrane Reactors 101

Mass flow 654.4 kg/s

Temperature 900°C

Pressure 16.54 bar

Molar composition of the feed stream

Nitrogen 75.96%

Oxygen 15.35%

Water 7.53%

Argon 0.94%

Carbon dioxide 0.22%

Concerning the feed temperature of the OTM reactor it is emphasised that this

temperature is arbitrary. As previously mentioned the feed temperature may range

from 800 to 1000°C. Higher feed temperatures are desirable as they promote a higher

permeation flux through the membrane (see figure 6­5), however the temperature is

limited by the material properties of the membrane. In this context it seems reasonable

to assume the feed temperature with 900°C. A sensitivity analysis of the feed

temperature is conducted to quantify the impact of this parameter. The results of the

variation of the feed temperature can be found in the appendix A.5. For all simulations

of the IGCC­OTM cycle and the lignite fired oxyfuel boiler cycle a fixed feed

temperature of 900°C is used. Therefore the results of the feed temperature variation

are presented in the appendix.

The amount of the oxygen separated in the OTM reactor (around 32.2 kg/s)

corresponds to the required oxygen mass flow rate for the gasifier in the IGCC­OTM

cycle. The fixed conditions of the feed stream and the separated mass flow of oxygen

are used as input for the calculation of the OTM reactor. Regarding the feed

temperature it should be emphasised that, of course, the permeation through the

membrane was increased if the feed temperature were higher. With these inputs the

overall membrane surface area of the OTM reactor was computed by iterative

calculations of heat and mass transfer as described in the previous sections. LP or IP

steam extracted from the steam turbine is used as sweep stream for the OTM reactor.

Although the steam temperature varies with the pressure level where it is extracted

from the steam turbine of the water/steam cycle, the inlet temperature of the sweep

stream is set to 350°C. This temperature can be obtained by internal heat exchange

with the oxygen rich permeate stream leaving the OTM reactor. The conditions – mass

flow and pressure – of the sweep stream are varied as follows:

102 6 .5 Parametric Studies on the Membrane Reactors

Pressure of the sweep stream: 2, 5, 10 and 15 bar

Mass flow of the sweep stream: from 2 to 24% of the feed stream

The steam mass flow (sweep stream) is expressed relative to the feed stream, so

that conclusions can be drawn independently of the absolute value of the feed stream.

Due to the pressure of the feed stream remaining constant for all calculations, the

difference in total pressure ranges from 1 to 14 bar.

Variation of the overall heat transfer coefficient

The overall heat transfer coefficient is calculated with 152 W/(m2 K) for the OTM

reactor. Nevertheless, the overall heat transfer coefficient is varied to see how the mass

transport is affected by this parameter. The overall heat transfer coefficient is varied

from 50 to 300 W/(m2 K). For the different sweep pressures from 2 to 15 bar, the

resulting membrane surface areas are illustrated in figure 6­7 to figure 6­10. In all four

figures the inlet temperature of the feed stream is 900°C. The mass flow of the sweep

stream ranges from 12 to 24% of the feed stream. For all sweep pressures, the overall

heat transfer coefficient shows the same trend. For all cases, above 200 W/(m2 K)

there is no impact at all on the membrane surface area; below 200 W/(m2 K) there is a

small influence on the membrane surface area. The membrane surface area increases

slightly for a heat transfer coefficient of below 200 W/(m2 K) because the mean

temperature difference between feed and sweep stream is larger when the heat transfer

coefficient is small. The larger temperature difference leads to a lower mean

temperature at each element of the membrane reactor because the mean temperature is

assumed to be the arithmetic average of feed and sweep temperature, and a lower

mean temperature results eventually in a larger membrane surface area.

Comparing the four different sweep pressures from 2 to 15 bar, figure 6­7 to figure

6­10, it can be seen that the difference in oxygen partial pressure determines the order

of magnitude of the membrane surface area. For the lowest sweep pressure of 2 bar,

the membrane surface area is for all mass flows below 80,000 m2, whereas for the

highest sweep pressure of 15 bar, the membrane surface area increases up to

400,000 m2. The higher the total pressure of the sweep stream, the stronger the impact

of the sweep mass flow rate. The oxygen partial pressure increases with higher total

pressure of the sweep stream, whereas it decreases with larger mass flow rates of the

sweep stream. A larger mass flow of the sweep stream can partly compensate the

higher total pressure.

6 Model l ing of In tegra ted Membrane Reactors 103

50 100 150 200 250 300

6

8

10

12

14

16

18

20

22

24

26

28M

em

bra

ne s

urf

ac

e a

rea / 1

04 m

2

Overall heat transfer coefficient / W/(m2 K)

Feed stream mass flow ratem

Sweep = 12 %

mSweep

= 16 %

mSweep

= 20 %

mSweep

= 24 %

TFeed

= 900°C

pSweep

= 2 bar

Figure 6-7: Membrane surface area vs. heat transfer coefficient for varies sweep stream

mass flow rates (2 bar sweep pressure and a feed temperature of 900°C).

50 100 150 200 250 300

6

8

10

12

14

16

18

20

22

24

26

28

Mem

bra

ne s

urf

ac

e a

rea / 1

04 m

2

Overall heat transfer coefficient / W/(m2 K)

Feed stream mass flow ratem

Sweep = 12 %

mSweep

= 16 %

mSweep

= 20 %

mSweep

= 24 %

TFeed

= 900°C

pSweep

= 5 bar

Figure 6-8: Membrane surface area vs. heat transfer coefficient for varies sweep stream

mass flow rates (5 bar sweep pressure and a feed temperature of 900°C).

104 6 .5 Parametric Studies on the Membrane Reactors

50 100 150 200 250 300

6

8

10

12

14

16

18

20

22

24

26

28M

em

bra

ne s

urf

ace a

rea

/ 1

04 m

2

Overall heat transfer coefficient / W/(m2 K)

TFeed

= 900°C

pSweep

= 10 bar

Feed stream mass flow ratem

Sweep = 12 %

mSweep

= 16 %

mSweep

= 20 %

mSweep

= 24 %

Figure 6-9: Membrane surface area vs. heat transfer coefficient for varies sweep stream

mass flow rates (10 bar sweep pressure and a feed temperature of 900°C).

50 100 150 200 250 30018

20

22

24

26

28

30

32

34

36

38

40

42

Mem

bra

ne s

urf

ac

e a

rea / 1

04 m

2

Overall heat transfer coefficient / W/(m2 K)

TFeed

= 900°C

pSweep

= 15 bar

Feed stream mass flow ratem

Sweep = 16 %

mSweep

= 20 %

mSweep

= 24 %

Figure 6-10: Membrane surface area vs. heat transfer coefficient for varies sweep stream

mass flow rates (15 bar sweep pressure and a feed temperature of 900°C).

6 Model l ing of In tegra ted Membrane Reactors 105

Variation of pressure and mass flow rate of the sweep stream

Assuming a constant feed temperature to the OTM reactor, the pressure and the

mass flow of the sweep stream determine the operating conditions of the OTM reactor.

Therefore, both parameters have been varied to investigate their impact on the oxygen

permeation. The discussion of the results of the parametric studies can be done from

different point of views:

Minimising the membrane surface area.

The size of the OTM reactor is determined by the required membrane surface

area. In turn, the size gives the investment costs of the OTM reactor. Therefore,

from an economic point of view, the goal could be to allow the reactor to have a

certain size to restrict its investment costs. This approach would lead to a minimal

mass flow and/or to maximal allowable pressure of the sweep stream without

exceeding the limit of the reactor size.

Limiting the sweep pressure to a minimum value.

According to rules for designing the membrane reactor, a requirement could be

that the difference in total pressure is limited to a maximum value in order to

avoid too high mechanical load on the membrane material. Due to the fact that the

pressure of the feed stream is given by the pressure ratio of the gas turbine cycle,

such a requirement would restrict the sweep pressure to a minimum value. Recall,

the higher the sweep pressure, the smaller the difference in total pressure across

the membrane.

Limiting the sweep stream mass flow to a maximum value.

In case of the OTM reactor, the IP or LP steam is used as sweep stream (the steam

is being extracted from the steam turbines). Each kilogram of steam that is

extracted reduces the power output of the overall power generation cycle;

therefore, the goal is that the mass flow of the extracted steam is as small as

possible. This could result in the requirement that the amount of steam is limited

to a maximum value. At the same time a low mass flow of steam on the sweep

side of the membrane reactor causes the oxygen partial pressure to be high, which,

in turn, lowers the permeation flux through the membrane.

The last two aspects conflict with the aim of achieving a high permeation flux

through the membrane. Therefore, the pressure and the mass flow rate of the sweep

stream are varied to give an understanding of how both parameters impact the

permeation flux through the membrane. The impact of the operating conditions on

overall power generation cycles are described in chapter 7. In this section the implica­

106 6 .5 Parametric Studies on the Membrane Reactors

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 160.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

Sweep stream mass flow rate 5 % 14 % 8 % 16 % 9 % 20 % 12 % 24 %

TFeed

= 900°C

Mem

bra

ne s

urf

ace a

rea / 1

06 m

2

Sweep stream pressure / bar

Figure 6-11: Membrane surface area vs. feed stream mass flow rate for different sweep

stream pressures at a feed temperature of 900°C

2 4 6 8 10 12 14 16 18 20 22 240.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

Sweep stream pressure 2 bar 5 bar 7.5 bar 10 bar 12.5 bar 15 bar

TFeed

= 900°C

Mem

bra

ne s

urf

ace a

rea / 1

06 m

2

Sweep stream mass flow rate / %

Figure 6-12: Membrane surface area vs. feed stream pressure for different sweep stream

mass flow rates at a feed temperature of 900°C.

6 Model l ing of In tegra ted Membrane Reactors 107

tions on the membrane reactor are discussed. The results of the parametric studies are

shown in figure 6­11 and figure 6­12, respectively. The same calculations are shown in

two different ways. In figure 6­11 the membrane surface area is plotted against the

sweep stream pressure for various sweep stream mass flow rates, while in figure 6­12

the area is plotted against the sweep stream mass flow for different sweep pressures.

For instance the curves with a constant sweep pressure in figure 6­12 show the

minimal acceptable amount of mass flow rate. Taking a sweep pressure of 10 bar as an

example, it can be seen that the sweep mass flow rate has to be equal to or greater than

9% of the feed stream. If the sweep mass flow rate were smaller, the oxygen partial

pressure would be too low on the sweep side of the membrane, so that the transport

through the membrane would be eliminated.

Comparing the lowest sweep pressure (2 bar) and the highest mass flow rate (24%)

it can be concluded that the sweep pressure has the stronger impact on the membrane

surface area. If the sweep pressure is set to 2 bar, the membrane surface area is smaller

than 100,000 m2 for mass flow rates from 5 to 24%. By contrast, if the mass flow rate

is constant at 24%, the sweep pressure is 5 bar, then the membrane surface area is in

all cases larger than 100,000 m2.

If the approach was to restrict the membrane surface area to a certain value, for

instance 200,000 m2, figure 6­12 shows that in this case the sweep pressure would

have to be below 10 bar. For a very low sweep pressure of 2 bar the membrane surface

area changes only slightly over nearly the whole range of mass flow rates. If the sweep

pressure were higher than 10 bar, the mass flow rate would need to be larger than 15%

of the feed stream. A large amount of the steam as sweep stream impacts adversely the

overall power generation cycle. Were the sweep pressure limited to a minimum value

(e.g. 10 bar), it can be seen from figure 6­11 that the sweep stream mass flow rate

would need to be larger than 12%. Furthermore it can be seen that either the mass flow

rate needs to be larger than 20% to keep the reactor size below 200,000 m2 or the

reactor size would range from 200,000 to 700,000 m2, for example for a mass flow rate

of 14%. Assuming the surface­volume­ratio of 750 m2/m3, this range of membrane

surface area would result in a reactor volume of 260­930 m3. If, for reasons of the

overall power generation cycle, the sweep stream mass flow rate was limited to a

maximal value (e.g. 10%), the sweep pressure would need to be lower than 10 bar (see

figure 6­12). In case the mass flow rate were even more restricted, only sweep

pressures of 5 to 2 bar would be then feasible. In those scenarios the pressure

difference across the membrane would be more than 11 or 14 bar, respectively.

108 6 .5 Parametric Studies on the Membrane Reactors

Summarising the described scenarios, a conclusion regarding the operating

conditions of the OTM reactor is that the reactor needs to be capable at withstanding a

pressure difference of at least 5 to 10 bar. If the reactor were only able to cope with a

smaller pressure difference (< 5 bar), then large quantities of steam as sweep gas

would be required to limit the reactor size or even to make operation of the reactor

feasible at all. Furthermore, if for economical reasons the size of the membrane reactor

must be as small as possible, the sweep pressure has to be even smaller than 5 bar.

