steady state

Steady-State Simulation and Optimization of an Integrated Gasification CombinedCycle Power Plant with CO2 Capture

Debangsu Bhattacharyya,*,†,‡ Richard Turton,*,†,‡ and Stephen E. Zitney*,‡

Department of Chemical Engineering, West Virginia UniVersity, Morgantown, West Virginia 26506, United States,and Collaboratory for Process and Dynamics Systems Research, National Energy Technology Laboratory,Morgantown, West Virginia 26507, United States

Integrated gasification combined cycle (IGCC) plants are a promising technology option for power generationwith carbon dioxide (CO2) capture in view of their efficiency and environmental advantages over conventionalcoal utilization technologies. This paper presents a three-phase, top-down, optimization-based approach fordesigning an IGCC plant with precombustion CO2 capture in a process simulator environment. In the firstdesign phase, important global design decisions are made on the basis of plant-wide optimization studieswith the aim of increasing IGCC thermal efficiency and thereby making better use of coal resources andreducing CO2 emissions. For the design of an IGCC plant with 90% CO2 capture, the optimal combinationof the extent of carbon monoxide (CO) conversion in the water-gas shift (WGS) reactors and the extent ofCO2 capture in the SELEXOL process, using dimethylether of polyethylene glycol as the solvent, is determinedin the first phase. In the second design phase, the impact of local design decisions is explored considering theoptimum values of the decision variables from the first phase as additional constraints. Two decisions aremade focusing on the SELEXOL and Claus unit. In the third design phase, the operating conditions areoptimized considering the optimum values of the decision variables from the first and second phases asadditional constraints. The operational flexibility of the plant must be taken into account before taking finaldesign decisions. Two studies on the operational flexibility of the WGS reactors and one study focusing onthe operational flexibility of the sour water stripper (SWS) are presented. At the end of the first iteration,after executing all the phases once, the net plant efficiency (HHV basis) increases to 34.1% compared to32.5% in a previously published study (DOE/NETL-2007/1281; National Energy Technology Laboratory,2007). The study shows that the three-phase, top-down design approach presented is very useful and effectivein a process simulator environment for improving efficiency and flexibility of IGCC power plants with CO2

capture. In addition, the study identifies a number of key design variables that has strong impact on theefficiency of an IGCC plant with CO2 capture.

Introduction

Meeting the challenge of delivering clean, affordable, andsecure electric power is critical to sustaining the growth andprosperity of human society. With continued focus on the useof cheap and abundant coal resources for electric powergeneration, this compound energy challenge shifts to that ofreducing greenhouse gas emissions, most importantly carbondioxide (CO2). Coal-fired power stations contribute about 20%of the worldwide CO2 emissions arising from the utilization offossil fuels.2 While the fossil energy industry continues toimprove the environmental performance of conventional pulver-ized coal (PC) combustion power plants, advanced technologiessuch as coal gasification offer the potential to generate signifi-cantly lower levels of CO2 and other criteria air pollutants at alower cost of electricity.3-5 Compared to PC power plants,integrated gasification combined cycles (IGCC) also producesmaller volumes of solid wastes,6use 30-60% less water thanthe competing technologies,3 provide greater fuel flexibility,3,7andoffer attractive polygeneration options.8,9 In addition, the penaltyin efficiency and cost of electricity due to CO2 capture is lessfor IGCC systems compared to conventional PC technologies.1,10

For example, a recent study1 shows that the net plant efficiency(HHV basis) of an IGCC power plant with a general electric

energy (GEE)-type gasifier is reduced from 38.2% to 32.5%for 90% CO2 capture, whereas the efficiencies of subcriticaland supercritical PC plants decrease from 36.8% and 39.1% to24.9% and 27.2%, respectively. Another study11 claims that 75%of CO2 from an IGCC plant can be captured with only about4% loss in efficiency.

In recent years, a number of researchers has used steady-state simulation and analysis tools to evaluate IGCC plantperformance and efficiency, including the impact of differentCO2 capture technologies.12-18 When comparing capture op-tions, appropriate importance must be given to the efficiency,reliability, and cost of not-currently available technologies. Forexample, some of the capture technologies presented in theexisting literature such as semiclosed cycles12,16 and the Matiantcycle14 use CO2 as a working fluid in the gas turbine and willrequire development of modified gas turbines. Several IGCCstudies17,19,20 have also shown that improved system integrationcan increase overall plant efficiency. In another recent IGCCstudy, Gnanapragasam et al.21 evaluated the effects of differentfeedstocks, gasifier inlet conditions, and gasifier temperatureson system performance and CO2 emissions. For an IGCC plantwithout CO2 capture, Emun et al.22 carried out a sensitivityanalysis to determine the impact on system thermal efficiencyand environmental performance resulting from variations in thesolids concentration in the coal slurry feed, gasification tem-perature, gas turbine inlet temperature, and level of nitrogen(N2) injection into the gas turbine. Emun et al.22 also appliedpinch analysis for heat integration and analyzed the effects on

* To whom correspondence should be addressed. E-mail:[email protected] (D.B.); [email protected] (R.T.); [email protected] (S.E.Z.).

† West Virginia University.‡ National Energy Technology Laboratory.

Ind. Eng. Chem. Res. 2011, 50, 1674–16901674

10.1021/ie101502d 2011 American Chemical SocietyPublished on Web 12/27/2010

overall IGCC plant efficiency. While these studies used steady-state simulation and analysis tools, they did not consideroptimization for improving the efficiency and flexibility of IGCCpower plants with CO2 capture.

Bahri et al.23,24 presented a two-level iterative approach foroptimal design of chemical processes. In the first level, theoptimal plant design and operating conditions (and controlstructure) are obtained by solving a mixed-integer nonlinearprogramming (MINLP) problem. In the second level, a NLPproblem is solved for investigating the feasibility of the solutionobtained in the first level. This approach, though rigorous, hasa number of drawbacks for a process simulator. First, manyprocess simulators do not have algorithms for constrainednonlinear optimization. Second, as each of the steps requiressolution of an optimization problem that is an iterative proce-dure, the problem can become computationally intractable fora plant-wide simulation. It should be noted that our currentmodel of the IGCC plant with CO2 capture contains more than120 000 equations. Third, when optimization is done in aflowsheet that involves a number of recycle streams, the solutionbecomes very difficult, especially with a sequential modularapproach typical of process simulators. The SELEXOL unit isa good example of a unit involving multiple recycle streams.Fourth, the optimization may involve integer decision variablesas in the first level of the approach proposed by Bahri et al.23,24

Process simulators typically do not offer algorithms for solvingmixed integer programming problems.

