9
ELSEVIER PII: SOO16-2361 96)00132-9 Fuel Vol. 75, No. 12, 1449-1456, 1996 p. Copyright 0 1996 Elsevier Science Ltd Printed in Great Britain. All rights reserved 0016-2361/96 $15.00+0.00 Thermodynamic simulations of lignite fired IGCC with n s t u desulfurization and CO2 capture Wim F. Elseviers Tania Van Mierlo Martine J. F. Van de Voorde and Harry Verelst Department of Chemical Engineering, Vrue Universiteit Brussel, Pleinlaan 2, B- 1050 Brussels, Belgium Received 18 April 1996; revised 73 June 1996) IGCC cycles running on Greek lignite with in situ desulfurization and CO 2 capture w ere simulated using the ASPEN PLUS process simulator and compared with other power cycles. Dolomite was chosen as the material for in situ desulfurization. Considering the power produced and residual sulfur content of the off- gas, an oxygen-blown gasifier operated at 1.8 MPa and 95O”C, having a thermal efficiency of 38.9%, gives the best performance. The eff iciency loss due to CO2 capture can be largely compensated using IGCC cycles. Copyright 0 1996 Elsevier Science Ltd Keyw ords : IGCC; CO* ca ptu re; s imulation) The application of oil- and gas-fired combustion turbine cycle systems has grown rapidly because of the lower cost, higher efficiencies and improved gas turbine reliability. Because of supply limitations and t he increasing costs of conventional clean fuels such as oil and gas, the use of coal in combined cycles for electrical and thermal energy production is gaining interest. Integrated gasification combined cycles (IGCC) with high efficiencies and low emissions, due to integrated gas cleaning, are an attractive alternative for power production from solid fuels’)2, as shown by several demonstration plants: Buggenum (The Netherlands)3, Reno (USA)4 and Puertollano (Spain)‘16. Modern coal gasification technologies present a unique opportunity to combine the advantages of high-efficiency combined- cycle power generation with an environme ntally friendly coal-based process. IGCC power stations are a major step forward in producin g electricity a t competitive prices while significantly reducing emissions through advances in low-NO, gas turbine combustion technol- ogy and intensive desulfurization processes. GASIFICATION BASICS For a better understanding of the subsequent discussion of the simulation results, a short overview of gasification basic reactions, as used in this model, is given. Solid fuels with high moisture and relatively low fixed carbon contents are not suited for direct combustion. As an alternative, they can be converted to a combustible gas containing mainly H2 and CO in a gasification process. The gasification gas can be used for firing a gas turbine or for the production of several basic chemicals such as NH3 or acetic acid from synthesis gas after steam reforming. The prime reactions taking place in solid fuel gasification can be summarized as follows. The heat required for the endothermic Boudouard reaction C(s) + CO2 g) = 2CO ,)> AH298 = 283.1 kJ mol-’ (1) and the reactions of the fixed carbon with steam C(s) + H2O(s) = CO(s) + H2 4> AH2s8 = 131.4kJmol-’ (2) Cc,) + 2H20 ,) = COz ,) + 2H2 g), AH29s = 90.28 kJmol_’ (3) is delivered through partial combustion: C(s) + @2 g) = CO ,), AH2s8 = -110.6kJmoll’ (4) C(s) + 02 g) = C02 g)r AH29s = -393.7 kJmol_’ (5) If pure oxygen from a cryogenic air separation plant is used as oxidizing agent, the gas produced consists mainly of CO, C02, H2 and H20 . N2 will be present if air is used instead of oxygen. The gas composition is determined through the equilibrium of the water gas shift reaction: CO(s) + H20 ,) = C02cg) + H2 g)> AH29s = -41.10 kJmol_’ (6) The above-mentioned reactions are just a small selection of possible reactions, as upon pyrolysis a number of volatile components are also formed, resulting in small amounts of methane and ethane in the gas.

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ELSEVIER PII: SOO16-2361 96)00132-9

Fue l Vol. 75, No. 12, 1449-1456, 1996p.Copyright 0 1996 Elsevier Science Ltd

Printed in Great Britain. All rights reserved0016-2361/96$15.00+0.00

Thermodynamic simulations of lignite fired

IGCC with n s tu desulfurization and CO2

capture

W im F. Elseviers Tania Van Mierlo Martine J. F. Van de Voorde and Harry

Verelst

Department of Chemical Engineering, Vrue Universiteit Brussel, Pleinlaan 2, B- 1050

Brussels, Belgium

Received 18 April 1996; revised 73 June 1996)

IGCC cycles running on Greek lignite with i n i t u desulfurization and CO 2 capture w ere simulated using the

ASPEN PLUS process simulator and compared with other power cycles. Dolomite was chosen as the

material for i n s i t u desulfurization. Considering the pow er produ ced and residual sulfur content of the off-

gas, an oxygen-blown gasifier operated at 1.8 MPa and 95 O” C, having a thermal efficiency of 38.9%, gives

the best performan ce. The efficiency loss due to CO2 capture can be largely compensa ted using IGCC cycles.Cop y r i g h t 0 1 9 9 6 E l s ev i e r Sc i en c e L t d