As previously discussed, the size of the membrane reactor may become large if

only a small difference in total pressure across the membrane is applied. Due to this it

must be emphasised again that a higher feed temperature could help to lower the

increase of the required membrane surface area. The corresponding graph to figure

6­12 is Figure A5­6 (given in appendix A.5) for a feed temperature of 1000°C. For this

higher feed temperature, the membrane surface area would be smaller than 400,000 m2

in all cases. The issue of a higher feed temperature is the same as for a large pressure

difference across the membrane, that the membrane material needs to be able to cope

with such challenging operating conditions.

Development of oxygen partial pressure along the reactor length

The difference in oxygen partial pressure is the driving force for the mass transport

through the membrane. Therefore, for some of the described parameter variations, the

oxygen partial pressure along the reactor length is exemplarily shown in this section.

The intention of this illustration is to comprehensively explain how the oxygen partial

pressure develops for both streams along the membrane reactor and to show why some

combinations of pressure and mass flow rate of the sweep stream are physically not

feasible.

Two different cases with a constant sweep pressure and two cases with a constant

mass flow rate of the sweep stream are shown in figure 6­13 & figure 6­14, and figure

6­15 & figure 6­16, respectively. As previously described the OTM reactor is divided

into 75 discrete elements. The normalised reactor length is represented by these 75

elements. The numbering of the elements starts on the side where the feed stream

enters the reactor (reactor length “0”). Since the OTM reactor is considered as a

counter­flow device, the sweep stream enters the reactor at the opposite side (reactor

length “1”).

In all four cases the feed temperature is set to 900°C. In figure 6­13 the sweep

pressure is constant at 5 bar; the sweep stream mass flow rates is varied from 5 to 24%.

6 Model l ing of In tegra ted Membrane Reactors 109

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.00.0

0.5

1.0

1.5

2.0

2.5

3.0O

xyg

en

part

ial

pre

ssu

re

/ b

ar

Reactor length (normalised)

TFeed

= 900°C

pSweep

= 5 bar

Retentate Permeatem

Sweep = 5 %

mSweep

= 8 %

Retentate Permeatem

Sweep = 12 %

mSweep

= 16 %

Retentate Permeatem

Sweep = 20 %

mSweep

= 24 %

Figure 6-13: Development of oxygen partial pressure along the normalised length of the

membrane reactor (for a feed temperature of 900°C and a sweep stream

pressure of 5 bar).

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.00.0

0.5

1.0

1.5

2.0

2.5

3.0

Oxyg

en

part

ial

pre

ssu

re

/ b

ar

Reactor length (normalised)

Retentate Permeatem

Sweep = 9 %

mSweep

= 12 %

mSweep

= 16 %

mSweep

= 20 %

mSweep

= 24 %

TFeed

= 900°C

pSweep

= 10 bar

Figure 6-14: Development of oxygen partial pressure along the normalised length of the

membrane reactor (for a feed temperature of 900°C and a sweep stream

pressure of 10 bar).

110 6 .5 Parametric Studies on the Membrane Reactors

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.00.0

0.5

1.0

1.5

2.0

2.5

3.0O

xyg

en

pa

rtia

l p

res

su

re

/ b

ar

Reactor length (normalised)

TFeed

= 900°C

mSweep

= 9 %

Retentate Permeatep

Sweep = 2 bar

pSweep

= 5 bar

pSweep

= 10 bar

Figure 6-15: Development of oxygen partial pressure along the normalised length of the

membrane reactor for a constant sweep stream mass flow rate of 9% (for feed

temperature of 900°C).

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.00.0

0.5

1.0

1.5

2.0

2.5

3.0

Ox

yg

en

pa

rtia

l p

res

su

re

/ b

ar

Reactor length (normalised)

TFeed

= 900°C

mSweep

= 14 %

Retentate Permeatep

Sweep = 2 bar

pSweep

= 5 bar

pSweep

= 10 bar

pSweep

= 15 bar

Figure 6-16: Development of oxygen partial pressure along the normalised length of the

membrane reactor for a constant sweep stream mass flow rate of 14% (for feed

temperature of 900°C).

6 Model l ing of In tegra ted Membrane Reactors 111

In figure 6­14 the sweep pressure is assumed to be 10 bar. Due to the higher sweep

pressure the mass flow rate is varied from 9 to 24%. For the next two graphs a constant

sweep stream mass flow rate is assumed. In figure 6­15 a mass flow rate of 9% is used

and the sweep pressure changes from 2 to 10 bar, whereas in figure 6­16 the mass flow

is 14% and the sweep ranges from 2 to 15 bar. Further cases are given in appendix

A.5.

In all four charts the feed temperature is set to 900°C. All curves for the feed

stream have the same start and end points because their pressure and composition at

the inlet remain constant for all variations. Due to the fact that the amount of the

separated oxygen is always the same, that the composition and hence the oxygen

partial pressure at the outlet is the same in all cases. Of course, for all cases the oxygen

partial pressure is zero at the inlet because pure steam enters the membrane reactor.

The development and the oxygen partial pressure at the outlet is determined by a

combination of pressure and mass flow rate of the sweep stream. The oxygen partial

pressure on the feed side decreases from 2.5 bar to 1.75 bar. For a sweep pressure of

5 bar and the mass flow of 5%, the oxygen partial pressure increases to 1.75 bar at the

exit of the OTM reactor (see figure 6­13). In contrast the oxygen partial pressure is

2.25 bar, in case of a sweep pressure of 10 bar and a mass flow rate of 9% (see figure

6­14). The small difference at the reactor length “0” makes clear that a smaller mass

flow rate of the sweep is physically not feasible because the difference in oxygen

partial pressure would be eliminated if the sweep mass flow were further reduced.

Figure 6­15 and figure 6­16 give the same message as the first two graphs. For

instance with a sweep mass flow rate of 14% a sweep pressure of 15 bar is feasible,

although the difference in oxygen partial pressure is small at the outlet of the sweep

stream (see figure 6­16). This combination would lead to the largest membrane surface

area of more than 700,000 m2, compare figure 6­16 and figure 6­12. In case of a sweep

mass flow rate of 9%, the highest feasible pressure is only 10 bar (see figure 6­15).

For the sake of completeness it is emphasised that the variation of the pressure and

mass flow rate of the sweep stream is also conducted for a feed temperature of 800 and

1000°C. The appropriate graphs are given in appendix A.5 (in Figure A5­5 to Figure

A5­8). For a lower feed temperature the variation of the membrane surface area

increases slightly. This is due to a stronger impact of the feed temperature for an

identical variation of oxygen partial pressure. The impact of both pressure and mass

flow rate of the sweep stream is slightly lowered because a higher feed temperature

promotes the permeation through the membrane.

112 6 .5 Parametric Studies on the Membrane Reactors

6.5.2 Results for the Hydrogen-selective Membrane Reactor

The feed stream of the hydrogen­selective membrane reactor is part of the syngas

treatment of the IGCC­H2 cycle. Therefore the conditions of the feed stream are

primarily determined by the gasification process but also by the subsequent cleaning

(COS­hydrolysis and desulphurisation) and further treatment (CO­shift reaction).

Regarding the temperature of the feed stream it should be emphasised that this

parameter is even more restricted than the feed temperature of the OTM reactor.

In the case of the oxygen transport membrane this reactor is integrated into a gas

turbine cycle, where the feed temperature could be adjusted by the outlet temperature

of the first combustion chamber. In the IGCC­H2 cycle the hydrogen­selective

membrane reactor is located after the CO­shift reactor from which the syngas leaves

this reactor with a temperature around 150°C (see figure 4­6). The feed temperature of

the hydrogen­selective membrane reactor is more restricted by the cycle layout and

required internal heat exchange than by limitation of the membrane material. Before

the syngas stream enters the hydrogen­selective membrane reactor it needs to be

heated. The heating is only possible by means of internal heat exchange. In a first step,

the syngas is heated to 350°C utilising internal (but indirect) heat exchange with

syngas leaving the syngas cooler. In a second step, the syngas is further heated by

internal (but indirect) heat exchange with heat from the supplementary burner, which

is located after the hydrogen­selective membrane reactor (see figure 4­6). After the

second stage of internal heat exchange the syngas temperature is assumed to be 600°C.

A temperature level of 600°C is a compromise between a high temperature to achieve

a reasonable high permeation flux and available heat from the supplementary burner.

Due to the aforementioned reasons the conditions of the feed stream is set constant

for all configurations. The conditions of the feed stream for the IGCC­H2 are:

Mass flow 157.6 kg/s

Temperature 600°C

Pressure 30 bar

Molar composition of the feed stream

Hydrogen 37.8%

Carbon dioxide 27.4%

Water 28.8%

Nitrogen 3.7%

Carbon monoxide 1.9%

Argon 0.4%

6 Model l ing of In tegra ted Membrane Reactors 113

The hydrogen partial pressure is higher than 12 bar at the inlet of the membrane

reactor. Nevertheless, the major challenge for the hydrogen­selective membrane

reactor is that the hydrogen entering the reactor with the feed stream has to be

separated completely from the feed stream. The reason for this is that the hydrogen

accounts for nearly all of the heat input of the syngas to the combined cycle and each

molecule of hydrogen, which is not separated in the membrane reactor, represents a

heat loss for the overall power generation cycle. In the simulation for numerical

reasons it is assumed that 99% of the hydrogen is separated by the membrane reactor.

Therefore, the development of hydrogen partial pressure for both streams of the

membrane reactor differs completely to that of the OTM reactor. Although 99% of the

hydrogen is separated in the hydrogen­selective membrane reactor, an supplementary

burner is required to utilise the heat from the remaining hydrogen and, more

importantly, from the carbon monoxide, which is also not completely converted to

carbon dioxide in the CO­shift reactor.

Due to the high degree of separation of hydrogen in the membrane reactor, a large

difference in hydrogen partial pressure across the membrane is required to assure a

stable operation of the separation process. As previously mentioned in context of the

OTM reactor, both a large mass flow and a low pressure of the sweep stream increase

the difference in hydrogen partial pressure across the membrane. The sweep stream

mass flow rate in the hydrogen­selective membrane reactor is relative to the feed

stream much higher compared to the OTM reactor.

Nitrogen is used as sweep gas for the hydrogen­selective membrane reactor. The

nitrogen is generated by the ASU as a by­product of the oxygen production, which is

used as oxidiser in the gasification process. The advantage of using nitrogen as sweep

gas is that from an energy expenditure point of view, the nitrogen is freely available

because it is a by­product from the oxygen production. Therefore the amount of

nitrogen does not impact the overall power generation cycle. Furthermore, it is not

necessary to separate the mixture of sweep (nitrogen) and permeate (hydrogen)

streams after leaving the membrane reactor because the hydrogen needs to be diluted

anyway before it is combusted in the combustion chamber of the gas turbine. In

comparison with the OTM reactor, the nitrogen supplied by the ASU needs to be

pressurised before it enters the membrane reactor. The conditions of the sweep stream

are as follows:

Pressure of the sweep stream: 5, 7.5, 10 and 12.5 bar

Mass flow of the sweep stream: from 10 to 50% of the feed stream

114 6 .5 Parametric Studies on the Membrane Reactors

The large difference in hydrogen partial pressure across the membrane, which is

required to separate 99% of the hydrogen from the feed stream, can only be obtained

by having a large mass flow rate of the sweep stream. Although the difference in total

pressure across the membrane ranges from 8.5 to 20 bar – higher than for the OTM

reactor – the large mass flow rate of nitrogen is still necessary to reduce the hydrogen

partial pressure on the sweep side of the membrane. Both variations of pressure and

mass flow of the sweep stream are done to investigate the impact on the permeation

flux through the membrane. Less variation of the sweep stream conditions is

conducted because nearly all of the hydrogen is separated in the hydrogen­selective

membrane reactor, which narrows the possibilities of varying the operating conditions.

Variation of pressure and mass flow rate of the sweep stream

The variation of pressure and mass flow rate of the sweep stream is done in a

similar way as for the OTM reactor. The conditions of the feed stream remain

constant, whereas both, pressure and mass flow rate of the sweep stream are varied to

investigate their impact on the hydrogen permeation flux through the membrane. The

operating conditions of the hydrogen­selective membrane reactor differ to those of the

OTM reactor because all of the hydrogen ( 99%) is separated from the feed stream. In

addition, the pressure of the feed stream is determined by the pressure of the gasifier

and not by the gas turbine cycle; therefore the feed pressure is 25 bar instead of

16.5 bar as for the OTM reactor. Nevertheless, the results can be discussed in the same

manner as for the OTM reactor. The results of the parametric studies can be

interpreted from different perspectives:

Minimising the membrane surface area.