The present study describes a three-phase, top-down, opti-mization-based design approach as shown in Figure 1. Each ofthe phases may have subphases, if needed. The optimization ateach phase/subphase can be done by sensitivity studies or byrigorous optimization based on the available algorithms in thechosen simulator envionment. The optimum values of thedecision variables from each phase are treated as constraints inthe next phase. If a phase/subphase involves integer variables,then that phase/subphase is solved by a case-study approachwhere a finite set of options is considered and then the optimumof each option is compared. A number of examples that involvesinteger variables can be found in the second design phase ofthis study. This approach provides the flexibility of the use of

a combination of sensitivity studies and rigorous optimizationthat may involve integer decision variables.

In the top phase of the optimization hierarchy, importantglobal (plant-wide) design decisions aimed at maximizing netenergy efficiency are evaluated considering limits and/or targetson environmental emissions. Because the focus of this paper ison CO2 capture, a study is done by varying the percentconversion of carbon monoxide (CO) in the water-gas shift(WGS) reactors and percent capture of CO2 in the SELEXOLunit to maximize plant efficiency while achieving the target of90% carbon capture for the overall system.

In the second design phase, local design decisions are madeconsidering the integer decision variables. In this phase, theoptimum values of percent conversion of carbon monoxide (CO)in the WGS reactors and percent capture of CO2 in theSELEXOL unit from the first phase are considered as additionalconstraints. The first study on local design decision evaluatesthe use of a single stage flash vessel as the H2S concentrator inthe SELEXOL unit compared to a multistage stripper. Anotherlocal design study compares the technical feasibility and impacton energy efficiency of routing the tail gas from the Claus unitto the H2S absorber in the SELEXOL unit against sending itdirectly to the CO2 compression unit.

In the third phase, the operating conditions are optimizedsubject to the additional constraints from the first and secondphases. Two studies are presented. In the first study, the flowrate of the stripping medium in the H2S concentrator isoptimized. In the second study, the pressures of the flash vesselsin the SELEXOL unit are optimized to reduce the powerconsumption in the SELEXOL and CO2 compression unit.

Before the final design decisions are taken, an assessment ofthe operational flexibility should be done. This ensures that acertain desired level of performance is achieved in face ofdisturbances and uncertainties. Two studies on the operationalflexibility of the WGS reactors are done. Both the studiesconsider that there is a desired upper limit on the amount ofsteam extracted from the steam turbine (ST). As the supple-mental steam in the WGS reactors is provided by extractingsteam from the ST, a larger extraction results in a loss of powerfrom the ST affecting the overall plant efficiency. In the firstdesign phase, the optimum conversion of CO in the WGS reactorsis determined for maximizing the plant efficiency. However, it maybe desired to vary the CO conversion in the WGS reactors for aflexible CO2 capture scenario. The first study evaluates the effectof this scenario on the amount of steam extracted from the ST.The second study evaluates the effect of change in the syngascarbon-to-hydrogen and CO2/CO ratios on the amount of steamextraction from the ST. Another study on the operational flexibilityof a sour water stripper (SWS) considers a desired level of 50ppmw NH3 in the stripped condensate.

In this paper, each phase of the proposed top-down designapproach is executed once using representative optimization studiesfocused largely on maximizing the net power generated from theIGCC plant with CO2 capture. This iterative procedure should berepeated as shown in Figure 1 until some convergence criteria aresatisfied. After that, an assessment of operational flexibility shouldbe done. If found unsatisfactory, the steps should be repeated. Theconvergence criteria can simply be based on a user definedtolerance on absolute or relative differences between the currentand old values of the operating conditions in addition to satisfyingthe criterion that the design decisions taken in the second phaseare unchanged compared to the previous iteration. Other userdefined convergence criteria can also be considered. Since eachphase is executed only once in generating the results presented in

Figure 1. Optimization-based design approach of an IGCC plant with CO2

capture in a process simulator environment.

Ind. Eng. Chem. Res., Vol. 50, No. 3, 2011 1675

the “Results and Discussions” section, the results should not betreated as the final IGCC design. It should be noted that this paperproposes a generic approach that can be applied over a wide rangeof commercial steady-state simulators. If certain features (algo-rithms) are available in a software platform, the user can utilizethat to develop a more systematic approach. For example, the user,instead of doing a sensitivity study, can find out the decisionvariables for an optimization study by generating a gain matrixand then considering the appropriate decision variables for carryingout a nonlinear constrained optimization. It should also be notedthat the present work does not consider changes in the configurationof the IGCC plant (except where mentioned) and does not takeequipment capital cost into account. For an optimal techno-economic design, multiobjective optimization with potential topo-logical changes and consideration of other available technologiesshould be considered.

Plant Configuration

The selection of gasification technology for IGCC depends onvarious factors such as the quality of the feed coal, capitalinvestment, efficiency goals, reliability, availability, cost of electric-ity, environmental targets, and CO2 capture technology.1,6,25 Thegeneral electric energy (GEE) entrained-flow gasifier, previouslythe Texaco technology, currently holds a strong position with about34% of the world’s gasification capacity, including the TampaElectric (TECO) IGCC power generation facility.26 In this paper,a single-stage, oxygen-blown, GEE-type entrained-flow slagginggasifier with radiant syngas cooler is considered.

For gasification-based power plants, a number of precom-bustion CO2 capture techniques is being explored with promisingresults.25,27-29 In this paper, CO shift conversion followed byphysical absorption is used on the basis of the maturity of thetechnology and because of the high partial pressure of CO2 froma GEE-type gasifier. Physical absorption is also very efficientand the least capital-intensive CO2 capture technology that iscommercially proven.27,30 Because of a number of advantagessuch as lower solvent loss, higher selectivity toward H2S, betterthermal stability, better water solubility, and lower circulation

rate,31-33 dimethylether of polyethylene glycol (SELEXOL) isthe physical solvent considered in this study. A techno-economicanalysis shows that a SELEXOL-based CO2 capture process isvery competitive compared to membrane reactors and chemicallooping processes.34

The IGCC with CO2 capture plant configuration used as thebase case for the present study is shown in Figure 2 and is amodified version of the Case 2 configuration described in arecent fossil energy power plant cost and performance com-parison by the National Energy Technology Laboratory(NETL).1 The coal slurry is fed to an oxygen-blown GEE-typegasifier with radiant-only configuration. The hot syngas exitingthe gasifier passes through a radiant syngas cooler (RSC) anda water scrubber. The scrubbed syngas is shifted in two WGSconverters. The shifted syngas is further cooled before goingto the acid gas removal (AGR) process which is a physicalabsorption process with SELEXOL solvent. The sour waterdrained from the syngas coolers is sent to the syngas treatmentunit. The clean water from the sour gas treatment unit is sentto the scrubber and is for slurry preparation. In the first stageof the dual-stage SELEXOL unit, H2S is separated in the stripperand sent to the Claus unit. In the second SELEXOL stage, CO2

is separated and sent to the compression unit for sequestration.Following the AGR process, the cleaned syngas is heated andexpanded before going to the gas turbine combustor wherediluent N2 from the ASU is mixed. The hot flue gas from thegas turbine flows through a heat recovery steam generator(HRSG). In the HRSG, a triple pressure steam cycle generateshigh pressure (HP), intermediate pressure (IP), and low pressure(LP) steam.