Keywords: IGCC; CO* capture; simulation)

The applicatio n of oil- and gas-fired combustion turbinecycle systems has grown rapidly because of the lowercost, higher efficiencies and improved gas turbinereliability. Because of supply limitations and theincreasing costs of conventiona l clean fuels such as oiland gas, the use of coal in comb ined cycles for electricaland thermal energy production is gaining interest.Integrated gasification com bined cycles (IGCC) withhigh efficiencies and low emissions, due to integrated

gas cleaning , are an attractive alternative for powerproduction from solid fuels’)2, as show n by several

demonstration plants: Buggenum (The Netherlands)3,Reno (USA )4 and Puertollano (Spain)‘16. Modern coal

gasification techn ologies present a unique opportunity tocombine the adva ntages of high-efficiency com bined-cycle power generation with an environme ntally friendlycoal-based process. IGCC power stations are a majorstep forward in producin g electricity a t competitiveprices while significantly reducing emissions throughadvances in low-NO, gas turbine combustion technol-ogy and intensive desulfurization processes.

GASIFICATION BASICS

For a better understanding of the subsequent discussionof the simulation results, a short overview of gasification

basic reactions, as used in this model, is given.Solid fuels with high moisture and relatively low fixed

carbon contents are not suited for direct comb ustion. Asan alternative, they can be converted to a comb ustiblegas containing mainly H2 and C O in a gasificationprocess. The gasification gas can be used for firing a gasturbine or for the production of several basic chem icalssuch as NH3 or acetic acid from synthesis gas after steamreforming.

The prime reactions taking place in solid fuelgasification can be summarized as follows. The heat

required for the endotherm ic Boudo uard reaction

C(s) + CO2 g) = 2CO ,)> AH298 = 283.1 kJ mol-’

(1)

and the reactions of the fixed carbon with steam

C(s) + H2O(s) = CO(s) + H2 4>

AH2s8 = 131.4kJmol-’ (2)

Cc,) + 2H20 ,) = COz ,) + 2H2 g),

AH29s = 90.28 kJmol_’ (3)is delivered through partial combustion:

C(s) + @2 g) = CO ,), AH2s8 = -110.6kJmoll’

(4)

C(s) + 02 g) = C02 g)r AH29s = -393.7 kJmol_’

(5)

If pure oxygen from a cryogenic air separation plant isused as oxidizing agent, the gas produced consists mainlyof CO, C02, H2 and H20 . N2 will be present if air is usedinstead of oxygen. The gas comp osition is determine d

through the equilibrium of the water g as shift reaction:

CO(s) + H20 , ) = C02cg) + H2 g)>

AH29s = -41.10 kJmol_’ (6)

The above-mentioned reactions are just a small selectionof possible reaction s, as upon pyrolysis a numb er ofvolatile componen ts are also formed, resulting in sma llamounts of methane and ethane in the gas.

Fuel 1996 Volume 75 Number 12 1449

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As the gasification takes place in a reducing atmo-sphere, sulfur, nitrogen and chlorine species present inthe solid fuel will be converted to H2S, NH3 and HClrespectively. H2S is in equilibrium with C OS, as can be

seen from

Hz?,) + COW = COS(,) + J320 ,),

AZZ2s8= 33.39 kJ mol-’ (7)

Due to their corrosive na ture, these compo nents willrequire special care in mate rials selection. H owever,upon combustion in the gas turbine, sulfur components

will be converted to SO2 whereas NH3 will yield N O,,making a gas cleanup system necessary to comply withenvironm ental emission regulations. The H2S content of

the gas can be lowered using i n s i t u desulfurization in thegasifier. If still lower H2S concentration s are desired,addition al intensive desulfurization becom es necessary.

As gasifier gas volume flows are much smaller than gasturbine exhaust gas flows, such cleaning preferably takes

place between gasifier and gas turbine.In s i t u desulfurization can be carried o ut by injecting a

sorbent into the gasifier. Mineral carbonates, such aslimestone or dolomite, are already in use for thispurpose. In general, for a metal carbonate MeC Os thefollowing reac tions take place upon injection in a

pressurized gasifier operating at 900°C. The metalcarbonate can decompose up to a certain C O2 partialpressure through the calcination reaction:

MeC Os(,) + MeO (,) + CO2(s) (8)

Kp = Ptot [CO21  9)

The oxide can then be converted into the sulfide:

MeO(,) + H&s) + MeS(,) + HzO(s) (10)

Kp = PW1/[HA (11)

At higher C O2 partial pressures, calcination will notoccur and desulfurization will be achieved throug h directreaction with the carbona te:

MeCOs(,) + HzS(,) + MeS(,) + HzO(,) + COZ(,)

(12)

Z$ = ~tot [CO,1 [H201 /[H2Sl (13)

Comparing Equations (11) and (13), it is clear thatdesulfurization with the calcined material is independentof pressure, while residual H2S will increase withpressure if no calcination occurs. It can also be seenthat an increase in CO2 and Hz0 partial pressures w illhave a negative influence on residual H2S.