The size of the hydrogen­selective membrane reactor is determined by the

required membrane surface area. In turn, the size gives the investment costs of the

membrane reactor. Therefore, from an economical point of view, the goal could be

to allow the reactor to have a certain size to restrict its investment costs. This

approach would lead to a minimal mass flow and/or to maximal allowable

pressure of the sweep stream, respectively, without exceeding the limit of the

reactor size.

Limiting the sweep pressure to a minimum value.

The high pressure of the feed stream leads to a large difference in total pressure

across the membrane. According rules for designing the membrane reactor, a

requirement could be that the difference in total pressure is limited to a maximum

6 Model l ing of In tegra ted Membrane Reactors 115

value in order to avoid too high mechanical loads for the membrane material. Due

to the constant pressure of the feed stream, the lower the sweep stream pressure,

the larger the difference in total pressure across the membrane. A limitation of the

difference in total pressure would finally lead to a minimum allowable value of

the sweep pressure.

The amount of used nitrogen as sweep gas.

Since nitrogen used as sweep gas in the membrane reactor is a by­product of the

ASU, the amount of sweep gas has not a direct impact on the overall power

generation cycle. With respect to the expenditure of energy the ‘production’ of

nitrogen is neutral to the power output of the overall power generation cycle.

Pressure and mass flow rate of the sweep stream are varied to understand how both

parameters impact the permeation flux through the membrane. The implications on the

membrane reactor alone are discussed. The variation of pressure and mass flow rate of

the sweep stream are illustrated in figure 6­17 and figure 6­18, respectively. The

results are presented in the same way as for the OTM reactor. In figure 6­17 the

membrane surface area is plotted versus the sweep stream pressure for various sweep

stream mass flow rates, whereas in figure 6­18 the area is plotted against the sweep

stream mass flow for different sweep pressures. It can be seen from figure 6­17 that

large sweep stream mass flow rates are required to cover the full range of sweep

stream pressures. For the investigated range of pressure and mass flow rates, the

membrane surface varies from 500,000 to nearly 1,600,000 m2. Assuming the same

surface­volume­ratio as for the OTM reactor of 750 m2/m3, this range in membrane

surface area would result in a reactor volume of around 660­2,100 m3. This gives an

impression about the size of the hydrogen­selective membrane reactor. In all four

figures the feed temperature is 600°C. Figure 6­17 shows that only by means of large

mass flow rates (40 and 50%) it is possible to cover a range in sweep stream pressure

from 2.5 to 15 bar. If the sweep stream mass flow is smaller than 30%, the sweep

pressure has to be below 10 bar. Only with a lower sweep pressure of 10 bar is the

difference in hydrogen partial pressure large enough to separate all hydrogen from the

feed stream. In this case the difference in total pressure is larger than 15 bar. Therefore

it is likely that the mass flow rate has to be larger than 30% of the feed stream to limit

the mechanical load on the membrane material. For a sweep stream mass flow rate of

40 and 50%, the membrane surface area increases approximately linearly with the

sweep pressure. The membrane surface area ranges from 800,000 to 1,200,000 m2.

The large membrane surface area is a result of high degree of separation in the

hydrogen­selective membrane reactor.

116 6 .5 Parametric Studies on the Membrane Reactors

Figure 6-17: Membrane surface area vs. sweep stream pressure for various sweep stream

mass flow rates. The sweep stream mass flow is expressed in percentage of the

feed stream.

Figure 6-18: Membrane surface area vs. sweep stream mass flow for various sweep stream

pressures. The sweep stream mass flow is expressed in percentage of the feed

stream.

6 Model l ing of In tegra ted Membrane Reactors 117

Development of hydrogen partial pressure along the reactor length

The difference in oxygen partial pressure is the driving force for the mass transport

through the membrane. Therefore, for a constant sweep stream mass flow rate of 40%

and varying sweep pressures, the hydrogen partial pressure along the reactor length is

shown in this section. The intention of this illustration is to explain how the hydrogen

partial pressure develops and how it differs in comparison to the OTM reactor. Figure

6­19 shows the development of the hydrogen partial pressure along the reactor length

for the retentate and the permeate stream. The hydrogen­selective membrane reactor is

divided into 128 discrete elements. The normalised reactor length represents these 128

elements. The numbering of the elements starts on that side where the feed stream

enters the reactor (reactor length “0”). Since the hydrogen­selective membrane reactor

is considered a counter­flow apparatus, the sweep stream enters the reactor at the

opposite side (reactor length “1”).

In figure 6­19 the temperature of the feed stream is set to 600°C. The mass flow

rate of the sweep stream is 40% of the feed stream and its pressure is varied from 5 to

12.5 bar. All curves for the feed stream have the same starting point and end at a

hydrogen partial pressure close to zero because 99% of the hydrogen is separated from

Figure 6-19: Development of hydrogen partial pressure along the membrane reactor for the

retentate (feed) and the permeate (sweep) streams.

118 6 .5 Parametric Studies on the Membrane Reactors

the feed stream. Of course, for all cases the hydrogen partial pressure is zero at the

inlet because pure nitrogen enters the membrane reactor. The development and the

hydrogen partial pressure at the outlet is determined by a combination of pressure and

mass flow rate of the sweep stream. The hydrogen partial pressure on the feed side is

approximately 12.5 bar at the inlet of the membrane reactor. The hydrogen partial

pressure of the sweep stream ranges from 3 to 8 bar at the exit of the membrane

reactor. The difference in partial pressure is much larger than that of the OTM reactor

due to the high degree of separation of hydrogen from the feed stream, compare figure

6­13 and figure 6­19. The high degree of separation in the hydrogen­selective

membrane reactor is responsible for the large required difference in partial pressure.

7 Simula t ion of the Analysed Power Genera t ion Cycles 119

7 Simulation of the Analysed Power Generation Cycles

7.1 Simulation of the IGCC Cycles

7.1.1 Design Point Comparison of all Investigated Configurations

The results of the simulations of all investigated configurations of the IGCC cycles

are described in the following.12 The benchmarking of the different cycles is carried

out for baseload operation in the design­point of each configuration. Since the focus is

on the thermodynamic potential of the cycles investigated, no analysis of partload

behaviour or transient operation is conducted. The following IGCC configurations,

which are presented in section 4.2, are benchmarked:

IGCC­REF

IGCC­CAP

IGCC­OTM

IGCC­H2

The first cycle is IGCC­REF because it is without CO2 capture and is therefore

used as a reference cycle for the other configurations with CO2 capture. The results of

all cycles are presented in the same way. For each cycle a breakdown of power and

efficiency summarises the main contributors of the power generation cycle. For all the

cycles investigated the same mass flow of coal is assumed, thus for all cycles the heat

input into the cycle is identical. The power of each entity is given as absolute power in

megawatt as well as in relation to the heat input of the configuration. The relative

expression corresponds to the difference in net efficiency of each contributor. Table

7­1 shows the balance of power and efficiency for the IGCC­REF cycle. The sum of

shaft power of the turbomachineries of the gas turbine and the steam turbine delivers a

power of 54.41%. The power at the generator is slightly lower (53.38%) because

mechanical and generator losses are considered. The energy consumption and plant

auxiliary power reduce the generated power and lead to the net power output of the

power island (52.26%). The net power output of the overall power plant is

significantly reduced due to two large contributors related to the gasification process.

First, the expenditure of energy for the cryogenic ASU reduces the power output by

3.87% points. Before the syngas can be combusted in the gas turbine it needs to be

diluted with nitrogen to assure a stable combustion process. The nitrogen has to be

compressed before it can be mixed with the syngas, see figure 4­1. Secondly, due to the

12 This work has been partly published in [204, 205].

120 7 .1 Simula t ion of the IGCC Cycles

Table 7-1: Power and efficiency balance of the cycle: IGCC-REF.

LHV MJ/kg 25.174 Share of LHV Fuel flow kg/s 43.21

Fuel LHV MW 1087.77 100.00%

Gas turbine expander MW 667.60 61.37%

Steam turbines (HP, IP and LP) MW 198.97 18.29%

Gas turbine compressor MW ­274.70 ­25.25%

Turbomachinery shaft power MW 591.88 54.41%

Turbomachinery mechanical loss MW ­2.37 ­0.22%

Turbomachinery generator loss MW ­8.84 ­0.81%

Turbomachinery generator terminal output MW 580.67 53.38%

Pumps of the water/steam cycle kW ­3555 ­0.33%

Power island gross power output MW 577.11 53.05%

Plant auxiliary power MW ­8.66 ­0.80%

Net plant power island output MW 568.45 52.26%

Work oxygen production and compression MW ­42.07 ­3.87%

Work nitrogen compression (syngas mixing) MW ­36.05 ­3.31%

Work for any other related auxiliary processes MW ­0.002 0.00%

Net plant power output MW 490.34 45.08%

Specific CO2 emissions g/kWhe 757

large amount of nitrogen necessary for the dilution, the compression lowers the power

output by 3.31% points. In absolute numbers, the power for the ASU and the

compression of the nitrogen reduce the power output by 78 MW. The net power output

of the IGCC­REF cycle is 490.34 MW or in percentage to the heat input 45.08%. The

relative expression corresponds to the net efficiency of the cycle. Because no capture

of CO2 takes place the specific CO2 emissions of the IGCC­REF cycle are 757

g(CO2)/kWhe.

The results of the cycle IGCC­REF­ASU where the ASU is integrated into the gas

turbine cycle is in the same way shown in Appendix A.6 in table A6­1. The net

efficiency increases by around 0.7% points due to savings on the expenditure of

energy for the ASU.

The results of the cycle IGCC­CAP are shown in table 7­2. The net power output is

significantly reduced by around 108 MW because of the CO2 capture. Three

contributors to the reduction of the net power output are: (i) additional compression

work for the separated CO2, (ii) reduced syngas mass flow due to the separation of the

CO2 and (iii) steam is extracted from the steam turbine to be utilised in the CO­shift reac­

7 Simula t ion of the Analysed Power Genera t ion Cycles 121

Table 7-2: Power and efficiency balance of the cycle: IGCC-CAP.

LHV MJ/kg 25.174 Share of LHV Fuel flow kg/s 43.21

Fuel LHV MW 1087.77 100.00%

Gas turbine expander MW 619.74 56.97%

Steam turbines (HP, IP and LP) MW 147.86 13.59%

Gas turbine compressor MW ­259.28 ­23.84%

Turbomachinery shaft power MW 508.33 46.73%

Turbomachinery mechanical loss MW ­2.033 ­0.19%

Turbomachinery generator loss MW ­7.594 ­0.70%

Turbomachinery generator terminal output MW 498.70 45.85%

Pumps of the water/steam cycle kW ­3520 ­0.32%

Power island gross power output MW 495.18 45.52%

Plant auxiliary power MW ­7.428 ­0.68%

Net plant power island output MW 487.75 44.84%

Work CO2 compression MW ­24.48 ­2.25%

Work oxygen production and compression MW ­42.07 ­3.87%

Work nitrogen compression (syngas mixing) MW ­39.85 ­3.66%

Work for any other related auxiliary processes MW ­0.010 0.00%

Net plant power output MW 381.35 35.06%

Specific CO2 emissions g/kWhe 78

tor and the CO2 capture unit. The compression work is required to pressurise the

separated CO2 to 110 bar. The reduced syngas stream causes a smaller mass flow of

fuel for the gas turbine in combined cycle configuration. Therefore, both the power

output of the gas turbine and the steam turbines is reduced. The smaller exhaust gas

mass flow of the gas turbine results in a smaller steam mass flow produced in the

HRSG. In addition to the smaller mass flow through the steam turbines, some steam is

extracted from the turbines for the syngas treatment. The power output of gas turbine

and steam turbines is 84 MW lower than that of the IGCC­REF cycle. This reduction

results in a net power output of the power island of 487.75 MW or 44.84%. The

additional 24 MW for the compression of the CO2 sum up to the mentioned reduction

in power output of 108 MW. The net power output of the IGCC­CAP cycle is

381.35 MW, which corresponds to a net efficiency of 35.06%. The emitted CO2 is

reduced by nearly 90% to 78 g(CO2)/kWhe.

The results of the IGCC cycle with CO2 capture and integrated ASU, IGCC­CAP­

ASU, is also presented in appendix A.6 in Table A6­2. Same as for the IGCC­REF

122 7 .1 Simula t ion of the IGCC Cycles

cycle, the net efficiency increases by around 0.7% points due to savings on the

expenditure of energy for the ASU.

The breakdown of the IGCC cycle with integrated OTM reactor is given in table

7­3. These results are for one operating condition of the OTM reactor with a sweep

pressure of 5 bar and a mass flow rate of 6% of the feed stream. Comparing the IGCC­

OTM cycle to the IGCC­CAP it can be seen that the gas turbine power output is

further decreased. The reduction of the gas turbine power is caused by the separated

oxygen, which is extracted from the gas turbine working fluid by the OTM reactor.