IGCC Steady-State Modeling and Simulation Approach

In this section, the steady-state modeling and simulation of thecoal-fired IGCC plant with CO2 capture is discussed. First, the coalanalysis and handling is described. Next, the IGCC plant is dividedinto a number of major, well-defined areas including gasification,radiant syngas cooler and scrubber, shift converters and syngascooling, sour water stripper, SELEXOL unit and CO2 compression,

Figure 2. Layout of the IGCC with CO2 capture.

1676 Ind. Eng. Chem. Res., Vol. 50, No. 3, 2011

Claus unit, gas turbine and the heat recovery steam generator (fluegas side), steam cycle, and air separation unit. All these IGCC plantareas are modeled using the commercial steady-state processsimulator, Aspen Plus Version 2006.5.35

Coal Analysis and Handling. In this study, the primary feedto the IGCC plant is Illinois No. 6 bituminous coal, characterizedby the proximate analysis and ultimate analysis provided inTables 1 and 2, respectively. These coal properties are similarto those for the Illinois No. 6 #1 coal used in tests run at theIGCC power plant at Polk Power station operated by TampaElectric Company (TECO).36 As a result, the TECO projectstatus report36 is used as the basis for specifying operating pointsand ranges for sensitivity studies for many of the IGCCequipment items described in this work.

The coal analysis data in Tables 1 and 2 are used to“decompose” the Illinois No. 6 coal into the following species:C, H2O, H2, N2, S, O2, and ash. Ash is characterized as silicondioxide (SiO2) and “converted” to slag in a stoichiometric reactor(RStoic model in Aspen Plus). The stoichiometry of the reactionin the RStoic model is manipulated so that all of the ash and2% of carbon are removed in the slag assuming 98% carbonconversion in the gasifier. The “decomposed” coal with slagremoved is then fed to the gasifier.

Gasifier. The entrained-flow GEE-type gasifier consideredin this paper is modeled as a restricted chemical equilibriumreactor with temperature approaches to equilibrium for indi-vidual reactions. Using a more simplistic approach, some authorshave assumed that the syngas at the gasifier outlet exists atchemical equilibrium, which is reasonably accurate if the carbonconversion in the gasifier is properly considered and the gasifieroperating temperature is known.37,38 Carbon conversion is oneof the most important parameters for measuring the gasifierperformance and depnds on a number of factors.39-42 In thisstudy, the carbon conversion in the gasifier is specified to be98% based on the TECO results for Illinois No. 6 #1 coal.36

The restricted-equilibrium gasifier reactor model (RGibbsmodel in Aspen Plus) used in this work minimizes Gibbs freeenergy considering the reactions shown in Table 3. The reactiontemperature approaches shown in Table 3 are specified to matchthe experimental/industrial data for the GEE-type entrained-flow gasifier using Illinois No. 6 coal and operating at atemperature of 1315.6 °C and pressure of 5.61 MPa.1 As someof the reactions, such as the WGS reaction, continue in the RSCand the current RSC model does not consider any reaction, thetemperature approach in the gasifier has been utilized to capturethis phenomenon. Since the GEE-type gasifier is a slagginggasifier, its operating temperature must be high enough to ensurefree slag flow along the wall and avoid clogging at the bottomof the gasifier. Wide variations in the ash fusion temperature

of various types of Illinois No. 6 coal have been reported.43

The data from the test runs conducted at TECO IGCC plantshow that the T250 temperature of the ash from the two types ofthe Illinois No. 6 coal are 1360 and 1326 °C, respectively.44

Therefore, the operating temperature of 1315.6 °C is reasonablefor the desired slag viscosity and carbon conversion. It shouldbe noted that higher operating temperatures can reduce therefractory life for GEE-type gasifiers.

In this study, the gasifier dimensions have been calculatedmaintaining the same superficial velocity as that reported forthe TECO gasifier.44,45 For calculating the heat loss from thegasifier, both the radiative and convective heat losses have beenconsidered. An infrared image of a GEE-type gasifier showsskin temperature of about 175-200 °C.46 A uniform skintemperature of 200 °C is assumed in this study. In addition, awind velocity of 8 km hr-1 and ambient temperature of 25 °Cis assumed. The heat loss from the gasifier is found to be about0.65% of the HHV of the coal. For the simulations in this work,the O2/wet coal ratio is manipulated so that the gasifiertemperature is maintained constant at 1315.6 °C while takinginto account the heat loss at the conditions mentioned previously.

For steady-state gasifier models in Aspen Plus, coal can bedeclared as a nonconventional solid component characterizedin terms of the component attributes ULTANAL (ultimateanalysis) and PROXANAL (proximate analysis). However, itshould be noted that unit operation models involving solids inAspen Plus Version 2006.5 cannot be exported automaticallyto Aspen Plus Dynamics. Even though it is possible to modelunit operations involving solids in Aspen Plus, it is not yetpossible to export that model to Aspen e outlet from the firstreactor is cooPlus Dynamics. In view of this limitation, Robinsonand Luyben45 have used a complex cyclical aromatic as asurrogate compound for coal. In the future, the plant-widesteady-state IGCC model developed as part of this study willbe exported to Aspen Plus Dynamics for use in transient studies.In view of this, all the nonconventional components are declaredas conventional components using electrolyte chemistry. Sincesalt is the only solid component currently handled in AspenPlus Dynamics, the nonconventional solids in the Aspen PlusIGCC model developed in this study are treated as salts bysetting up chemistry of type “salt”. The equilibrium constantsare manipulated such that the equilibrium completely favors thesalt. For the gasification section, the selected property methodis the “ELECNRTL” option which uses the Redlich-Kwongequation of state. In summary, the electrolyte chemistry ap-

Table 1. Proximate Analysis of Illinois No. 6 Coal (Dry Basis)

value (dry basis)

moisture 11.12FC 49.72VM 39.37ash 10.91HHV (kJ kg-1) 30506

Table 2. Ultimate Analysis of Illinois No. 6 Coal (Dry Basis)

value (dry basis)

ash 10.91carbon 71.72hydrogen 5.06nitrogen 1.74sulfur 2.82oxygen 7.75

Table 3. Restricted Equilibrium in the Gasifier

no. reaction

temperatureapproach to

equilibrium (°C)

1 C + O2 f CO2 0

2 C + 0.5O2 f CO 0

3 H2 + S f H2S 0

4 H2O + CO f CO2 + H2 -480

5 CH4 + H2O f 3H2 + CO -50

6 N2 + 3H2 f 2NH3 -700

7 COS + H2O f CO2 + H2S -650


proach not only generates comparable results to the case wherenonconventional solid components are used but also enablesthe automatic export of the plant-wide IGCC model from AspenPlus to Aspen Plus Dynamics.