COS can be removed according to

MeC 03(,) + COS(,) + MeS(,) + 2 COQ ) (14)

MeOg) + COS(,) * MeS(,) + COQ) (15)

COS can also be removed through the equilibrium withH2S (reaction 7). As can be seen from the reactionequilibria, it is clear that a change in the shift reaction

equilibrium will have a strong influence on desulfuriza-tion performance.

IGCC SIMULATIONS

To evaluate process performance and econom ics, anumber of IGCC cycles running on Greek lignite or

Polish bituminous coal, with i n s i t u desulfurization andoff-gas CO2 capture, were simulated using the ASPEN

PLUS process simulator. As simulation results for asimilar run of the Buggenum power plant show verygood agreement with the experimental data supplied, it

can be concluded that ASPEN PLUS is a suitable andpowerful tool for this kind of work7 .

Fuel ana l y s i s

The lignite feed compo sition and calculated lower

heating value (LHV) are shown in Tab l e 1 .

Process descr ip t ion

The process flow diagra m is represented in Figure 1.

A flow of 20 kgs-’ (db) of Greek lign ite containing

60 wt% moisture (db) is dried to 15 wt% moisture byindirect steam with vapour recompression’ and fed to thegasifier. Compressed oxygen (95 vol.% O2 and 5 vol.%Ar) from a cryogenic air separation plant or comp ressed

air is used as oxidizing agent. The oxygen or air flow rateis regulated to obtain adiabatic gasification under well-

defined operating conditions. Partial combustion of thefuel supplies the heat necessary for the endothermicgasification reactions. Steam is added in a steam/fuel

ratio of 1 : 20 (w/w ) to activate gasification reactions.Furthermore , recycled CO2 off-gas or steam can be

added to obtain good fluidization, yielding a superficialvelocity ranging between 10 and 30 times minimumfluidization velocity (u,r). A sorbent (dolom ite, lime-

stone or zinc carbonate) is added in an active-metal/sulfur ratio of 2 : 1 for i n s i t u desulfurization. A cyclonesystem removes entrained solid particles from the fuelgas. Cooling the gas to 500°C in a ceramic heatexchang er with heat recovery’ allows the volatile alkalicomp onents to condense on the fly ash particles to beremoved by ceramic candle filters”.

The solid-free gas is burnt in a gas turbine unit usingthe sam e oxidizing agent as in the gasifier, i.e. 95 vol.%02-5 vol.% Ar or compressed air. CO2 off-gas is used as

a moderator to regulate the gas turbine inlet tempera -ture. Corrosion restrictions on the gas turbine imposea maximum allowable concentration of 1800 ppmvsulfur”, even if sulfur is less problem atic than thevolatile alkali compounds. After expansion to atmo-spheric pressure , the exhaust g ases are further cooled toatmospheric temperature in several heat exchangersthrough exchanging heat with water and steam tooperate an optimized, three-pressure-leve l recuperativeRankine bottom cycle, which also produces the gasifiersteam, Condensed water is neutralized with an alkalinesolution. The off-gas, containing mainly C O2 and Ar, iscompressed to 8 MP a in a four-stage compressor with

Table 1 Lignite composition

Proximate analysis Ultimate analysis

(wt% db) (wt% db)

Ash analysis(wt%)

Moisture 15.0 Ash 16.6 SiOz 26.57FC 36.2 C 47.2 Al203 14.01VM 47.2 H 4.68 CaO 33.10Ash 16.6 N 1.27 MgO 1.26

Cl 0.20 Fez03 7.99S 0.99 SO3 17.07

0 29.1

LHV (MJ kg-‘) 18.62

1450 Fuel 1996 Volume 75 Number 12

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Combustion xygen----

Heat to rankine cycle‘_____ _ _ __ - - -

Fhidization CO

Figure 1 IGCC process flow diagram

intercooling to obtain a CO z-rich ga s, ready for safe

disposal12. The transport of CO2 in normal high-tensilecarbon steel pipelines requires a gas above dewpoint atany condition in the pipeline system . Practically, this

means a dewpoint below -4°C. SO2 concentrationlimits depend upon the CO2 off-gas application. IfCO2 injection below ground into a depleted oil or

gas reservoir or some other underground structure isapplied, SO2 and O2 concentrations should be <l ppmvto avoid g rowth of biological species. If CO 2 is to be

injected into the sea at a suitable d epth, SO2 and 02concentrations can be as high as 1 vol.% as long as no

water is present in the gas. Thus, off-gas residual SO2level is determined by both gas turbine tolerances andCO2 off-gas destination.

Simulation input

Component physicochemical properties are calculated

using the Redlich-Kwong-Soave equation of state. Forpure water streams, ASME steam tables were used. For acorrect calculation of the sour water formation, it wasnecessary to use the non-random two-liquid (NRTL)electrolyte property mode l. In the ASP EN P LUSsimula tions, coal is treated as a so-called n on-conven-tional solid. Thermodynamic and physical parametersare derived from the proximate , ultimate and sulfuranalyses. Since chem ical reaction calc ulations are possi-ble only for pure com ponents, coal has to be converted

into real chemica l sub stances , i.e. its constituent ele-ments. Coal mineral matter, or ash, is left as a non-conventional solid, as it can be considered inert for allgasification reactions. Table 2 lists the components usedin the simulations.