The lower mass flow through the gas turbine reduces the power output. Another aspect

needs to be considered when looking at the power output of the gas turbine: as said

before, for all IGCC configurations the same fuel mass flow (coal) is assumed. For this

reason the syngas mass flow is the same for both cycles, the IGCC­CAP and the

IGCC­OTM. Therefore, the compressor inlet mass flow in the case of the IGCC­OTM

cycle is lower than that of the IGCC­CAP, see Table A2­1 and Table A2­2 in the

appendix. The temperature of the feed stream into the OTM reactor (between the two com­

Table 7-3: Power and efficiency balance of the cycle: IGCC-OTM.

LHV MJ/kg 25.174 Share of LHV Fuel flow kg/s 43.21

Fuel LHV MW 1087.77 100.00%

Gas turbine expander MW 575.25 52.88%

Steam turbines (HP, IP and LP) MW 153.62 14.12%

Gas turbine compressor MW ­251.11 ­23.08%

Turbomachinery shaft power MW 477.76 43.92%

Turbomachinery mechanical loss MW ­1.911 ­0.18%

Turbomachinery generator loss MW ­7.138 ­0.66%

Turbomachinery generator terminal output MW 468.72 43.09%

Pumps of the water/steam cycle kW ­3741 ­0.34%

Power island gross power output MW 464.97 42.75%

Plant auxiliary power MW ­6.975 ­0.64%

Net plant power island output MW 458.00 42.10%

Work CO2 compression MW ­24.48 ­2.25%

Work oxygen compression MW ­7.50 ­0.69%

Work nitrogen compression (syngas mixing) MW ­39.85 ­3.66%

Work for any other related auxiliary processes MW ­0.010 0.00%

Net plant power output MW 386.17 35.50%

Specific CO2 emissions g/kWhe 77

7 Simula t ion of the Analysed Power Genera t ion Cycles 123

bustion chambers of the gas turbine) is lowered by 50 K due to heat exchange with the

sweep stream. This reduction in temperature needs to be compensated for in the

second combustion chamber. If the fuel mass flow remains constant, the mass flow of

the working fluid has to be reduced. The power output of the gas turbine and steam

turbines is about 30 MW less than that for the IGCC­CAP cycle.

The mixture of oxygen and steam (on the sweep side of the OTM reactor) is heated

by the higher temperature of the feed stream along the membrane reactor. This heat is

additional heat, which is transferred from the gas turbine to the water/steam cycle. The

heat of the oxygen­steam­mixture is utilised in the HRSG by additional heat

exchangers. This effect more than compensates for the additional extraction of steam

of around 30 kg/s as sweep gas. That is the reason why, although the additional steam

is extracted from the water/steam cycle, the power output of the steam turbines

increases by almost 6 MW.

The saving on the power consumption on the ASU is 30.5 MW. The remaining

power is used for compression of the oxygen separated from the OTM reactor. The

oxygen­steam mixture leaving the reactor at 5 bar (in this case) is cooled and the steam

is mostly condensed from the mixture. The oxygen is compressed afterwards to 35 bar

before being fed to the gasifier. This compression consumes around 8 MW (see table

7­3). The sum of the lower power of the combined cycle and the positive effect on the

missing ASU leads to an overall increase of the net power output of nearly 5 MW and

the net efficiency raises by 0.44% points. The specific CO2 emissions remain almost

constant with 77 g(CO2)/kWhe.

Of course, the overall net power output and net efficiency are impacted by the

pressure and the mass flow rate of the sweep stream. Variation of those parameters are

presented in the next section. For this configuration a pressure of 5 bar and a mass

flow rate of 6% are arbitrarily chosen to show that the IGCC with integrated OTM

reactor has thermodynamically a potential to improve power output and efficiency of

the overall IGCC cycle with CO2 capture.

The results of the last configuration of IGCC cycles investigated, an IGCC with

integrated hydrogen­selective membrane reactor is shown in table 7­4. The IGCC­H2

is compared to the IGCC­CAP because this cycle is the reference cycle. It is not

meaningful to compare the two membrane­based configurations, IGCC­OTM and

IGCC­H2, with each other because in both cases the interest is on the potential of the

appropriate cycle with integrated membrane reactor in comparison to the IGCC­CAP

cycle. The power output of the gas turbine decreases by 4 MW because the fuel mass

flow is lowered by 2 kg/s, compare Table A2­1 and Table A2­2. The power output of

124 7 .1 Simula t ion of the IGCC Cycles

the steam turbines increases significantly by around 43 MW because of two effects.

First, the supplementary burner after the hydrogen­selective membrane reactor

generates a lot of heat (the temperature after the supplementary burner is above

1100°C) which is partly integrated into the water/steam cycle. This heat results in a

larger live steam mass flow and, thus in a higher power output of the steam turbines.

Secondly, no energy in the form of steam is required for the separation of the CO2.

The power for the ASU increases by more than 3 MW because some additional

oxygen is needed for the combustion in the supplementary burner – after the

hydrogen­selective membrane reactor. The total power for compression of the nitrogen

and the hydrogen, which permeates through the membrane, increases by about

32 MW. This large increase is caused by two circumstances. First the hydrogen

‘expands’ when it is transported through the membrane. The syngas enters the

membrane reactor with a pressure of 25 bar but the separated hydrogen leaves the

membrane reactor on the sweep side with a pressure of 5 bar. The lower pressure of

the sweep stream needs to be compensated by additional compression work. Secondly,

Table 7-4: Power and efficiency balance of the cycle: IGCC-H2.

LHV MJ/kg 25.174 Share of LHV Fuel flow kg/s 43.21

Fuel LHV MW 1087.77 100.00%

Gas turbine expander MW 610.42 56.12%

Steam turbines (HP, IP and LP) MW 191.75 17.63%

Gas turbine compressor MW ­254.50 ­23.40%

Turbomachinery shaft power MW 547.67 50.35%

Turbomachinery mechanical loss MW ­2.191 ­0.20%

Turbomachinery generator loss MW ­8.182 ­0.75%

Turbomachinery generator terminal output MW 537.30 49.39%

Pumps of the water/steam cycle kW ­4075 ­0.37%

Power island gross power output MW 533.22 49.02%

Plant auxiliary power MW ­7.999 ­0.74%

Net plant power island output MW 525.22 48.28%

Work CO2 compression MW ­8.99 ­0.83%

Work oxygen production and compression MW ­45.39 ­4.17%

Work nitrogen compression (syngas mixing) MW ­38.70 ­3.56%

Work for any other related auxiliary processes MW ­33.500 ­3.08%

Net plant power output MW 398.64 36.65%

Specific CO2 emissions g/kWhe 0

7 Simula t ion of the Analysed Power Genera t ion Cycles 125

the mixture of hydrogen and nitrogen leaving the hydrogen­selective membrane

reactor is heated up by the feed stream of the reactor. The nitrogen/hydrogen mixture

leaves the membrane reactor with a temperature close to 600°C. Therefore the

following compression up to 25 bar requires more power due to the high inlet

temperature of the compressor.

In total, the net power output of the IGCC­H2 cycle is 17 MW larger than that of

the IGCC­CAP cycle. As described, the positive aspect on the combined cycle is partly

deminished by a larger power consumption for the compression of the nitrogen and

hydrogen. The expenditure of energy for the compression of the CO2 decreases

because the compression starts at a higher pressure level of around 25 bar, which

corresponds to the pressure level after the hydrogen­selective membrane reactor. The

net efficiency of the IGCC­H2 cycle is 36.65%, which is 1.59% points higher than that

of the IGCC­CAP cycle. The specific CO2 emissions are 0 g(CO2)/kWhe because it is

assumed that only hydrogen is separated in hydrogen­selective membrane reactor.

Therefore all CO2 remains on the retentate stream of the reactor and is captured

afterwards.

No variation of the pressure and the mass flow rate of the sweep stream is

conducted because both parameters impact only slightly the overall IGCC cycle. For

the results of the IGCC­H2 cycle shown in table 7­4 the sweep pressure is assumed

with 5 bar and the mass flow rate of nitrogen is 20% of the feed stream.

Thermodynamically the pressure and mass flow rate of the sweep stream has only a

minor impact on the overall performance of the IGGC cycle. For stability reasons the

ratio of nitrogen and hydrogen entering the combustion chamber of the gas turbine as

fuel is chosen to be constant with 45% and 55% by mole, respectively. As shown in

figure 4­6, the mixture of hydrogen and nitrogen is additionally mixed with some

nitrogen to achieve this composition. If the sweep stream of nitrogen were smaller or

larger, the stream of nitrogen would be adjusted to achieve the desired fuel

composition. So, in total the same amount of nitrogen needs to be compressed. The

same situation occurs for the pressure of the sweep stream. After the hydrogen­

selective membrane reactor the pressure has to have a particular level. If the sweep

pressure were lower, it would need to be compressed anyway after the membrane

reactor and vice versa.

126 7 .1 Simula t ion of the IGCC Cycles

7.1.2 Variation of the Operating Conditions of the OTM Reactor

The same variation as was carried out on the pressure and mass flow of the sweep

stream in section 6.5.1 is conducted here to investigate the impact on the overall power

generation cycle. The net efficiency of the IGCC­OTM for different pressures and

mass flow rates of sweep gas is illustrated in figure 7­1. A similar variation of pressure

and mass flow rate is conducted as was done in section 6.5.1. Each single calculated

point in figure 7­1 represents a different required membrane surface area of the OTM

reactor.

In figure 7­1 the net efficiency of the IGCC­OTM cycle is plotted versus the mass

flow rate of the sweep stream. The curves represent four different pressures of the

sweep stream: 1.5, 5, 10 and 15 bar. For an easier comparison the net efficiencies of

the IGCC­CAP and the IGCC­H2 cycle are also given as horizontal lines. It can be

seen from figure 7­1 that the net efficiency decreases linearly with higher mass flow

rates of steam. The net efficiency varies from 34 to 36% for a sweep pressure of 1.5

and 5 bar, respectively and ranges between 31 and 33% for 10 and 15 bar. If the sweep

mass flow rate is larger than 10%, then the net efficiency is always lower than that of

the IGCC­CAP cycle. Only if the mass flow rate of the sweep stream is lower than 8%,

the IGCC­OTM achieves an higher net efficiency as for the IGCC­CAP cycle. As pre­

Figure 7-1: Net efficiency vs. sweep stream mass flow rate for sweep stream pressure 1.5, 5,

10, and 15 bar of the IGCC-OTM; with a constant feed temperature of 900°C.

7 Simula t ion of the Analysed Power Genera t ion Cycles 127

viously discussed in detail, the low pressure of the sweep stream results in a large

difference in total pressure across the membrane. The OTM reactor needs to be

capable of coping with such a large pressure difference. If the material of the OTM

reactor is not able to withstand this high mechanical load due to the large pressure

difference, the configuration is not attractive in comparison to the IGCC­CAP cycle.

It may be expected that the different sweep pressure would have a stronger impact

on the net efficiency of the overall IGCC cycle because the higher sweep pressure

means a larger loss in power output from the steam turbine. This effect is partly

compensated for by the required compression of the oxygen after the OTM reactor.

The oxygen separated in the OTM reactor needs to be compressed to the gasifier

pressure of 35 bar. If the sweep pressure is high, then the power consumption of the

compression is decreased. This aspect compensates partly the negative effect of a

lower power output of the steam turbines.

7.2 Simulation of the Oxyfuel Boiler Cycles

7.2.1 Design Point Benchmarking of all Investigated Configurations

Three different configuration of the lignite fired boiler cycles are investigated. The

simulation results of those three configurations are described in this section.13 For all

configurations a comparison of the performance at baseload operation in the design­

point of each configuration is carried out. The goal of the benchmarking is to show the

thermodynamic potential of each cycle, and therefore no analysis of partload

behaviour or transient operation is performed. The following boiler configurations are

benchmarked:

LFB­AIR

LFB­OXY

LFB­OTM

For reason of comparison, the first cycle is the lignite fired boiler cycle without CO2

capture, LFB­AIR, which is used as a reference cycle for the other configurations with

CO2 capture. For all three configurations the fuel mass flow (wet lignite) is kept

constant, even though, the layout of the cycles differ significantly, in particular for the

cycle with integrated OTM reactor. In the same way as for the IGCC configurations,

the results of the simulations are presented as a breakdown of power output and

13 This work has been partly published in [206].

128 7 .2 Simula t ion of the Oxyfuel Boi ler Cycles

efficiency of each cycle. The breakdown summarises the main contributors of the

power generation cycle. The power of each entity is given as absolute power in

megawatt as well as in relation to the heat input of the configuration. The relative

expression corresponds to the difference in net efficiency of each contributor. For all

cycles investigated the same mass flow of coal is assumed.