Radiant Syngas Cooler (RSC) and the Scrubber. As shownin Figure 3, the gasifier outlet goes to the RSC where it is usedto generate HP steam from the boiler feedwater (BFW) comingfrom the HRSG economizer at a pressure of about 13.58 MPa.The outlet temperature of the syngas leaving the RSC is fixedat 599 °C. The saturated steam from the RSC steam drum issent to the HRSG superheater. The heat loss from the RSC iscalculated by considering both radiative and convective heatlosses. Slag is separated from the syngas in a “flash” separator.The syngas from the top of the slag separator goes to thescrubber. The scrubber is simulated as a flash vessel where theamount of quench water is manipulated by a design specification(called “designspec” in Aspen Plus) to decrease its temperatureto 210 °C.

Shift Converters and Syngas Cooling. The two water-gasshift (WGS) reactors in series are modeled as adiabatic plug flowreactors. The WGS reaction is a reversible reaction and is givenby

The kinetics are taken from the open literature for a sour shiftcatalyst.47,48 The forward rate for “Catalyst # N” which is a “half-strength” cobalt molybdenum-based catalyst is given by eq 2.

where Ef ) 53172 (kJ)/(kmol) and [CO] is molar concentration ofCO in (kmol)/(m3s).

The rate parameters in eq 2 were derived under near-atmospheric condition.47,48 The authors could not find any rateexpression for the sour shift catalysts in the open literature thathas been validated with experimental data at the operatingpressure considered in this study. The equilibrium constants aregiven49 by eqs 3 and 4 for the high temperature reactor (thefirst reactor) and the low temperature reactor (the secondreactor), respectively.

where T ) temperature, °R.The sizing of the reactors, the quantity of the catalysts, and

the pressure drop across the reactors have been determinedfollowing the work of Rase.49 It was ensured that the operatingtemperatures of the reactors remain in the range of 215-500°C.47,48,50 In addition, the minimum temperature at the inlet ismaintained at 25 °C above the dew point.

The outlet from the first reactor is cooled to 355 °C by raisinga part of the shift steam. The syngas is further cooled to 241 °Cby generating IP steam before it is sent to the second reactor. Thesecond reactor outlet is cooled by raising IP steam that is reheatedalong with the exhaust from the HP steam turbine and used in theIP section of the steam turbine. Shifted syngas is further cooled infour exchangers with intermediate flash drums before going to theSELEXOL unit. In the first exchanger, it is cooled to 168 °C bygenerating IP steam. After that, syngas is cooled to 122 °C byheating up the BFW. In the third exchanger, syngas is used to heatthe makeup water for the scrubber. Finally, syngas is cooled bythe cleaned syngas leaving the CO2 absorber in the SELEXOLunit. The process condensate streams from the first through lastflash vessels, containing about 798, 1402, 3821, and 25949 ppmof NH3, respectively, are sent to the SWS.

Sour water Stripper (SWS). Sour water stripping has beena subject of study for a long time.51-53 However, littleinformation is available in the open literature about the impactof the design of a sour water stripper (SWS) in an IGCC plant.The SWS plays a key role in generating water that can berecycled back to the gasifier and is a major consumer of strippingsteam. Operation of the SWS is also one of the problematicareas in an IGCC plant.44

In this simulation, approximately 5% of the bottom effluentfrom the scrubber is sent to the SWS as purge. The remainderof the bottom effluent is sent to a flash block. The vapor fromthe flash block is sent to the SWS. The liquid from the flashblock is recycled back to the scrubber. As mentioned above,the process condensate from the syngas coolers is also sent tothe SWS. The condensate from the tail gas (TG) compressor

Figure 3. Configuration of the gasifier, syngas cooling, shift reactors, and the blackwater treatment units.

CO + H2O S CO2 + H2 (1)

-rf ) 2.6 × 104exp(- Ef

RT)[CO]kmol

m3s(2)

Keq ) exp(-4.33 + 8240T ) for1060 e T e 1360 (3)

Keq ) exp(-4.72 + 8640T ) for760 e T e 1060 (4)


suction knockout (KO) drum and TG interstage condensate isdrained to the SWS. In Aspen Plus, a rigorous distillationcolumn (RadFrac) is used to simulate the SWS. The followingreactions are considered

The electrolyte nonrandom two liquid (NRTL) activitycoefficient model is used for liquid phase physical propertycalculation. The Soave-Redlich-Kwong (SRK) equation of state(EOS) is used for the vapor phase.

SELEXOL Unit and CO2 Compression. Different configu-rations of the SELEXOL unit can be considered on the basis ofthe design objective and the available integration opportunities.54

In this work, the dual-stage SELEXOL unit, as shown in Figure4, is configured for selective removal of H2S (first stage) andCO2 (second stage) from the sour syngas using dimethyletherof polyethylene glycol (DEPG) as a solvent. Most of the H2Sin the syngas is absorbed in the semilean solvent as it passesthrough the H2S absorber. The tail gas (TG) from the Clausunit (Figure 5) is also recycled to the H2S absorber. Thecondensate from the TG compressor interstage KO drums issent to the SWS. The off-gas from the top of the H2S absorberis sent to the CO2 absorber. Clean syngas from the CO2 absorber

exchanges heat with the incoming sour syngas to the H2Sabsorber and is then heated before going to the expander (Figure6).

A portion of the loaded solvent (about 30% in the base case)from the bottom of the CO2 absorber is cooled in a water cooler,chilled, and sent to the H2S absorber. The rich solvent from thebottom of the H2S absorber is heated by exchanging heat withthe lean solvent from the stripper. Thereafter, the syngas goesto a flash vessel. The vapor from the flash vessel is recycledback to the H2S absorber. The bottom stream from the flashvessel goes to the SELEXOL stripper which uses a combinationof stripping steam and a reboiler for deep removal of H2S fromthe syngas. Make-up solvent is mixed with the stripped solventand sent to the top tray of the CO2 absorber.