Coal properties were estimated using the IGT coal

density model. For low-rank coa l containing significantamo unts of incompletely converted bioma terial, the Boiecorrelation (Equation 16) seem s to give a better accuracy

Table 2 Components in simulation

Phase Components

Gas/liquid 02, Nz, Ar, HZ, HzO, CH4, C02, CO , COS,H2S, NH?, H CN, HCl, Cl*, NO, NOz, S02, S2

Solid

Non-conventional

C, S, SiOl, A&O,, Fe203, Fe& FeS2, CaO,CaCOs, CaS, CaS04, CaC&, MgO, MgC03,

MgS, MgSO,, M gCl2, CaMg(C03)2ash, coal

than the Dulo ng correlation for the calculation of lower

heating value (LHV) in MJ kg-’ (daf):

LHV =35.17xc + 116.25~” - 11.10~~ + 6.28~~

+ 10.47~s - 0.4396 (16)

where fuel elemental composition is expressed as mass

fractions xi on a dry, ash-free basis.The chemical reactions in the gasifier and the gas

turbine are assumed to reach thermodynamic equili-brium and concentrations are calculated using the Gibbsfree energy minim ization method .

Compressors and gas turbine are calculated using

rigorous polytropic law s with a polytropic efficiency of88% (compressor) or 87% (turbine) and 99% mechan-

ical efficiency.Condenser, heat exchanger and burner friction losses

are calculated as respectively 2, 3 and 4% of thecorresponding inlet pressure. A pressure drop of 5 kPais used for the ceramic candle filter unit13. The pressuredrop across the fluid bed gasifier was estimated at 13 kPa.The gasifier packed bed height is 0.5m. The calculatedminim um fluidization velocity umf of 0.03 ms-’ atminimum fluidization voidage cmf = 0.56 (ref. 14)corresponds well with experimental values ranging

from 0.011 to 0.038 m s-’ (ref. 15 ).

SIMULATION RESULTSLignite ga cation

Recuperative Rankine cycle. As further discussio n ofthe results w ill deal with overall net efficiency of bothIGCC top cycle and recuperative Rankine bottomcycle, optimization of the latter will be discusse d first.A process flow diagram of the Rankine cycle is given inFigure 2.

The proposed Rank ine bottom cycle for optimizationis a three-pressure-level cycle with reheating. Gasifica-tion steam is delivered at the desired pressure througha draw-off in the appropriate steam turbine. The

following technological limitations were imposed onthe simulation:

turbine isentropic efficiencies of 87%, mechanicalefficiencies of 99%no condensation in high- and intermediate-pressureturbines (HPT, IPT)conden sation in low-pressu re turb ine (LPT) is

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w

I G\

Figure 2 Recuperative Rankine cycle process flow diagram. HPT ,high-pressure turbine; IPT, intermediate-pressure turbine; LPT, low-pressure turbine

Table 3 Rankine cycle optimization results

Level

HP

IPl

IP2

Pressure (MPa) Power (MW)

6.0 90.94

10.0 92.7811.5 93.14

1.3 92.421.4 92.571.5 92.690.3 92.190.5 92.69

0.7 92.64

Steam temp. (“C) Power (MW)

400 89.06500 90.57580 91.93

allowed , isentropic efficiency decreases by 1% for each

1% of liquid in the turbine outlet streammaximum pressure ratio for HPT and IPT 30, LPT 70

maximum steam temperature 580°C (exchangermaterial limitations)

maximum steam pressure 11.5 MPaminimum condenser pressure 5 kPa (35°C dewpoint)heat exchanger pressure drop 3% of inlet pressure

gas-liquid heat exchange requires minimum AT of10 K a t the pinch pointgas-gas heat exchange requires minimum AT of

100 K at the pinch point.

Heat from the gas turbine top cycle is available from

two sources:ceramic heat exchange r after gasifier (hot side 900 to

500°C)

Gas t u r b i ne opera t i on . In the subsequent discussion,the gasifier is operated at 9OO” C, xcept where mentioneddifferently. F or gas turbine operation, inlet temperatu reand pressure ratio need to be determined. As expansionis always to atmospheric pressure, the process parameterdetermin ing the pressure ratio is the gasifier pressure.The maximum gas turbine inlet temperature is 12OO ” C,due to materials limitations. Optimum performancewill be dictated by the combination of maximum powerproduction and turbine outlet temperature above680°C the temperatu re needed for the production ofsteam at 580°C. From Figure 3 it can be concluded thatgas turbine inlet temperature should be as high as possi-

ble for maximum power production, i.e. 1200°C. Powerproduction increases by 1 MW for each 50 K increase

in inlet tempe rature. It is evident th at power productionincreases with gasifier pressure and thus pressure ratio,so that inlet pressure should also be as high as possiblefrom the power produ ction point of view.

heat recovery from gas turbine exhau st gases (hot side

780 to 65°C)