The balance of power output and efficiency for the LFB­AIR cycle is shown in

table 7­5. The heat provided to the cycle is about 1871 MW. The shaft power, the sum

of the HP, IP and LP steam turbine, is around 1006 MW. The gross power output of

approximately 994 MW corresponds to a gross efficiency of 53.1%. The plant

auxiliary is estimated with 8% of the gross power. The power required for the main

feed water pump is not included in the auxiliary power because it is directly driven by

a IP steam turbine and therefore this main contributor does not show up in the

breakdown as a separate entity. Considering the plant auxiliary power with around

80 MW, the net power output is finally 914.35 MW, which corresponds to a net

efficiency of 48.9%. The specific CO2 emissions of this cycle is 810 g(CO2)/kWhe

because no separation of CO2 is carried out.

The simulation results of the second configuration, LFB­OXY, are presented in

table 7­6. The heat input to the cycle is identical to the LFB­AIR cycle, but the shaft

power increases by more than 30 MW. The shaft power increases because the feed

water stream entering the boiler is approximately 26 kg/s larger than that of the LFB­

AIR cycle.

Table 7-5: Power and efficiency balance of the cycle: LFB-AIR without CO2 capture.

Lower heating value (LHV) MJ/kg 9.01 Share of LHV Fuel mass flow rate kg/s 207.70

Fuel LHV MW 1871.38 100.0%

High pressure steam turbine MW 284.37 15.2%

Intermediate pressure steam turbine MW 445.05 23.8%

Low pressure steam turbine MW 276.50 14.8%

Turbomachinery shaft power MW 1005.93 53.8%

Turbomachinery generator loss MW 12.071 0.6%

Turbomachinery generator terminal output MW 993.85 53.1%

Plant auxiliary power (8% of generator output) MW ­79.51 ­4.2%

Net plant power island output MW 914.35 48.9%

Net plant power output MW 914.35 48.9%

Specific CO2 emissions g/kWhe 810

7 Simula t ion of the Analysed Power Genera t ion Cycles 129

Table 7-6: Power and efficiency balance of the cycle: LFB-OXY with CO2 capture.

Lower heating value (LHV) MJ/kg 9.01 Share of LHV Fuel mass flow rate kg/s 207.70

Fuel LHV MW 1871.38 100.0%

High pressure steam turbine MW 294.75 15.8%

Intermediate pressure steam turbine MW 459.15 24.5%

Low pressure steam turbine MW 282.30 15.1%

Turbomachinery shaft power MW 1036.20 55.4%

Turbomachinery generator loss MW 12.434 0.7%

Turbomachinery generator terminal output MW 1023.76 54.7%

Plant auxiliary power (8% of generator output) MW ­81.90 ­4.4%

Net plant power island output MW 941.86 50.3%

Energy expenditure for ASU MW ­116.82 ­6.2%

Work for CO2 compression MW ­92.960 ­5.0%

Net plant power output MW 732.08 39.1%

Specific CO2 emissions g/kWhe 98

Both the generator losses and the plant auxiliary power are in the same order of

magnitude as for the reference cycle, so that the net power output of the power island

is 27.5 MW higher than that of the LFB­AIR cycle. The significant difference between

the two configurations is caused by the expenditure of energy related to CO2 capture.

The contributors are the power required for the cryogenic ASU and for the

compression of the CO2. In total both contributors reduce the power output by more

than 11%. The largest contributor is the expenditure of energy for the cryogenic ASU,

which is 6.2% of the total heat input. The power required for the CO2 compression unit

is slightly lower but still consumes 5% of the total heat input. The net power output of

the LFB­OXY cycle is 732.08 MW, which is around 210 MW lower than the power

output of the power island. The net efficiency of this configuration is 39.1%. The

specific CO2 emissions are lowered by 88% to 98 g(CO2)/kWhe because 90% of the

CO2 from the exhaust stream are captured.

It should be emphasised that the net efficiency could be increased by 1.0 and 1.5%

if the integration of low temperature heat was integrated into the overall model of the

power generation process. Before the exhaust gas is fed to the CO2 compression,

further flue gas cooling (FGC) takes place down to 25°C. The integration of heat from

the FGC is not included in the model of the LFB­OXY cycle. Furthermore, the

separated CO2 is compressed in several stages. Between each stage intercooling of the

CO2 is performed to reduce the expenditure of energy of the compression unit. Due to the

130 7 .2 Simula t ion of the Oxyfuel Boi ler Cycles

Table 7-7: Power and efficiency balance of the cycle: LFB-OTM with CO2 capture.

Data of the steam cycle

Lower heating value (LHV) MJ/kg 9.01

Fuel mass flow rate kg/s 207.70

Fuel LHV MW 1871.38

High pressure steam turbine MW 277.69

Intermediate pressure steam turbine MW 557.71

Low pressure steam turbine MW 328.54

Turbomachinery shaft power MW 1163.94

Turbomachinery generator loss MW 13.967

Data of the gas turbine cycle

Lower heating value (LHV) MJ/kg 47.45

Fuel mass flow rate kg/s 17.50

Fuel LHV MW 830.48

Gas turbine expander MW 873.22

Gas turbine compressor MW ­666.32

Turbomachinery shaft power MW 206.90

Turbomachinery generator loss MW 2.483 Share of LHV

Overall cycle

Total heat input MW 2701.86 100.00%

Turbomachinery generator terminal output MW 1354.40 50.13%

Plant auxiliary power (8% of ref. output) MW ­80.00 ­2.96%

Net plant power island output MW 1274.40 47.17%

Work for CO2 compression MW ­91.949 ­3.40%

Net plant power output MW 1182.45 43.76%

Specific CO2 emissions g/kWhe 263

fact the CO2 compression is not modelled in detail but rather a specific expenditure of

energy is assumed to calculate the overall power, the heat from the CO2 compression

unit is not included in the modelling. If these two items were additionally integrated

into the modelling, the net efficiency would range between 40.0 and 40.5% for the

LFB­OXY cycle.

The last configuration investigated is the LFB­OTM cycle. The results of the LFB­

OTM cycle are given in table 7­7. The balance of power output and efficiency of the

LFB­OTM cycle is split into two sections. First the boiler is shown; it can be seen that

the heat input supplied by the lignite is the same as that of the other two

7 Simula t ion of the Analysed Power Genera t ion Cycles 131

configurations. The shaft power of the steam turbines increases by almost 128 MW

compared to the LFB­OXY cycle. The increase of the power output of the steam

turbines is caused by the large amount of additional heat: the exhaust energy of the

integrated gas turbine and the heat of the permeate stream of the OTM provide such an

amount of heat that feed water heaters are redundant. The feed water is solely

preheated by the exhaust gas of the gas turbine and by the permeate stream leaving the

OTM reactor, see figure 4­9. Therefore extraction of steam from the IP and LP steam

turbines is not required and the overall mass flow of both steam turbines increases

significantly.

In addition to the power output of the steam turbines, the gas turbine integrated into

the LFB­OTM cycle generates about 206 MW. The compressor inlet mass flow of the

gas turbine is 1500 kg/s. This shows that in a real cycle at least two – even three –

heavy duty gas turbines would be required to cope with such a large mass flow. The

large mass flow is a result of the necessary amount of oxygen for the boiler. The mass

flow of oxygen is around 165 kg/s, which is separated in the OTM reactor from the

preheated air mass flow. The large exhaust mass flow of the gas turbine reveals that

the exhaust energy of the gas turbine is capable of preheating most of the boiler feed

water.

Balancing the LFB­OTM cycle both heat inputs combine to around 2700 MW. The

total shaft power minus the plant auxiliary power results in a power output of the

overall power island of 1274 MW. The only parasitic power is the energy for the CO2

compression unit, at 92 MW. The net power output is finally 1182.45 MW, which

corresponds to a net efficiency of 43.76%. The specific CO2 emissions of the LFB­

OTM cycle are 263 g(CO2)/kWhe. The specific CO2 emissions are considerably higher

than that of the LFB­OXY cycle because the combustion process of the gas turbines

generates CO2, which is directly released to the atmosphere. Although 90% of the CO2

from the boiler exhaust gas is captured, in total only around 68% of the generated CO2

is captured.

7.2.2 Variation of the Operating Conditions of the OTM Reactor

The same variation of the operating conditions of the OTM reactor as for the IGCC

configuration is carried out for the LFB­OTM cycle. A parametric study on the

pressure and the mass flow rate of the sweep stream has been performed. The results

of the parametric study and the inducement on the net efficiency is illustrated in figure

7­2. For reason of comparison, the net efficiency of the LFB­OXY is additionally shown

132 7 .2 Simula t ion of the Oxyfuel Boi ler Cycles

Figure 7-2: Net efficiency of the LFB-OTM cycle vs. sweep stream mass flow rate for weep

stream pressure 1.5, 5, 10, and 15 bar; with a constant feed temperature of

900°C.

in figure 7­2. It can directly be seen that the high sweep pressures (and high sweep

mass flow rates) result in a net efficiency lower than that of the LFB­OXY cycle. For

the high sweep pressures (1.5 and 5.0 bar) the mass flow rate varies from 6 to 12%,

whereas for high sweep pressures (10 and 15 bar) the mass flow rates ranges from 16

to nearly 21%. For the high sweep pressures the net efficiency is between 40 and 36%.

Only in case of low sweep pressures the net efficiency is higher than that of the LFB­

OXY cycle and ranges from around 43 to 46%. From these results similar conclusions

as for the IGCC cycles can be drawn. The membrane based LFB­OTM cycle looks

only thermodynamically promising – in comparison to the ‘conventional’ CO2 capture

cycle – if the OTM reactor is capable to withstand challenging operating conditions.

For the same mass flow rate of the sweep stream, a different sweep pressure leads

to a change in net efficiency of nearly 1.5% points. The reason for that is that a higher

sweep pressure means a larger loss in power of the steam turbines because the steam

mass flow can not be expanded. The impact of the sweep pressure is larger than for the

IGCC cycle because the oxygen required for the boiler does not need to be pressurised.

In the case of the IGCC cycle, a lower sweep pressure needs to be compensated by

additional compression up to the gasifier pressure, but in case of the LFB­OTM cycle,

7 Simula t ion of the Analysed Power Genera t ion Cycles 133

the oxygen is used at atmospheric pressure. Therefore a lower sweep pressure has a

stronger positive effect on the net efficiency of the overall power generation cycle.

7.2.3 Variation of the Heat Input to Gas Turbine and Boiler

As previously described, the net efficiency of the LFB­OTM cycle may be

significantly higher than that of the LFB­OXY cycle. Even in comparison to the

reference cycle, LFB­AIR, the reduction in net efficiency is only around 3­6%­pts,

which is thermodynamically very promising. But assessing this result it has to be

recognised that the LFB­OTM cycle is not really a lignite fired boiler cycle but rather

a combination of a lignite fired boiler and an combined cycle power plant. With

respect to the heat input into the cycle around one third of the total heat is provided to

the gas turbine by the natural gas. Due to this large fraction it is not reasonable to

compare the net efficiency to a conventional lignite fired boiler cycle.

In figure 7­3 the net efficiency is plotted on the ordinate on the left hand side

versus the ratio of heat input to the gas turbine and the total heat input. The dotted line

represents the net efficiency of a hypothetical reference cycle. This net efficiency is

created by assuming that for a lignite fired boiler cycle (heat input to the gas turbine is

zero) the net efficiency corresponds to that of the LFB­AIR (48.9%). For a combined

cycle (heat input to the gas turbine is 100%) the net efficiency is set to 56.6%. These

two points (0 and 100%) are connected with a straight line. This is the dotted line “net

efficiency of the reference cycle” in figure 7­3. The different heat loads to the gas

turbines mean that the fuel mass flow changes, and thus the compressor inlet mass

flow varies. The compressor inlet mass flow is varied from 1300 to 1900 kg/s for the

LFB­OTM cycle. For the calculations in figure 7­3 the sweep pressure is constantly

1.5 bar, whereas the mass flow rate is between 6 and 7%. The net efficiency increases

with higher heat input to the gas turbine and ranges from 45 to 46%. The net efficiency

of the reference cycle varies from 51 to 52%. This shows that the reduction in net

efficiency is only 5 to 6%­pts. This leads to the conclusion that even compared with the

reference cycle with a higher efficiency, the LFB­OTM cycle is thermodynamically a

promising configuration. The LFB­OTM cycle shows a smaller reduction in net

efficiency than other power generation cycles with CO2 capture.