The remaining portion of the loaded solvent from the bottomof the CO2 absorber is heated and sent through a series of fourflash vessels to recover CO2 for compression in preparation forstorage. The loaded solvent stream is first heated by exchangingheat with the recycled solvent from the exit of the LP flashvessel and then further heated by a small LP steam heater. Next,the loaded solvent flows through the first flash vessel, generallycalled the H2 recovery drum, which operates at about 3/4 ofthe pressure of the CO2 absorber in order to recover about 75%of the H2 dissolved in the solvent. More specifically, the pressureof the H2 recovery drum is adjusted such that the sequesteredCO2 contains about 0.7 mol % of H2. This H2 concentrationlevel is easily achieved using the IGCC configuration andsimulation in this study. The recommended design basis for CO2

sequestration gas from NETL (Table 4) shows the limit on H2

content to be uncertain at this time. On the basis of futurespecifications, the H2 concentration can be increased or de-creased by adjusting the pressure of the H2 recovery drum. Uponremoving the absorbed H2, the solvent stream then goes throughthree additional flash vessels, high pressure (HP), mediumpressure (MP), and low pressure (LP), to release CO2. Ap-proximately 49.3%, 19.7%, and 25.3% of CO2 dissolved in thesolvent is recovered in the HP, MP, and LP flash vessels,respectively, in the base case. The semilean solvent leaving the

Figure 4. Configuration of the SELEXOL unit and the CO2 compression section.

2H2O S H3O+ + OH- (5a)

H2S + H2O S H3O+ + HS- (5b)

HS- + H2O S H3O+ + S-- (5c)

CO2 + 2H2O S H3O+ + HCO3

- (5d)

HCO3- + H2O S H3O

+ + CO3-- (5e)

NH3 + H2O S NH4+ + OH- (5f)

NH3 + HCO3- S NH2CO2

- + H2O (5g)


LP flash vessel is cooled by exchanging heat with the loadedsolvent and then chilled before returning to the fourth tray ofthe CO2 absorber. A NH3 vapor-compression cycle is consideredfor refrigeration. The flow rate of the refrigerant through thefour-stage intercooled compressor is determined by a designspecification considering a minimum temperature approach of5.5 °C. It should be noted that all the columns in this sectionhave been modeled considering equilibrium stage modeling thatmay overestimate the extent of absorption.

CO2 is compressed by a split-shaft multistage compressor.Three intercooled stages are considered for compressing the LPCO2 to the pressure of the MP CO2. MP CO2 is furthercompressed to the pressure of the HP CO2 followed by a twostage HP CO2 compressor. The compressed CO2 at a pressure

of 3.8 MPa is cooled and flashed to remove water (not shownin Figure 4). This stage is able to remove about 90% water.The compressed CO2 is further treated in an absorber usingtriethylene glycol (TEG) as the solvent (not shown in Figure4). Schwartzentruber-Renon EOS is used to simulate thisdehydration step. Because of a very high hygroscopicity, a 10stage tower is able to achieve a water content less than 0.01vol % that is well below the NETL target (Table 4). It isassumed that TEG is available at a concentration of 98.5 mol% from a regeneration process using a simple atmosphericpressure still.31 The dehydration process is performed at apressure much less than the critical pressure of CO2 in order toavoid loss of glycol in the supercritical CO2.

31 The dehydratedsequestration stream is further compressed to a pressure of 15.3MPa for transfer to the CO2 pipeline.

The target limits provided in Table 4 are tentative andexpected to change in the future. The impurities in thesequestration gas are limited mainly due to their influence onthe sequestration-gas compression and dehumidification system,the piping system, or the geological storage site.55 For the NETLtechnical note which is the basis of Table 4, existing industrialexperiences on sequestration at various storage sites have beenconsidered.55 Concentration of H2S in the sequestration gas is

Figure 5. Configuration of the Claus unit.

Figure 6. Configuration of the GT and the HRSG (flue gas side).

Table 4. Recommended Design Basis for the CO2-Sequestration Gasfor a Remote, Deep, Geological Storage Site55

final pressure 15.16 MPa

H2O 0.015 vol %N2, NH3, CO not limitedH2S <1.5 vol %CH4 <0.8 vol %H2 uncertainSO2 <3 vol %


limited primarily because of safety concerns due to potentialpipeline leaks. According to an IEA study, IGCC cosequestrationgases containing up to 3% H2S are suitable for all forms ofsequestration with the main concerns being the toxicity of H2Sunder leak cases.56 The study also concludes that corrosion isnot an issue if the sequestration gas is sufficiently dry andsuitable pipeline materials are selected. In the current simulation,the required level of drying was achieved in the Glycol absorber.

It should be mentioned that additional CO2 can be strippedfrom the SELEXOL solvent by decreasing the pressures of theflash vessels, especially that of the LP flash vessel; however,this will increase the compression power required for CO2

sequestration. An optimization study is presented later in thispaper that shows how the pressures of the flash vessels aredetermined to minimize the overall auxiliary power requirementin the SELEXOL unit and the CO2 compression unit consideredtogether.

In the design of the SELEXOL unit, the environmental limitsare set on the basis of the emission limit of SO2 to theatmosphere and the target for CO2 capture. The operationallimits are due to the maximum operating temperature ofDEPG (<175 °C32), maximum water content in DEPG (typically<6 wt %32), and lower limit of the H2S concentration in thefeed to the Claus plant (>20 mol % H2S

31).

In this study, DEPG is represented by an Aspen Plus databankcomponent with an average molecular weight of 280. Theperturbed chain statistical associating fluid theory (PC-SAFT)EOS based on the SAFT is used to accurately represent thethermo-physical and transport property of the SELEXOLsystem.57 The DEPG vapor pressure, liquid density, heatcapacity, viscosity, and thermal conductivity of the solvent havebeen regressed in Aspen Plus using published data.58 Availabledata in the open literature on vapor-liquid equilibrium betweenthe DEPG solvent and the selected species have been used toadjust the binary interaction parameters.59 Results from thesimulation agree well with the data available in the existingliterature. However, the authors acknowledge that the accuracyof the model can be further improved if more validation dataare available, especially at high operating pressure such as thatconsidered in this study.