Taking into account th e technological limitation s, The total heat available for steam production is

optimal cycle parameters can be deduced from the indepen dent of the inlet tempera ture. It can be seen

results given in Table 3. Increasing the high-pressure that at the maximum inlet temperature of 12OO” C, he

(HP) level to the maximum allowable of 11.5 MPa yields gasifier pressure has to be <2.8 MPa to comply with the

Table 4 Heat distribution in Rankine cycle

Available heat Available from

an efficiency increase of 0.6% . Gasification steam at1.5 MPa is taken from the high-pressure turbine (HPT).Increasing the HPT exit pressure (IP level 1) to thegasifier pressure level also increases Rankine cycle netpower prod uction. If the gasifier is operated at other

pressures, a similar conclusion can be drawn. Theoptima l IPT outlet pressure (IP level 2) lies between 0.5

and 0.6MPa. Since this would lead to LPT expansionratios of 100 to 12 0, the IPT outlet pressure has to be

kept at 0.3 MPa , ca using a minor power reduction of-0.15%. Reheating the steam to 580°C before expansionin the LPT gives best cycle performance.

The available heat is distributed as shown in Tab l e 4

over the various countercurrent heat exchang ers and

boilers, so that no condensation occurs in the IPT andmaximum power is produced. As increasing thermalinput in the intermediate pressure boiler will always lead

to a lower power output, the recovered heat is best usedin the high-pressu re steam boiler. For ga sifier operationat 1 .5 MPa this results in a 37 MW thermal input for the

intermediate-pressure boiler.

Cycle ove ra l l ef i c i ency ca l cu la t i ons . Cycle total ther-mal input is 373 MW . Power output is delivered from the

gas turbine and Rankine cycle. Fuel drying costs4.3MW , and the COz removal cost includes both theproduction of 95 vol.% 02-5 vol.% Ar from a cryogenic

air separation plant (22.1 MW ) and the CO2 off-gas com-pression (13.1 MW ). Recycle gas and oxygen com pres-sion are the remaining costs, which are proportional totheir respective flow rates.

Distributed over

900 to 500°C

780 to 130°C

130 to 65°C

alkali removal cooling

gas turbine exhaust gas cooling

gas turbine exhaust gas cooling

intermediate-pressure steam reheater

intermediate-pressure steam boilerhigh-pressure steam boilerintermediate-pressure steam reheater

high-pressure water economizerintermediate-nressure economizer

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0

Figure 3 Influence of gasifier pressure and gas turbine inlet

temperature on gas turbine performance. Gas turbine inlet temperature(“C): a, 1000; b, 1100; c, 1200

outlet temp erature limit of 680°C . Howeve r, the gasifierpressure w ill be further limited by environmental SO2considerations, as will be shown below.

F lu id i zed bed load ing capac i t y . Kinetic considera-tions impose a maximum fluidized bed loading capacityof 4000 kg h-’ m-* at 900 °C to maintain a gas spacetime of 1.0-l 5 s. Increasing the loading capacity willyield a smaller gasifier surface area for a given feedrate, but will reduce gas space time since gas superficialvelocity is inversely proportional to gasifier surfacearea. Furthermore, lowering the desired superficial velo-city for proper bed fluidization will increase gas spacetime at a given bed loading. Operating at a higher bedloading and a lower g as velocity will shift to a higherlevel the critical pressure at which extra C02/H 20 fluidi-

zation gas has to be added. Gas residence time will thusbe determine d by the comb ination of gasifier loadingcapacity and the desired g as superficial velocity forproper bed fluidization.

For com bustion in the gas turbine, fuel gas LHV is animportant parameter. It is calculated in MJkg-’ fromcompo sition, yi being the mass-fraction of compon ent i :

LHV = 10.lyco + 122.7~~~ + 50.6~~~~ (17)

As can be seen from Tab le 5 fuel gas LHV decreasesrapidly when C O2 addition is required. The shift of thewater gas equilibrium towards CO cannot compensatefor the overall decrease in H2 and C O mass fractions dueto the strong dilu tion effect: whe n additional fluidizationgas is required, its flow rate is of the same order ofmagn itude as the flow rate of flue gas leaving the gasifier.

Table 5 Effect of bed loading and gas velocity on cycle performancea

Gasifier co2 to Fuel gas Fuel gas SOr inpressure gasifier LHV flow Efficiency

(MPa) (kg s-‘) (MJ kg-‘) (kg s-‘) ;;I%; W)

ia; 3 t h-’ mm2,20~ (0.7 m s-i)

23.1 5.1 55.2 483 36.81.5 46.2 3.4 78.8 837 37.7

1.8 68.7 2.5 102.2 1147 37.82.1 91.2 1.9 125.6 1418 38.0

(b) 4 t h-’ rne2, 15~ (0.5 m ss’)

1.2 0.0 10.1 29.8 154 37.21.5 0.0 10.2 29.6 196 38.31.8 7.3 7.9 37.5 399 39.02.1 19.0 5.8 49.9 742 39.3

a Steam 1 kg s-’ to gasifier

01

Figure 4 Influence of gasifier pressure and temp erature on cycleefficiency and SO2 emission. Gasifier temperature (“C ): a, 850; b, 900;c, 950

Adding extra steam converts CO into HZ, giving adecrease in CO mass fraction. Becau se dilution effectsand formation through reaction compensate each other,

the H2 mass fraction remains nearly constant. Thisresults in an overall LH V decrease when more steam isadded. A s total flow increases upon steam addition, CO2addition for proper fluidization is necessary only at

higher pressures at the same bed loading and desired gasvelocity. Clearly, a higher bed loading and a lower gasvelocity are preferable for both cycle overall efficiencyand minimal SO2 concentration in the off-gas produced.