Due to the fact that the CO2 produced by the combustion process in the gas turbine

is not captured at all, the overall CO2 capture rate decreases with higher heat input to

the gas turbine. The effect of higher specific CO2 emissions is expressed in form of the

CO2 capture rate in figure 7­3 – on the ordinate on the right hand side. It can be seen that

134 7 .2 Simula t ion of the Oxyfuel Boi ler Cycles

Figure 7-3: Net efficiency (left ordinate) and CO2 capture rate (right ordinate) vs. ratio of

gas turbine heat load to total heat load of the LFB-OTM cycle.

the CO2 capture rate ranges between 69 and 75%. In comparison with other cycles

with CO2 capture such a capture rate is relatively low but this is counterbalanced by

promising results in net efficiency of the overall power generation cycle. Therefore,

the LFB­OTM cycle represents a reasonable compromise between efficient use of the

economically attractive primary energy resource (lignite) and low CO2 emissions.

8 Conclusions and Recommendations for Fur ther Work 135

8 Conclusions and Recommendations for Further Work

8.1 Summary of the Main Results

In this work two different types of coal fired power generation cycles are

thermodynamically analysed and benchmarked against each other. The following

different configurations of power generation cycles are analysed:

IGCC configurations LFB configurations

IGCC­REF LFB­AIR

IGCC­CAP LFB­OXY

IGCC­OTM LFB­OTM

IGCC­H2

The results of the net efficiency of all analysed configurations are summarised in

figure 8­1. It can be seen that if CO2 capture is applied to a power generation cycle –

without integrated membrane reactor – the net efficiency drops by approximately 10%

points. If a membrane reactor is integrated into a power generation cycle with CO2

capture, the net efficiency increases to some extent. For the IGCC cycles the net

efficiency increases by 0.4 to 1.6% points depending which type of reactor is integrated

Figure 8-1: Comparison of net efficiencies of the analysed power generation cycles.

136 8 .2 Conclusions on the IGCC configura t ions

into the cycle. For the LFB configurations, the net efficiency increases due to

integration of the OTM reactor by 5.3%­pts but it should be kept in mind that the

lignite fired boiler is ‘converted’ to a combination of a coal boiler fired and a natural

gas combined cycle power plant.

8.2 Conclusions on the IGCC configurations

8.2.1 General Conclusions on the Cycle Layout

In general, the complexity of the IGCC increases if CO2 capture is applied to the

cycle. The syngas treatment is extended by an additional reactor to carry out the CO­

shift reaction and the CO2 capture, where the separation of CO2 takes place by means

of physical adsorption. If an OTM reactor is integrated into an IGCC with CO2

capture, the complexity is even further increased. The OTM reactor is part of the gas

turbine cycle. In this arrangement the gas turbine has to have two combustion

chambers because the first combustion chamber preheats the air before it enters the

OTM reactor. The second combustor is required to heat the retentate stream to a

common temperature to achieve the desired temperature at the exhaust of the GT. The

heat of the permeate stream – mixture of steam and oxygen – leaving the OTM reactor

needs to be efficiently utilised in the water/steam cycle. Therefore additional heat

exchangers are necessary to transfer the heat to the water/steam flow.

8.2.2 Conclusions on the OTM Reactor as part of the IGCC

In comparison to the IGCC with CO2 capture but without an integrated OTM

reactor, IGCC­CAP, the thermodynamic potential of the IGCC­OTM cycle to increase

the net efficiency is limited. For challenging operating conditions (a sweep pressure of

1.5 bar, resulting in a large difference in total pressure across the membrane), the net

efficiency might be increased by nearly 1% point. Assuming more moderate operating

conditions for the OTM reactor (a sweep pressure of 5 bar) and a small mass flow rate

of the sweep stream, the net efficiency may be slightly higher than that of the IGCC­

CAP cycle. In this case the net efficiency yields around 35% ± 0.5% point. For more

conservative operating conditions (sweep pressure of 10 or 15 bar), the net efficiency

is significantly lower with 32% ± 1% point. Therefore one of the main conclusion is

that the inducement on the overall power generation cycle is determined by the

operating conditions of the OTM reactor. The IGCC­OTM is thermodynamically only

promising if a low pressure and small mass flow rate of the sweep stream is feasible. If

8 Conclusions and Recommendations for Fur ther Work 137

the OTM reactor is not capable of coping with a large difference in total pressure

across the membrane, the cycle becomes less attractive compared to the IGCC­CAP

cycle.

In addition to the impact on the overall power generation cycle the operating

conditions of the OTM reactor determine the size of the membrane reactor. Parametric

studies on the OTM reactor are carried out to show how the operating conditions

influence the membrane surface area. For the range of pressure and mass flow rate of

sweep stream investigated, the membrane surface area differs from 100,000 to

700,000 m2. One outcome is that the impact of the overall heat transfer coefficient on

the membrane surface area is small. The impact increases for lower feed temperature

of the membrane reactor.

For the investigated range in pressure and mass flow rate of the sweep stream, the

membrane surface area varies by a factor of 3­5. The difference in oxygen partial

pressure across the membrane is the driving force for the mass transport through the

membrane. Both parameters, mass flow rate and pressure of the sweep stream

determine the oxygen partial pressure on both sides of the OTM reactor. On the one

hand, it can be concluded that the pressure of the sweep stream has a stronger impact

on the membrane surface area than the mass flow rate of the sweep stream. If the goal

is a reduction of the membrane surface area as much as possible, then the pressure of

the sweep stream has to be as low as technically feasible. On the other hand, it should

be kept in mind that the requirement of a low sweep pressure leads to an increased

mechanical load for the membrane material. Therefore the OTM reactor needs to be

designed mechanically as a pressure vessel that is capable of withstanding the

difference in total pressure.

For a certain sweep stream mass flow rate the impact of the sweep pressure is small

on the net efficiency of the IGCC because the positive effect of a low sweep pressure

(low extraction pressure from the steam turbine) is partly used up because the

separated oxygen needs to be compressed to the gasifier pressure of 35 bar.

8.2.3 Conclusions on the hydrogen-selective Reactor as part of the IGCC

The IGCC with integrated hydrogen­selective membrane reactor achieves a higher

net efficiency than the IGCC­OTM cycle. The net efficiency of the IGCC­H2 cycle

yields 36.7%, which is 1.2% points higher than IGCC­OTM cycle. The net efficiency

looks thermodynamically promising in particular in comparison with the conventional

IGCC with CO2 capture, IGCC­CAP.

138 8 .3 Conclusions for LFB Cycles

The cycle layout of the IGCC­H2 results in a large number of additional heat

exchangers, which are required to transfer heat internally to and from the syngas.

Before the syngas stream enters the hydrogen­selective membrane reactor as a feed

stream it needs to be heated. The heat is provided by internal heat exchange with the

retentate stream leaving the reactor. Due to the fact that the retentate stream contains

some combustible components (hydrogen and carbon monoxide), an supplementary

burner is necessary to combust these components. In particular the heat generated by

the combustion in the supplementary burner leads to the necessity of several heat

exchangers to utilise the heat efficiently. Furthermore, an operational aspect is that the

mixture of hydrogen and nitrogen leaving the reactor needs to be compressed to the

combustor pressure level of the gas turbine. The mixture leaves the reactor with a

temperature close to 600°C. A compressor capable of handling such a hot gas does not

exit currently. Either materials need to be developed which are able to withstand

tenmperatures of more than 600°C or compressor parts need to be cooled.

8.3 Conclusions for LFB Cycles

8.3.1 General Conclusions on the Cycle Layout

The integration of the OTM reactor leads to significant changes for the LFB cycle.

The major change is that the OTM reactor is integrated into a gas turbine cycle.

Consequently, the lignite fired boiler is ‘converted’ to a combination of a coal fired

boiler and a natural gas combined cycle power plant. A practical issue when building

such a power plant is that, beside the lignite, natural gas has to be supplied to the

location of the power plant.

The arrangement of the OTM reactor is slightly different from the IGCC­OTM

cycle because the gas turbine has only one combustion chamber. The retentate stream

leaving the OTM reactor is directly expanded in the turbine. Because of the large

compressor inlet mass flow assumed for the LFB­OTM cycle, if such a power plant

were built, two OTM reactors and two gas turbines would need to be operated in

parallel to cope with the large mass flow.

The layout of the LFB­OTM cycle changes to a large extent because of the large

amount of exhaust energy from the gas turbine and the heat transfer from the gas

turbine to the water/steam cycle by the OTM reactor. The large amount of additional

heat results in a substitution of all ten feed water heater. Between condenser and boiler

the feed water is preheated solely by the exhaust gas of the turbine, the water/steam

8 Conclusions and Recommendations for Fur ther Work 139

mixture leaving the OTM reactor and the exhaust gas of the boiler. Therefore the

power output increases significantly because no steam extraction for preheating of the

feed water takes place.

8.3.2 Conclusions on the OTM Reactor as part of the LFB cycle

If the OTM reactor is integrated into the LFB cycle the net efficiency can be

increased by 4­5% points in comparison to the cycle with cryogenic ASU – LFB­

OXY. Even compared to a hypothetical reference cycle, the net efficiency reduces

only by around 6% points. Thermodynamically this is the most promising cycle of all

investigated configurations. It should be emphasised that the CO2 capture rate is

lowered because the CO2 from gas turbine is not captured at all. The CO2 capture rate

of the LFB­OTM cycle is approximately 70%.

The parametric studies of the OTM reactor are carried out for the conditions of the

IGCC configuration. Of course, the implications of pressure and mass flow rate on the

membrane surface area are identical. The size of the power generation in terms of

power output is determined by the assumed mass flow of lignite as fuel mass flow.

Due to a large reference cycle (LFB­AIR) of approximately 1000 MW, the power

output of the LFB­OTM is of the same order of magnitude (same mass flow of coal

used as for the reference cycle). Due to the large amount of oxygen required for the

boiler, the membrane surface area ranges from 500,000 to 1,000,000 m2. It is likely

that such a OTM reactor would be realised – even in large scale application – in

smaller scale if such a power plant were built. In case of a 1000 MW, more than one

OTM reactor would be operated in parallel configuration.

8.4 Recommendations for further work

As previously explained the operating conditions of the membrane reactor

determine the performance of the overall power generation cycle. Therefore the

feasibility of temperatures and pressure could be further assessed to investigate where

limitations for the membrane material occur. For both types of membrane reactor,

OTM and hydrogen­selective, a higher operating temperature will increase the

permeation flux through the membrane. Further investigations of material limits with

respect to the temperature could be done. The limitation of the operating temperature

could be investigated for different types of membrane material.

140 8 .4 Recommendat ions for fur ther work

The difference in total pressure across the membrane finally determines the

difference in partial pressure of the permeate and thus the driving force of the mass

transport through the membrane. The higher the difference in total pressure across the

membrane, the higher the permeation flux through the membrane. At the same time a

large difference in total pressure leads to a high mechanical load for the membrane

material. It should be investigated which pressure differences can be realised for the

different types of membrane materials. For this purpose the mechanical integrity of the

membrane reactors needs to be investigated. For instance practical issues such as

sealing of the membrane could be addressed in detail.

Due to the large variety of perovskites and perovskite­related materials a

benchmarking of different membrane material could be carried out to find the highest

achievable permeation fluxes, which would decrease the overall size of the membrane

reactor. A benchmarking of different membrane materials could also include economic

aspects such as the pure material costs or the total manufacturing costs of the

membrane reactor. The costs could be used as a comparative parameter especially due

to the large required reactor sizes. In an economic context the question could be raised

of how large the permeation flux through the membrane needs to be to fulfil a given

criteria for the total investment costs of such a reactor. Regarding maximum

achievable permeation fluxes the impact of species – CO2 for the oxygen transport

membranes and SO2 for the hydrogen­selective membrane reactor – on the

performance of the membrane could be investigated.

From an operational point of view aspects like operating a power generation cycle

with integrated membrane reactor in partload, start­up and shut­down strategies could

be looked at. If such power generation cycles were needed to be operated in a kind of

cyclic operation, the availability of the membrane reactor and the required sweep

streams (steam and nitrogen, respectively) would need to be addressed; for example

start­up and shutdown times of the membrane reactor have not been investigated yet. It

needs to be addressed to what extent such power generation cycles are able to cover

the conventional operational range of a power plant.

Thermo­economic assessment is not within the scope of this work but in a future

work the investment costs of such membrane reactors need to be considered for a

thermo­economic assessment. Such an assessment should be based for example on

existing publications like [193]. The large size of the proposed membrane reactors will

significantly increase the total investment costs of the overall power generation cycle.

Since the cost of electricity depends on the total investment cost, it would be

8 Conclusions and Recommendations for Fur ther Work 141

worthwhile to investigate, how the cost of electricity changes if such membrane

reactors are integrated in these coal fired power generation cycles.