Claus Unit. The multistep, oxygen-blown Claus unit is a gasdesulfurizing process that recovers elemental sulfur from theacid gas stream generated from each gasifier train and the sourgas produced from the SWS. All the reactors in this section aremodeled on the basis of derived results from a commercialsoftware package (TSweet by BR and E). As shown in Figure5, the Claus unit is composed of one thermal stage and twocatalytic stages. In the thermal stage, the H2S-laden gas andthe oxidant react in substoichiometric combustion in the Clausfurnace. A split-flow oxy-fired configuration of the Claus furnaceis used to handle the dilute acid gas feed coming from the topof the SELEXOL stripper, while the off-gas from the SWS isfed directly to the furnace for the destruction of NH3. Formaintaining the furnace temperature, the feed streams arepreheated to a temperature of about 230 °C. The flow of theoxidant stream to the furnace is manipulated by a designspecification to achieve the desired 2:1 ratio for H2S/SO2. Theflame zone of the Claus furnace is modeled using a restrictedequilibrium reactor (RGibbs model in Aspen Plus) where thefollowing reactions are considered

The flame section temperature is maintained at 1315 °C bycontrolling the split fraction of the acid gas from the SELEXOLstripper with a design specification. The furnace bypass connectsto the anoxic zone which is modeled using a stoichiometricreactor (RStoic model in Aspen Plus). The following reactionsare considered in the anoxic zone including the newly formedSO2 reacting with the remaining H2S to form elemental sulfur:

The outlet from the anoxic zone enters the waste heat boiler(WHB) where IP steam is generated and the following restrictedequilibrium reactions take place

The sulfur in the vapor phase exists as S2, S6, and S8

molecular species, with S2 being predominant at higher tem-peratures and S8 predominant at lower temperatures.31 Theeffluent is further cooled in Condenser 1 by generating LP steam.Reactions 8a and 8b are considered in Condenser 1. Liquidsulfur, as a mixture of S2, S6, and S8, is separated in Separator1. The vapor from Separator 1 is heated to 242 °C by IP steam.Gas preheating is required prior to entering the catalytic reactorsto avoid sulfur condensing in the catalyst bed.

In the first of two catalytic stages, the heated reactants flowover a fixed bed of activated alumina catalyst where sulfur isformed via the Claus reaction. Catalytic Stage 1 is modeledusing a stoichiometric reactor (RStoic model in Aspen Plus)with the following reactions:

The effluent is cooled in Condenser 2 by generating LP steam.Reactions 8a and 8b are again considered in Condenser 2. Liquidsulfur is separated in Separator 2. The sequence of preheater,catalytic reactor, condenser, and separator is repeated again forfurther conversion to liquid sulfur. After the second catalyticstage, the tail gas outlet from Separator 3 is preheated to atemperature of 230 °C and sent to a hydrogenation reactor forfurther processing. In the hydrogenation reactor (RGibbs model2H2S + 3O2 f 2SO2 + 2H2O (6a)

2H2S + O2 f S2 + 2H2O (6b)

4NH3 + 3O2 f 2N2 + 6H2O (6c)

CH4 + 2O2 f 2H2O + CO2 (6d)

2H2S S S2 + 2H2 (6e)

CO + H2O S CO2 + H2 (6f)

COS + H2O S CO2 + H2S (6g)

4H2S + 2SO2 S 3S2 + 4H2O (7a)

3S2 S S6 (7b)

COS + H2O S CO2 + H2S (7c)

2NH3 S N2 + 3H2 (7d)

4H2 + 2SO2 S S2 + 4H2O (7e)

3S2 S S6 (8a)

4S2 S S8 (8b)

4H2S + 2SO2 f S6 + 4H2O (9a)

16H2S + 8SO2 f 3S8 + 16H2O (9b)

COS + H2O f H2S + CO2 (9c)


in Aspen Plus), the following restricted equilibrium reactionsare considered

The effluent from the hydrogenation reactor goes to the tailgas (TG) knockout drum. The liquid from the knockout drumgoes to the SWS. The vapor is compressed and sent to theSELEXOL unit.

Gas Turbine (GT) and the Heat Recovery Steam Gene-rator (Flue Gas Side). The gas turbine (GT) converts thechemical energy in the supplied hydrogen-rich syngas fuel intoshaft work which turns a generator and produces electricity. Inthis study, the 3-stage GT section is modeled with the turbinemodels available in the Aspen Plus library and simulated onthe basis of data available in the open literature for the GEE7FB turbine.60,61 As shown in Figure 6, the clean syngas fromthe expander is sent to the GT combustor after being dilutedwith N2 from the air separation unit. The amount of N2 dilutionis manipulated by a design specification so that the lower heatingvalue (LHV) of the syngas fuel is reduced to 4.55 MJ Nm-3. Ithas been shown in the literature that this dilution is able toachieve a NOx concentration of 25 ppmv in the exhaust gas.62

The combustion air is compressed in an axial flow compressorwhich raises the ambient air to a pressure of 1.65 MPa. Whenthe flow of combustion air is manipulated, the GT combustortemperature is maintained at 1377 °C with a specified heat lossequal to 1.5% of the lower heating value (LHV) of the syngas.The GT firing temperature is maintained at 1327 °C by a designspecification which manipulates the air flow rate to the combus-tor outlet gas before it reaches the first expansion stage. Theair flow rates to the second and third expansion stages aremaintained at predetermined values. The exhaust temperature

is maintained at 566 °C by manipulating the isentropic efficien-cies of the GT. The isentropic efficiencies of all the three stagesare assumed to be equal. The flue gas goes to the heat recoverysteam generator (HRSG) where steam is generated at threepressure levels. The flue gas is used to superheat the HP steamgenerated both in the HRSG evaporator and in the radiant syngascooler. The flue gas also generates steam that is used in theeconomizer section and for BFW heating. The flue gas finallyexits the system at 132 °C, well above the cold end corrosiontemperature.

Steam Cycle. As shown in Figure 7, the triple pressure steamcycle generates steam from the flue gas and other processstreams. The minimum temperature approach is considered tobe 10 °C in this study. HP steam, generated at 12.4 MPa and538 °C, is mainly used for generating power in the HP steamturbine (ST). IP steam is used for generating power, as well asin the reboilers. LP steam generated in the HRSG is mainlyused for heating process streams and in the reboilers. Condensateat the outlet of the surface condenser is mixed with the makeupdemineralized water and sent to the deaerator, via a series ofheat exchangers, and the HRSG. Condensate from the LP steamcircuit and flash steam from the HP blow-down drum are alsomixed in the deaerator. BFW at the outlet of the deaerator getssplit into three streams and is pumped at various pressure levelsfor generating HP, IP, and LP steam, respectively. BFW fromthe HP BFW pump goes to the economizer where it gets heatedto 313 °C before being split into two streams. A portion goesto the RSC, and the remainder goes to the HP steam evaporatorin the HRSG. HP steam from both sources is combined andsent to the superheater before being sent to the HP turbine. IPBFW passes through the economizer and evaporator to generateIP steam. A small portion of the IP steam is sent to the LPsteam header to satisfy the LP steam demand. The remainingIP steam is sent to the IP turbine. The third split from thedeaerator gets divided mainly into two streams. One stream isused for generating shift steam from the WGS reactor inter-cooler. The other stream is used mainly for generating LP andIP steam. The exit temperature of the flue gas above the coldend corrosion temperature is maintained by manipulating theflow of the BFW that goes to the LP steam evaporator. IP steam

Figure 7. Schematic of the steam cycle.