Gas i J i er opera t i on . As concluded above, gasificationwill take place at a bed loading of 4000 kg h-’ m -* and0.5 m s-l gas velocity. U nder these circumstance s, lignitegasification kinetics require temperatures >85O” C. At

9OO”C,gasification is -10 times faster16. Within this tem-perature range, gasifier operating conditions will bedetermined by cycle overall efficiency and SO 2 emissio ns.

Figure 4 shows cycle overall efficiency and SO2concentration in the CO 2 off-gas as functions of gasifierpressure an d temperatu re. Cycle overall efficiencyincreases with pressure. At a fixed gasifier pressure,efficiency decreases with increasing reaction temperature,due to higher exergy losses between gasifier and alkaliremoval. SO2 concentration in the CO2 off-gas increases

with pressure. If 250ppmv SO2 is accepted as a reason-able emission limit, optimal process parameters can be

determined. At 900°C and pressures ~1.8 MPa, SO2emission is too high at all temperatures. Operation at1.5MPa yields 196ppmv SO2 with 38.4% efficiency.Increasing the pressure to 1.8 MPa would increaseefficiency to 39.0% with 399ppmv SO2 emission. How-ever, increasing the gasifier temperatu re to 950°C

produces a CO2 off-gas containing only 217 ppmv SO*,while efficiency increases to 38.9%. This suggests the useof a somew hat higher gasification tempera ture, together

with a higher inlet pressure of the gas turbine unit.

F l u i d i z a t i o n agen t t y p e. Gasifier operation at 900°Crequires the addition of CO 2 or steam for a good fluidi-

zation at pressures > 1.5 MPa . The influence of this addi-tion on process performance at 1.8 MPa can be seen inTab l e 6 .

Steam gasification is superior to CO2 gasification fromthe efficiency point of view. Taking into account SO2emission , C O2 gasification gives a slightly better perfor-mance. H owever, since at elevated pressures, whereoxygen gasification is insufficient for proper fluidization,

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Table 6 Influence of fluidization gas type on process performance at1.8 MPa and 900°C

Overall

Steam to gasifier CO 2 to gasifier SO2 in CO 2 offgas efficiency

(kg s-l) (kg s-‘) @pmv) W)

0.0 8.9 319 38.8

1.0 6.1 399 39.0

2.0 4.5 417 39.13.0 2.3 434 39.3

4.0 0.0 448 39.5

SO2 levels will nearly always require an intensive colddesulfurization unit, steam gasification is preferable to

CO * gasification because of the strong efficiency increase.

Air gasification. Because the use of air as oxidizing

agent introduces a large amount of nitrogen in the gasi-fier, steam or CO2 addition for good fluidization will benecessary only at much high er pressures. Power produ c-

tion is 183.0MW at 1.5 MPa and drops to 173 .5 MW at1.8 MPa , indicating a lower pressure level for air-blowngasification. Process off-gas consists mainly of CO 2 andNZ, C02-N2 separation using an off-line DEA absor-ber-stripper unit will produce a COz-rich ga s ready forsafe disposa l, if required. DEA solvent regene rationcosts amount to 10 0.3 MW (2.94 MJ kg-’ C0217, whichcan be withdrawn from the gas turbine cycle heat recov-ery, decreasin g Ra nkine cycle power production by38.7 MW . Off-gas compression to 8 MPa requires only

11.8 MW , due to the ab sence of Ar. Cycle overall effi-ciency is 35.5% with CO2 capture. Without CO2 capture,cycle overall efficiency is 49.0% .

In situ desulfurization through sorbent injection. In

situ desulfurization is effected by injecting dolomite(CaMg (CO& ), limestone (CaCO & magnesium carbo-nate (MgC 03) or zinc carbonate (ZnC03) into thegasifier. Figure 5 shows the influence of pressure andsorbent materia l on residual H2S in the gasifier off-gas.Possible desulfurization properties of the ash constitu-

ents are not taken into account. It can be seen that with-out desulfurization, under the specified operatingconditions, residual H2S levels are in the region of

4000 ppmv. At pressures > 1.6 MPa , where COz additionis required to main tain 0.5 m s-l gas velocity, the concen-

tration drops d ue to the dilution effect, althoug h the totalamount of sulfur in the fuel gas remains the same.