Appendix 143

Appendix

A.1 Gas compositions of the different IGCC configurations

In this section the molar compositions after different process steps are given for the

investigated IGCC configurations. Table A1­1 contains the gas compositions for the

IGCC­REF process after the syngas cooler, COS­hydrolysis, desulphurisation,

humidification and finally of the syngas which is fed to the combustion chamber of the

gas turbine. Since Table A1­2 shows the gas composition at various position in the

IGCC­CAP process, no humidification takes place but therefore the composition after

the CO­shift reaction and after separation of the CO2 are listed.

Table A1-1: Molar composition of the raw gas, syngas and the fuel gas of the IGCC-REF

process at different positions in the process.

Component Syngas cooler

COS­Hydrolys.

Desulp­hurisation

Humidi­fication

Combust. Fuel gas [mol­%]

H2 23.00 23.40 25.67 21.15 12.72

CO 55.00 56.62 62.13 51.17 30.79

CO2 5.00 2.79 2.74 2.26 1.36

N2 9.00 7.40 8.12 6.69 43.86

H2O 8.00 8.71 0.35 17.91 10.77

H2S 0.00 0.17 0.00 0.00 0.00

CH4 0.00 0.04 0.04 0.04 0.02

Ar 0.00 0.87 0.95 0.79 0.47

Table A1-2: Molar composition of the raw gas, syngas and the fuel gas of the IGCC-CAP

process at different positions in the process.

Component Syngas cooler

COS­Hydrolys.

Desulp­hurisation

CO­shift CO2 removal

Combust. Fuel gas [mol­%]

H2 23.00 23.25 25.67 37.74 85.03 44.64

CO 55.00 56.23 62.09 1.94 4.37 2.29

CO2 5.00 2.79 2.76 27.38 0.93 0.49

N2 9.00 7.36 8.13 3.67 8.27 51.85

H2O 8.00 9.29 0.35 28.82 0.38 0.20

H2S 0.00 0.17 0.00 0.00 0.00 0.00

CH4 0.00 0.04 0.04 0.02 0.05 0.02

Ar 0.00 0.87 0.96 0.43 0.97 0.51

144 Appendix

In Table A1­3 the gas compositions at various positions of the IGCC­OTM, which

are similar to the IGCC­CAP process, are given. Finally, the gas compositions in the

IGCC­H2 process are listed in Table A1­4. In the IGCC­H2 process the separation of

carbon dioxide is utilised by the hydrogen­selective membrane reactor. Therefore, the

composition of the syngas leaving the membrane reactor is shown in Table A1­4.

Table A1-3: Molar composition of the raw gas, syngas and the fuel gas of the IGCC-H2

process at different positions in the process.

Component Syngas cooler

COS­Hydrolys.

Desulp­hurisation

CO­shift CO2 removal

Combust. Fuel gas [mol­%]

H2 23.00 23.25 25.67 37.74 85.03 44.64

CO 55.00 56.23 62.09 1.94 4.37 2.29

CO2 5.00 2.79 2.76 27.38 0.93 0.49

N2 9.00 7.36 8.13 3.67 8.27 51.85

H2O 8.00 9.29 0.35 28.82 0.38 0.20

H2S 0.00 0.17 0.00 0.00 0.00 0.00

CH4 0.00 0.04 0.04 0.02 0.05 0.02

Ar 0.00 0.87 0.96 0.43 0.97 0.51

Table A1-4: Molar composition of the raw gas, syngas and the fuel gas of the IGCC-H2

process at different positions in the process.

Component Syngas cooler

COS­Hydrolys.

Desulp­hurisation

CO­shift MR (Perm.)

MR (Ret.)

Combust. Fuel gas [mol­%]

H2 23.00 23.25 25.67 37.74 72.60 0.60 44.64

CO 55.00 56.23 62.09 1.94 0.00 3.10 0.00

CO2 5.00 2.79 2.76 27.38 0.00 43.71 0.00

N2 9.00 7.36 8.13 3.67 27.40 5.87 55.36

H2O 8.00 9.29 0.35 28.82 0.00 46.00 0.00

H2S 0.00 0.17 0.00 0.00 0.00 0.00 0.00

CH4 0.00 0.04 0.04 0.02 0.00 0.03 0.00

Ar 0.00 0.87 0.96 0.43 0.00 0.69 0.00

Appendix 145

A.2 Properties of certain positions of the gas turbine process of the

IGCC configurations

Key parameters of the combined cycle of the different IGCC configurations are

listed in Table A2­1 and Table A2­2.

Table A2-1: Overview of key parameters of the cycles: IGCC-REF, IGCC-REF-ASU and

IGCC-CAP.

Parameter IGCC­REF IGCC­REF­ASU

IGCC­CAP

Compressor air inlet mass flow / kg/s 674.2 826.1 636.4

Pressure ratio of the compressor / ­ 17 17 17

Combustor air inlet mass flow / kg/s 493.3 493.3 472.0

Fuel mass flow / kg/s 174.0 174.0 111.4

LHV of the fuel gas / MJ/kg 4.939 4.939 6.933

Hot gas temperature / °C 1424.9 1424.9 1425.1

Turbine inlet temperature (TIT) / °C 1229.1 1229.1 1229.5

Turbine exhaust temperature (TAT) / °C 594.7 594.7 579.2

Turbine exhaust mass flow / kg/s 848.2 848.2 747.8

Stack temperature / °C 86.3 83.0 88.3

Table A2-2: Overview of key parameters of the cycles: IGCC-CAP-ASU, IGCC-OTM and

IGCC-H2.

Parameter IGCC­CAP­ASU

IGCC­OTM IGCC­H2

Compressor inlet mass flow / kg/s 788.3 616.3 624.6

Pressure ratio of the compressor / ­ 17 17 17

Combustor air inlet mass flow / kg/s 472.0 472.3 463.2

Fuel mass flow / kg/s 111.4 111.4 109.6

LHV of the fuel gas / MJ/kg 6.933 6.933 6.576

Hot gas temperature / °C 1425.0 1425.0 1425.0

Turbine inlet temperature (TIT) / °C 1229.4 1231.8 1229.4

Turbine exhaust temperature (TAT) / °C 579.1 586.3 577.2

Turbine exhaust mass flow / kg/s 747.8 692.8 734.2

Stack temperature / °C 74.7 149.8 100.0

146 Appendix

A.3 Reference cases from ENCAP SP 3 of the lignite fired boiler

process with and without CO2 capture

The power and efficiency balances of the lignite fired boiler processes are shown in

Table A3­1 and Table A3­2.

Table A3-1: Power and efficiencies balance of the reference case from SP 3 (ENCAP) of

the lignite fired boiler process without CO2 capture.

Lower heating value (LHV) MJ/kg 9.01 Share of LHV Fuel mass flow rate kg/s 207.70

Fuel LHV MW 1871.38 100.0%

High pressure steam turbine MW 288.86 15.4%

Intermediate pressure steam turbine MW 447.59 23.9%

Low pressure steam turbine MW 275.70 14.7%

Turbomachinery shaft power MW 1012.15 54.1%

Turbomachinery generator loss MW 12.146 0.6%

Turbomachinery generator terminal output MW 1000.00 53.4%

Own consumption (8% of generator output) MW ­80.00 ­4.3%

Net plant power island output MW 920.00 49.2%

Net plant power output MW 920.00 49.2%

Table A3-2: Power and efficiencies balance of the reference case from SP 3 (ENCAP) of

the lignite fired oxyfuel boiler process with CO2 capture.

Lower heating value (LHV) MJ/kg 9.01 Share of LHV Fuel mass flow rate kg/s 207.70

Fuel LHV MW 1871.38 100.0%

High pressure steam turbine MW 302.04 16.1%

Intermediate pressure steam turbine MW 469.83 25.1%

Low pressure steam turbine MW 288.72 15.4%

Turbomachinery shaft power MW 1060.59 56.7%

Turbomachinery generator loss MW 12.727 0.7%

Turbomachinery generator terminal output MW 1047.86 56.0%

Own consumption (8% of generator output) MW ­83.83 ­4.5%

Net plant power island output MW 964.03 51.5%

Energy expenditure for ASU MW ­117.52 ­6.3%

Work for CO2 compression MW ­94.110 ­5.0%

Net plant power output MW 752.40 40.2%

Appendix 147

A.4 Emitted and captured carbon dioxide of the investigated power

generation processes

Details of the carbon dioxide which is captured and emitted, respectively, of the

various IGCC processes are given in Table A4­1 and Table A4­2.

Table A4-1: Information about the emitted and captured CO2 of the IGCC-REF, IGCC-

REF-ASU and IGCC-CAP cycles.

Parameter IGCC­REF IGCC­REF­ASU

IGCC­CAP

Exhaust mass flow to the stack / kg/s 848.2 848.2 747.8

Molecular weight of the exhaust / g/mol 29.19 29.19 27.48

Mole fraction of CO2 in the exhaust / mol­% 8.09 8.09 0.72

Amount of emitted CO2 / kg/s 103.09 103.01 8.31

Amount of captured CO2 / kg/s 0.0 0.0 94.68

Net power output / MW 490.34 495.22 381.35

Specific CO2 emissions / g/kWh(e) 757 749 78

CO2 capture efficiency / % 0.00 0.00 98.07

Table A4-2: Information about the emitted and captured CO2 of the IGCC-CAP-ASU,

IGCC-OTM and IGCC-H2 processes.

Parameter IGCC­CAP­ASU

IGCC­OTM IGCC­H2

Exhaust mass flow to the stack / kg/s 747.8 692.8 734.2

Molecular weight of the exhaust / g/mol 27.48 27.24 27.36

Mole fraction of CO2 in the exhaust / mol­% 0.72 0.77 0.03

Amount of emitted CO2 / kg/s 8.23 8.31 0.00

Amount of captured CO2 / kg/s 94.68 94.68 94.68

Net power output / MW 389.43 386.17 398.64

Specific CO2 emissions / g/kWh(e) 76 77 0

CO2 capture efficiency / % 98.09 98.10 98.16

Details of the carbon dioxide which is captured and emitted, respectively, of the

various oxyfuel processes are given in Table A4­3.

148 Appendix

Table A4-3: Information about the emitted and captured CO2 of the LFB-AIR, LFB-OXY

and the LFB-OTM cycles.

Parameter LFB­AIR LFB­OXY LFB­OTM

Exhaust mass flow to the stack / kg/s 945.38 274.4 254.76

Molecular weight of the exhaust / g/mol 29.63 34.57 42.24

Mole fraction of CO2 in the exhaust / mol­% 14.68 58.26 88.52

Amount of emitted CO2 / kg/s 205.75 19.99 86.28

Amount of captured CO2 / kg/s 0.00 183.15 185.08

Net power output / MW 914.35 732.08 1182.45

Specific CO2 emissions / g/kWh(e) 810 98 263

CO2 capture efficiency / % 0.00 87.83 69.46

It should be emphasised that for the LFB­OTM cycle the ’exhaust mass flow to the

stack’ is the exhaust mass flow of the boiler. The exhaust mass flow of the gas turbine

is 1352 kg/s containing 2.3 mol­% of CO2. The mass flow of the gas turbine

containing 49.1 kg/s of CO2 is not treated. This is the reason for relative high specific

CO2 emissions of 263 g(CO2)/kWh(e) although 90% of the CO2 from the boiler

exhaust mass flow is captured.

Appendix 149

A.5 Variation of the feed temperature of the OTM reactor

Variation of the feed temperature

The temperature of the feed temperature for the OTM reactor is assumed be to

constant in both configurations the IGCC­OTM cycle and the lignite fired oxyfuel

boiler cycle. A variation of the feed temperature is conducted to better understand the

impact of this parameter on the required membrane surface area. In all simulations a

feed temperature of 900°C is assumed. Therefore the feed temperature is varied from

800 to 1000°C. Of course, a lower feed temperature would lead to a larger required

membrane surface area because the operating temperature of the reactor would

decrease. Therefore the focus for the variation of the feed temperature is more to a

higher feed temperature to answer the question whether or not an increased feed

temperature would have a strong impact on the overall membrane surface.

The variation of the feed temperature is done for four different sweep pressures: 2,

5, 10 and 15 bar, see Figure A5­1 to Figure A5­4. At the same time the sweep stream

mass flow rate is varied from 2 to 24%. For all four different sweep pressures the

membrane surface area decreases by around 40%, if the feed temperature is increased

from 900 to 1000°C. But, of course, the absolute value of the membrane surface area

changes significantly with the sweep pressure. Comparing a sweep mass flow rate of

16%, at a sweep pressure of 2 bar the membrane surface area is reduced from 70,000

to 44,000 m2, whereas at a sweep pressure of 15 bar it decreases from 400,000 to

243,000 m2. The size of the required membrane surface reveals the necessary size of

the OTM reactor. In particular, if the sweep pressure have to be high for limiting the

mechanical load for the membrane material, a higher feed temperature would help to

reduce the overall size of the OTM reactor. Assuming the surface­volume­ratio of

750 m2/m3, an area of 400,000 m2 would result in a reactor volume of around 530 m3

(e.g. dimensions of 6 x 6 x 15 m3).