6H2 + S6 S 6H2S (10a)

8H2 + S8 S 8H2S (10b)

3H2 + SO2 S H2S + 2H2O (10c)

COS + H2O S CO2 + H2S (10d)

CO + H2O S CO2 + H2 (10e)


at the exit of the HP turbine is mixed with the IP steam generatedfrom the effluent of the first stage and second stage WGSreactors and sent to the HRSG for reheating. The reheated IPsteam enters the IP turbine at 538 °C and 2.8 MPa. LP steamfrom the exhaust of the IP turbine flows through a crossoverpipe where additional LP steam generated in the HRSG, andelsewhere in the process, is mixed. The exhaust steam fromthe LP turbine enters the surface condenser at 6.77 kPa.

Air Separation Unit (ASU). The elevated-pressure airseparation unit (EP ASU) is designed to provide a 95 mol %O2 stream to the gasifier along with a high-pressure ultrapureN2 stream extracted from the high-pressure (HP) column and aN2 stream of more than 98% purity from the low-pressure (LP)column. A four-stage intercooled compressor is considered asthe main air compressor (MAC). The compressed air at apressure of 1.31 MPa enters the ASU unit for cooling andremoving the impurities before being sent to the distillationcolumns. It is assumed in this study that the ASU is able toprovide the product streams at the required specifications underall operating conditions. The major portion of the O2 goes tothe gasifier through a four-stage intercooled compressor. A smallstream of O2 goes to the oxy-Claus sulfur recovery unit. A partof the LP N2 that is required for dilution of the GT fuel iscompressed by a four-stage compressor and mixed with theboosted ultra pure N2 from the HP column. The mixed N2 isheated up before being sent to the GT combustor. The excessN2 is vented from the system.

Validation of the Gasifier Model and Design of theShift Reactors. The gasifier is the heart of an IGCC plant.Therefore, the gasifier model is validated with the availableindustrial data. Even though the TECO gasifier was run usingIllinois No. 6 # 1 coal for 20 days in 1997 and for 7 days in1998, no information about the data collected during the testrun could be found in the open literature. However, data arereported for Patriot Kentucky no. 9 coal that was processed in1999.44 This coal is similar in composition to Illinois no. 6 coalconsidered in this study. A comparison of the ultimate analysisof the coals is shown in Table 5. The table shows that ash andoxygen contents of these two coals are different to some extent.The test run data are reported at the exit of the gas cleanupunit. The configuration of the gas cleanup section is quitedifferent in this study compared to that in the TECO plant.Besides, shift converters have been considered in this study.Therefore, it makes sense to compare the syngas compositionat the gasifier exit. The TECO test data show the compositionand flow of the clean syngas as well as acid gases. For thepurpose of comparison with our simulation results, these twostreams are added together. The comparison is done on a dry,ash-free basis. The comparison is presented in Table 6.

It can be seen that the simulation results match well with theindustrial data. In Table 6, for calculating the value of the H2O/wet coal and O2/wet coal ratios in the TECO data, both the wetcoal and the mass of the solids in the recycle water have beenconsidered. In calculating the value of the O2/wet coal ratio,the net mass of O2 in the oxidizing stream has been considered

because the oxidizing stream in the TECO data has 95.9 vol %O2 compared to 95.0 vol % in the present simulation.

As the WGS reactors play a key role in carbon capture, theirdesign criteria and results are presented here. Figure 8 showsthe temperature profile in the WGS reactors. In the case of 96%overall conversion, about 82% conversion takes place in thefirst reactor. This results in a rise of 184 °C temperature acrossthe catalyst bed compared to about 32 °C across the secondreactor. About 15% overdesign in the length is used in view ofanticipated changes in the feed flow rate and composition.

Results and Discussions

The following studies are on optimum global design decisions,local design decisions, operating conditions, and operationalflexibility. In these studies, the product specifications from theSELEXOL unit are satisfied under all operating conditions bythree design specifications. The first design specification main-tains the desired percent capture of CO2 in the SELEXOL unitby manipulating the flow rate of the circulating semilean solventthrough the flash vessels. The second design specificationsatisfies the desired H2S slip in the clean syngas by manipulatingthe flow rate of the lean solvent from the H2S stripper. Thethird design specification maintains the required concentrationof H2S in the Claus feed by manipulating the pressure of theH2S concentrator. The authors would like to note that all theresults presented below are generated by executing each phaseonly once and, therefore, should not be considered as the finalIGCC design as shown in Figure 1.

Global Design Decision: Optimum Combination of COConversion in the WGS Reactors and CO2 Capture inthe SELEXOL Unit. Identification of global design decisions(at the plant level) can be done by considering interactionbetween different units and its impact on the plant efficiency.An optimum decision can help to improve the overall plant

Table 5. Comparison of Ultimate Analysis of Illinois No. 6 Coal andKentucky No. 9 Coal (Dry Basis)

Illinois no. 6 coal Kentucky no. 9 coal

ash 10.91 9.63carbon 71.72 71.78hydrogen 5.06 4.84nitrogen 1.74 1.56sulfur 2.82 3.04oxygen 7.75 9.15

Table 6. Comparison of the Simulation Results with the Data fromTECO Test Run

present study TECO data

CO (vol %) 40.3 41.0H2 (vol %) 39.2 36.1CH4 (vol %) 0.11 0.07CO2 17.3 17.5H2S + COS 0.87 0.91H2O/wet coal 0.4108 0.3956O2/wet coal 0.7822 0.7744gasifier exit temperature (°C) 1315.6

Figure 8. Temperature profile in the WGS reactors.


efficiency by reducing losses (such as auxiliary power consump-tion) and/or increasing the power generation. In this study, anoptimal decision is taken for achieving 90% carbon capture,which corresponds to the design basis for Case 2.1 The twomain decision variables are the extent of CO conversion in theWGS reactors and the extent of CO2 capture in the SELEXOLunit. Table 7 presents results from this study. As the percentconversion of CO in the WGS reactors is increased from 86.5%to 87.5% and the percent capture of CO2 is decreased from98.9% to 98.3%, the net power increases by about 0.4%. Thishappens because of a significant decrease in the powerconsumption in the SELEXOL unit. For such a deep recoveryof CO2 in the SELEXOL unit, significantly higher circulationrates of the solvent and associated increase in power consump-tion for refrigeration and enhanced stripping are required. Thisalso adds to the refrigeration load and cost. It should bementioned that a physical solvent such as DEPG is more suitablefor bulk removal of CO2 under conditions of high partialpressure of CO2. For such a deep recovery of CO2, the syngasbecomes lean in CO2 in the upper section of the CO2 absorber.Absorption under this condition can be enhanced by increasingthe solvent/gas ratio, lowering the temperature of the solvent,increasing the extent of stripping of the solvent in the flashvessels, and increasing the number of trays in the column (notconsidered here). An increase in the solvent/gas ratio results inhigher pumping and refrigeration costs. Lowering the temper-ature of the solvent results in a higher refrigeration cost.Enhanced stripping can be achieved by lowering the operatingpressure of the flash vessels particularly that of the LP vesseland by increasing the solvent temperature. A lower operatingpressure results in higher compression costs in the CO2