Magnesium carbonate shows no desulfurization prop-

4500 900

4000

3500

3000

2500

_~ _._b .-e--c -*-d

2000

500 C.*. *.*, .*, .*. .~.*~~~.~*, *+*= ) 100

)_*_C... ~__, _~_*_*_~.. .~_. .. *__. ~-- -~ ~-~-~-' -~-' -'

0 0

0 5 10 15 20 25

Gaslfler ~r~rmur~ [bar]

igure 5 Influence of gasifier pressure and sorbent material ondesulfurization performance. Sorbent material: a, none; b, MgC OS;c, CaC 03; d, dolomite; e, Z&O ,

erties and must be rejected as an in situ sorbent. A slightdecrease in residual H2S level is entirely due to thedilution through CO2 formation during calcination.Magne sium sulfide is not formed at all.

Dolomite, which can be considered as a mixture ofdolomite and limestone, behaves about the same as

limestone, showing that its desulfurization capacity can

be entirely attributed to the presence of a CaC Os phase.Zinc carbonate show s a very good desulfurizationbehaviour over the entire pressure range studied.

Residual H2S levels and sulfur removal percentages at1.5 MP a and 900°C for each sorbent are listed in Table 7.

Although ZnC03 clearly gives the best performance, it

must be rejected for in situ desulfurization because ZnO,formed up on calcination at high tempera tures, is easilyreduced to the volatile zinc metal in the reducing gasifieratmosphere, giving Zn concentrations in the vapourphase in excess of the gas turbine tolerance limit of2ppmv Zn. As limestone and dolomite show nearly thesame behaviour, either material would be suitable.

However, taking into account desulfurization kinetics,dolomite has a better perform ance, because of its moreporous structure due to the MgC Os calcination, resultingin 100% Ca-sorbent usage, contrasted with 20% forlimestone”.

With dolomite injection, the residual H 2S concentra-tion increases gradually from 1 O to 1.6 MPa and more

strongly at higher pressures, as can be seen from Figure 5.

At 1 OMPa total pressure and under the given gasifica-tion conditions, the CO2 partial pressure corresponds tothe equilibrium for the calcination reaction. As theresidual H2S level decreases as CO2 partial pressure

approac hes calcination reaction equilibrium for both

calcining and non-calcining regions, it is clear that thispoint g ives the best desulfurization performance’g >20. At

pressures >0.8 MPa, calcination no longer occurs andthe residual H2S level increases with pressure. Since atpressures <1.4 MPa no additional gas is required forgood bed fluidization, only small changes in the watergas shift equilibrium occur, due to the lower CO 2 release

from the calcination reaction. CO 2 off-gas recycle to thegasifier for proper fluidization at higher pressu res ha s amuch stronger effect on the water gas shift equilibrium .CO2 and H20 partial pressures increase, resulting in arapidly inc reasing residual H2S level.

Extra steam ad dition increases the flue gas flow rate,making CO2 fluidization gas addition necessary only athigher pressures. Nevertheless, this advantage cannot

comp ensate for the direct influence on desulfurizationequilibria, as can be seen from Table 8.

Initially, a metal:sulfur ratio of 2 : 1 was chosen for thedesulfurization simula tions. The influence of Ca/S ratiofor dolomite desulfurization is show n in Table 9. From a

thermodynamic point of view, a stoichiometric Ca/Sratio gives the best results. Increa sing Ca/S ratio releasesexcess C02, which influences the water gas shift

Table 7 Desulfurization performance at 1.5 MPa and 900°C for

different sorbent materials

Sorbent H in gasifier off-gas (ppmv)

None 4105

MgCO3 4058

Dolomite 124

Limestone 112

ZnCOl 49

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Table 8 Effect of steam addition on desulfurization performance using

dolomite at 900°C

1.5 2.5Gasifier

pressure (MPa) HIS in HzS inHz0 to gasifier co2 to gasifier co* togasifier

(kg s-‘)

off-gas gasifier off-gas gasifier

(ppmv) (kg s-‘) (ppmv) (kg s-‘)

1 124 0 839 33.82 153 0 869 31.64 210 0 919 27.4

Table 9 Influence of Ca/S ratio for dolomite desulfurization

Ca/S molar ratio HzS + COS in off-gas (ppmv)

0.5 2123

0.9 419

1.0 121

1.5 125

2.0 129

4.0 146

equilibrium. However, using 100% sorbent excess hasonly a small influence. Taking into account incompletesorbent utilization and poor behaviour at sub-stoichio-

metric conditions, 100% excess of i n s i t u sorbentmaterial seems acceptable to obtain realistic simulationresults.

In situ desu l f u r i za t i on by coa l ash . Desulfurizationoccurs at elevated temperature s through reaction of

metal oxides or carbonates with H$S. From the coalcomposition in Tab le 1 it is clear that coal mineral mat-

ter, or ash, may also have desulfurization capacity,because of the presence of CaO and Fe20,. CaO assuch is already present in the ash in a Ca : S ratio of-3 : 1, which is substantia lly sufficient for proper desul-furization.