Comparing the variation of the feed temperature from 800 to 900°C and from 900

to 1000°C, respectively, Figure A5­1 to Figure A5­4 show that the relative change in

the membrane surface area increases for the lower feed temperature.

150 Appendix

775 800 825 850 875 900 925 950 975 1000 10250.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

Sweep stream mass flow rate

Mem

bra

ne s

urf

ace

are

a /

10

6 m

2

Feed temperature / °C

pSweep

= 2 bar

2 % 4 % 8 % 12 % 16 % 20 % 24 %

Figure A5-1: Membrane surface area vs. feed stream temperature for different sweep

stream mass flow rates at a sweep stream pressure of 2 bar.

775 800 825 850 875 900 925 950 975 1000 10250.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

Sweep stream mass flow rate

Me

mb

ran

e s

urf

ac

e a

rea

/

10

6 m

2

Feed temperature / °C

pSweep

= 5 bar

5 % 9 % 12 % 16 % 20 % 24 %

Figure A5-2: Membrane surface area vs. feed stream temperature for different sweep

stream mass flow rates at a sweep stream pressure of 5 bar.

Appendix 151

775 800 825 850 875 900 925 950 975 1000 10250.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

Sweep stream mass flow rate

Me

mb

ran

e s

urf

ace

are

a

/ 1

06 m

2

Feed temperature / °C

pSweep

= 10 bar

9 % 12 % 14 % 16 % 20 % 24 %

Figure A5-3: Membrane surface area vs. feed stream temperature for different sweep

stream mass flow rates at a sweep stream pressure of 10 bar.

775 800 825 850 875 900 925 950 975 1000 10250.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

Sweep stream mass flow rate

Mem

bra

ne s

urf

ac

e a

rea

/

10

6 m

2

Feed temperature / °C

pSweep

= 15 bar

14 % 15 % 16 % 20 % 24 %

Figure A5-4: Membrane surface area vs. feed stream temperature for different sweep

stream mass flow rates at a sweep stream pressure of 15 bar.

152 Appendix

Variation of both sweep stream mass flow and pressure for a feed

temperature of 800 and 1000°C

The combination of varying the mass flow as well as the pressure of the feed

stream is conducted for a feed temperature of 800 and 1000°C. For both feed

temperatures the sweep pressure is varied from 2 to 15 bar and the mass flow rate of

the sweep stream is changed from 2 to 24%. In both figures, Figure A5­5 and Figure

A5­6, the membrane surface area is plotted versus the sweep stream mass flow rate.

For all cases with a feed temperature of 1000°C the membrane surface area is

below 500,000 m2. In case of a feed temperature of 800°C the membrane surface area

increases to nearly 1,300,000 m2 (for a sweep pressure of 15 bar). Especially for high

sweep pressure (> 10 bar), a high feed temperature could help to limit the size of the

membrane reactor. On a different note, if the sweep pressure must not be below a

certain value (e.g. 10 bar) to restrict the mechanical load for the membrane material,

the feed temperature can positively impact the overall size of the membrane reactor.

In Figure A5­9, Figure A5­10 and Figure A5­11 the development of the oxygen

partial along the reactor length is illustrated. As previously described the OTM reactor

is divided into 75 discrete elements. The normalised reactor length represents these 75

elements. The numbering of the elements starts on that side where the feed stream

enters the reactor (reactor length “0”). Since the OTM reactor is considered as a

counterflow apparatus, the sweep stream enters the reactor at the opposite side (reactor

length “1”).

In all three charts the feed temperature is set to 900°C. In Figure A5­9 the sweep

pressure is constant with 2 bar; the sweep stream mass flow rates is varied from 2 to

24%. In Figure A5­10 the sweep pressure is assumed with 15 bar. Due to a higher

sweep pressure the sweep stream mass flow rates is changed from 14 to 24%. A

constant sweep stream mass flow rate is used for Figure A5­11, but therefore the

sweep stream pressure is varied from 2 to 15 bar.

Appendix 153

2 4 6 8 10 12 14 16 18 20 22 240.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

Sweep stream pressureTFeed

= 800°C

2 bar 5 bar 7.5 bar 10 bar 12.5 bar 15 bar

Mem

bra

ne s

urf

ace a

rea / 1

06 m

2

Sweep stream mass flow rate / %

Figure A5-5: Membrane surface area vs. sweep stream mass flow rate for different sweep

stream pressures and a feed temperature of 800°C.

2 4 6 8 10 12 14 16 18 20 22 240.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

TFeed

= 1000°CSweep stream pressure

2 bar 5 bar 7.5 bar 10 bar 12.5 bar 15 bar

Mem

bra

ne s

urf

ace a

rea / 1

06 m

2

Sweep stream mass flow rate / %

Figure A5-6: Membrane surface area vs. sweep stream mass flow rate for different sweep

stream pressures and a feed temperature of 1000°C.

154 Appendix

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 160.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

TFeed

= 800°CSweep stream mass flow rate 5 % 14 % 8 % 16 % 9 % 20 % 12 % 24 %

Mem

bra

ne s

urf

ace a

rea / 1

06 m

2

Sweep stream pressure / bar

Figure A5-7: Membrane surface area vs. sweep stream mass flow rate for different sweep

stream pressures and a feed temperature of 800°C.

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 160.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

Sweep stream mass flow rate 5 % 14 % 8 % 16 % 9 % 20 % 12 % 24 %

TFeed

= 1000°C

Mem

bra

ne s

urf

ace a

rea / 1

06 m

2

Sweep stream pressure / bar

Figure A5-8: Membrane surface area vs. sweep stream mass flow rate for different sweep

stream pressures and a feed temperature of 1000°C.

Appendix 155

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.00.0

0.5

1.0

1.5

2.0

2.5

3.0O

xyg

en

part

ial p

ressu

re / b

ar

Reactor length (normalised)

TFeed

= 900°C

pSweep

= 2 bar

Retentate Permeatem

Sweep = 2 %

mSweep

= 4 %

mSweep

= 8 %

mSweep

= 12 %

Retentate Permeatem

Sweep = 16 %

mSweep

= 20 %

mSweep

= 24 %

Figure A5-9: Development of oxygen partial pressure along the normalised length of the

membrane reactor (for a feed temperature of 900°C and a sweep stream

pressure of 2 bar).

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.00.0

0.5

1.0

1.5

2.0

2.5

3.0

Oxyg

en

part

ial p

ressu

re / b

ar

Reactor length (normalised)

TFeed

= 900°C

pSweep

= 15 bar

Retentate Permeatem

Sweep = 14 %

mSweep

= 16 %

mSweep

= 20 %

mSweep

= 24 %

Figure A5-10: Development of oxygen partial pressure along the normalised length of the

membrane reactor (for a feed temperature of 900°C and a sweep stream

pressure of 15 bar).

156 Appendix

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.00.0

0.5

1.0

1.5

2.0

2.5

3.0

Retentate Permeatep

Sweep = 10 bar

pSweep

= 15 bar

Retentate Permeatep

Sweep = 2 bar

pSweep

= 5 bar

Ox

yg

en

pa

rtia

l p

res

su

re

/ b

ar

Reactor length (normalised)

TFeed

= 900°C

mSweep

= 24 %

Figure A5-11: Development of oxygen partial pressure along the normalised length of the

membrane reactor (for a feed temperature of 900°C and a sweep stream

pressure of 2 bar).

Appendix 157

A.6 Power and Efficiency Balances for IGCC cycles with Integrated

ASU

The breakdown of power output and efficiency for the cycle IGCC­REF­ASU is

shown in Table A6­1. The main contributors to the overall power and efficiency

change due to the fact that the mass flow required for the cryogenic ASU is

compressed by the compressor of the gas turbine. Therefore, the power output of the

gas turbine decreases significantly because the mass flow of both turbomachineries do

not ‘match’ anymore. The inlet mass flow of the compressor is about 150 kg/s larger

than that of the IGCC­REF cycle. The sum of the changed power output of the

combined cycle and the saving on the ASU leads to an overall increase in net power

output of the power plant of around 7 MW. The larger power output corresponds to

an increase of 0.68%­pts.

Table A6-1: Power and efficiency balance for the cycle IGCC-REF-ASU.

LHV MJ/kg 25.174 Share of LHV Fuel flow kg/s 43.21

Fuel LHV MW 1087.77 100.00%

Gas turbine expander MW 667.60 61.37%

Steam turbines (HP, IP and LP) MW 208.67 19.18%

Gas turbine compressor MW ­333.12 ­30.62%

Turbomachinery shaft power MW 543.15 49.93%

Turbomachinery mechanical loss MW ­2.37 ­0.22%

Turbomachinery generator loss MW ­8.84 ­0.81%

Turbomachinery generator terminal output MW 531.94 48.90%

Pumps of the WSC cycle kW ­3615 ­0.33%

Power island gross power output MW 528.33 48.57%

Plant auxiliary power MW ­8.66 ­0.80%

Net plant power island output MW 519.67 47.77%

Work oxygen production and compression MW ­21.95 ­2.02%

Work for any other related auxiliary processes MW ­0.002 0.00%

Net plant power output MW 497.72 45.76%

Specific CO2 emissions g/kWhe 745

158 Appendix

The breakdown of power output and efficiency for the cycle IGCC­CAP­ASU is

presented in Table A6­2. The effects are the same as for the IGCC cycle with CO2

capture. Therefore a similar change in overall power output and efficiency is obtained

for the IGCC­CAP­ASU cycle. The sum of the changed power output of the combined

cycle and the saving on the ASU leads to an overall increase in net power output of the

power plant of approximately 8 MW. The larger power output corresponds to an

increase of 0.74%­pts.

Table A6-2: Power and efficiency balance for the cycle IGCC-CAP-ASU.

LHV MJ/kg 25.174 Share of LHV Fuel flow kg/s 43.21

Fuel LHV MW 1087.77 100.00%

Gas turbine expander MW 619.68 56.97%

Steam turbines (HP, IP and LP) MW 159.60 14.67%

Gas turbine compressor MW ­317.70 ­29.21%

Turbomachinery shaft power MW 461.58 42.43%

Turbomachinery mechanical loss MW ­2.033 ­0.19%

Turbomachinery generator loss MW ­7.594 ­0.70%

Turbomachinery generator terminal output MW 451.95 41.55%

Pumps of the WSC cycle kW ­3585 ­0.33%

Power island gross power output MW 448.36 41.22%

Plant auxiliary power MW ­7.428 ­0.68%

Net plant power island output MW 440.94 40.54%

Work CO2 compression MW ­24.48 ­2.25%

Work oxygen production and compression MW ­27.02 ­2.48%

Work for any other related auxiliary processes MW ­0.010 0.00%

Net plant power output MW 389.43 35.80%

Specific CO2 emissions g/kWhe 76

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Lebenslauf

Persönliche Daten

Geburtsdatum: 27. März 1974

Geburtsort: Wimbern / Wickede (Ruhr)

Staatsangehörigkeit: deutsch

Familienstand: verheiratet, 2 Kinder

Werdegang

1981 – 1985 Kath. Grundschule, Röhrschule

1985 – 1988 Franz-Stock-Gymnasium

1988 – 1991 Realschule Hüsten Abschluss: Fachoberschulreife (06/1991)

1991 – 1994 Ausbildung zum Technischen Zeichner bei Fa. Egon Hillebrand in 59755 Arnsberg

1994 – 1995 Fachoberschule Metalltechnik in 59759 Arnsberg Abschluss: Fachhochschulreife (07/1995)

1995 – 1996 Ableistung des Wehrdienstes, Flugabwehrraketengruppe 21, Graf-Yorck-Kaserne, Möhnesee-Echtrop

1996 – 2000 Studium des Maschinenbaus an der Universität-Gesamthochschule Paderborn, Abteilung Meschede Fachrichtung: Konstruktionstechnik Abschluss: Dipl.-Ing. (FH) (08/2000)

2000 – 2003 Studium des Maschinenbaus an der Universität Paderborn Fachrichtung: Produktentwicklung Abschluss: Dipl.-Ing. (05/2003)

2003 – 2007 Wissenschaftlicher Mitarbeiter am Lehrstuhl für Thermodynamik und Energietechnik, Universität Paderborn

2007 – 2008 Wissenschaftlicher Mitarbeiter am Lehrstuhl für Thermodynamik, Ruhr-Universität Bochum

ab 2008 Berechnungsingenieur für Gasturbinen bei Alstom (Schweiz) AG, Baden, Schweiz