Figure 6 shows residual sulfur (HzS + COS) levels fori n s i t u ash desulfurization with and without extradolomite addition . It can be seen that desulfurizationefficiency decreases upon sorbent addition, due to theCO 2 release from calcination. Heat required for heatingthe cold sorbent to the operating conditions is produced

through a slight increase in partial com bustion of thecoal, producing additional CO2 as well as HZ0 throughthe water g as shift equilibrium, lowering desulfurizationefficiency. Resid ual levels for ash-only desulfurizationare some what lower than for dolomite-only desulfuriza-tion. If ash participate s in desulfurization in practice, no

80 i

Figure 6 Influence of gasifier pressure and dolomite addition on in

situ desulfurization using coal ash: a, ash o nly (Ca /S = 3.1);b, ash + 1 kg s-’ dolomite (Ca/S = 3 .5); c, ash + 2.28 kg ss’ dolomite(Ca/S = 5.1)

i n s i t u sorbent addition will be necessary, reducing

operating costs and solid waste disposal problems.

Com pa r i son o f l i gn i t e- j i r ed cyc les . Tab le 10 gives an

overview of the cycles studied, all fed with 20 kgs-’Greek lignite, having a 373 MW thermal input. D olomiteis used in a Ca : S ratio of 2 : 1 for i n s i t u desulfurization.

Taking into account both efficiency and SO2 emission, ahigher bed loading (4000 kg h-’ mh2) and a lower gas

velocity (15~~) result in an optimu m performance for02-CO2 gasification at 950°C and 1.8 MPa. The use ofsteam as gasification and fluidization agent increasesefficiency, but sinc e SO2 emission levels also increase,an additional desulfurization unit between gasifier andburner becomes necessary. Because cold desulfurization

inevitably leads to high exergy losses, therma l efficiencywill be much lower. Operating costs will increase as well.

Quantitatively, it is clear that CO 2 capture decreasesair gasification cycle efficiency from 49.0 to 35.3%,resulting in a 3.6% drop in efficiency compared with

38.9% for the optimal C02-O2 combined cycle. Con-sidering the optimal C 02-O2 gasification and airgasification withou t CO 2 capture overall efficiencies, itcan be concluded that the cost of zero CO2 emissionamounts to 10.1% thermal efficiency.

B i t um i n ou s coa l ga s l j i c a t i o n

Similar simulations were made u sing Polish bitumi-

nous coal. The conclusions are mainly the same,indicating that the simulation model can be used forvarious coal types. For bituminous coal, CO2 or steamaddition is necessary not only for proper fluidization butalso for a complete adiabatic gasification at 1200°C . At1.5 MPa and u sing dolomite for i n s i t u desulfurization,

Table 10 Overview of performance parameters of lignite-fired cycles studied

Process parameters Process performance

(MPa) T (“C) Remarks Power (MW ) Efficiency (%) SO2 in COz off-gas (ppmv)

02-co* 1.5 900 3 t hh’ m-*, 0.7 m s-l 140.6 37.1 83702 1.5 900 4 t h-’ m-‘, 0.5 m s -I 142.7 38.3 196

02-co* 1.8 950 4 t h-i m-‘. 0.5 m s-l 145.0 38.9 217

oz-co* 2.1 900 4 t h-’ m-*, 0.5 m s-i 146.6 39.3 742

02-Hz0 2.1 900 4 t h-’ m-*, 0.5 m s-’ 152.1 40.8 989

Air 1.5 900 ditto, DEA CO2 capture 132.5 35.3 <lOO

Air 1.5 900 ditto, no CO2 capture 183.0 49.0 not applicable

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H.$ residual levels of 60-70ppmv are obtained, whichincrease to 120-140ppmv for operation at 2.1 MPa .Residual sulfur is minimal when either pure steam or

pure carbon dioxide is used as the gasification medium,while mixed steam-carbon dioxide gasification alwaysdecreases the in situ sulfur removal efficiency. Steam orcarbon dioxide does not seem to have a large influence

on the global power production when optimizing boththe gas turbine and the bottom steam cycle, although

steam gasification leads to somewhat higher efficiencies(+0.3% ). Therma l efficiencies of 41.2 and 42.7% a re

obtained at 1.5 and 2.1 MPa gasification pressurerespectively for pure steam gasification.

CONCLUSIONS

Several IGCC -type gasification cycles runnin g on Greek

lignite or Polish bituminous coal with in situ desulfuriza-tion and COz capture have been simulated. Simulationresults are similar for both coal types. IGCC top cycle

gives the best performance when the gas turbine isoperated at its maximum allowable temperature of1200°C. Using an optimized three-pressure-level Ran-kine bottom cycle, the optimu m cycle therma l efficiencyis 38.9% for a lignite-fired gasifier operated at 950°Cand 1.8 MPa. Using dolomite as in situ desulfurizationsorbent, an SO2 residual level of 217ppmv in the COzoff-gas can be achieved. Howeve r, when the coal ashcontains enough active materia l, desulfurization sorbentinjection may not be necessary. The use of steam as

fluidization agent incre ases efficiency at the cost of higherSOz concentration. When zero CO2 emission is required,it can be concluded that the costs of the CO2 removal can

largely be comp ensated by the use of an integrated cycle.

ACKNOWLEDGEMENTS

This work has been funded by the Commission of theEuropean Comm unities, DG XII, JOULE II programmeunder contract PL 920567. T. Van M ierlo is presentlyVUB-OZR research assistant.

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