compressors. An increase in the solvent temperature consumessteam and adds to the refrigeration cost. Because of the lowerpartial pressure of CO2 at this deep recovery, increasing thenumber of trays does not give much benefit. As percentconversion of CO in the WGS reactors increases, steamextraction from the ST increases resulting in lower powergeneration by the ST. In addition, with an increase in conversionof CO, the cold gas efficiency decreases resulting in lower powerfrom the GT. As the percent conversion of CO increases andpercent capture of CO2 decreases further to satisfy the overallgoal of 90% carbon capture, the net power produced decreases.It is seen that the power produced from the GT and the STdecrease monotonically. A small increase in the power require-ment of the SELEXOL unit is found even though the flow ofthe circulating solvent decreases because of lower CO2 capture.This happens because of an increase in partial pressure of CO2

at the inlet of the H2S absorber. This in turn results in higherCO2 content in the H2S absorber bottom. To maintain the desiredconcentration of H2S in the Claus feed, the additional CO2 mustbe stripped off in the H2S concentrator. Therefore, the operatingpressure of the H2S concentrator is lowered resulting in a higherpower requirement of the stripped gas compressor. As thepercent conversion of CO in the WGS reactors is increased

beyond 95%, a sharp increase in the shift steam is observed,resulting in a significant fall in the power generated by the ST.This results in a sharp fall in the net power produced. This studyshows that the optimum combination is about 87.5% conversionof CO and 98.3% capture of CO2 for achieving 90% C capturecompared to about 96% conversion of CO and 92.7% captureof CO2 in Case 2. This results in an increase in net plantefficiency (HHV basis) to 33.8% compared to 32.5% in Case2.1 The authors acknowledge that the H2O/CO ratio at the inletof the WGS reactors cannot be lowered than that suggested bythe catalyst manufacturer to avoid coke formation. One possibleapproach to operate the WGS reactors at a lower overall H2O/CO ratio is to bypass a part of the syngas, thus raising the H2O/CO ratio at the inlet of the reactors. However, the design of theSELEXOL unit should be evaluated for capturing the uncon-verted COS. For more accurate results from this study, rate-based models should be used for designing the towers insteadof the equilibrium model used in this study.

Optimum Local Design Decisions. Various options for localdesign (at the unit level) can be identified by consideringinteraction between different sections and their impact on theplant efficiency along with consideration of design constraints(such as operating and environmental limits). Here, two studiesare presented on local design decisions of the SELEXOL andthe Claus unit.

The SELEXOL unit should be designed not only forachieving the desired extent of capture of CO2 but also for atarget concentration of H2S in the Claus feed. The H2Sconcentrator plays a key role in this. In the first design phase,a multistage stripper was used as the H2S concentrator. As analternative, a single stage flash vessel could be used. The firststudy presented here compares these two configurations fromthe point of view of auxiliary power requirement. The secondstudy focuses on another local design decision by consideringthe routing of the tail gas. For a simultaneous optimizationapproach, these three studies need algorithms for solving mixedinteger programming problems.

Single Stage Flash Vessel vs a Multistage Stripper asthe H2S Concentrator. In this study, N2 from the ASU isconsidered to be the stripping gas. The number of theoreticalstages in the stripper is 6. As mentioned before, the operatingpressure of the H2S concentrator is manipulated by a designspecification to maintain the H2S concentration in the Claus feed.In the case of a stripper, the operating pressure of the H2Sconcentrator is about 2.5 times higher than the case when aflash vessel is considered. This decreases the power consumptionin the stripped gas compressor by more than 5.5 MW. As aresult, the auxiliary power consumption in the SELEXOL unitdecreases by about 16% in the case of a stripper. The net powerdecreases by about 1.09% when a flash vessel is used. Theauthors recognize that, for making this study more realistic, thecapital cost should be considered. However, note that the numberof trays required in the multistage stripper is a few whereas thesavings in the auxiliary power consumption is significant.

Table 7. Various Combinations of CO Conversion in the WGS Reactors and CO2 Capture in the SELEXOL Unit

% conversion of CO in theWGS reactors

% capture of CO2

in the SELEXOL unit power from ST (MW)power consumption inSELEXOL unit (MW) power from GT (MW) net power (MW)

86.5 98.9 294.66 33.41 479.05 576.1887.0 98.6 294.12 32.16 478.62 577.0487.5 98.3 293.90 30.97 478.35 578.2887.9 97.9 293.54 30.98 478.00 577.4790.1 96.3 291.9 31.01 476.85 576.0693.0 94.4 287.4 31.06 474.98 570.3595.3 92.9 277.83 31.10 473.43 559.8396.7 92.1 251.19 31.12 472.55 532.95


Routing of the Tail Gas (TG). In the first design phase, theTG was recycled back to the H2S absorber in the SELEXOLunit. Even though the flow of the tail gas is about 1.5% of thesyngas that enters the H2S absorber on a mass basis, it cannotbe vented to the atmosphere because of stringent environmentalregulations as mentioned before. However, the TG containsabout 52 mol % CO2 and 36 mol % N2. When sent to theSELEXOL unit, the TG has to be compressed to the pressureof the H2S absorber and the CO2 (and H2S) present in the TGneeds to be recaptured. Therefore, a possible outlet for the TGis direct routing to the LP CO2 compressor along with the CO2

from the LP flash vessel in the SELEXOL unit. It can then becosequestered with the CO2 stream from the other flash vessels.A study was done to see whether the resultant sequestered CO2

was still within the technical specifications. It was also seenwhether this option was more efficient. The results of this studyare presented in Table 8. It is seen that the power consumptionin the SELEXOL unit is slightly lower when the TG is routedto the H2S absorber mainly because the syngas is rich in CO2.As expected, the power consumption by the CO2 compressorincreases slightly, thus exceeding the savings in power byeliminating the TG compressor. However, when the TG isrecycled to the H2S absorber, the fuel species such as H2, CO,and CH4 get recovered and, thus, generate more power in theGT. This results in an increase in the overall power generatedby the plant. The last few rows in Table 8 show that thecomposition of the sequestered CO2 for both cases remainswithin the limits of the technical specifications given in Table4.

Optimum Operating Conditions. Once global and localdesign decisions are taken, the operating conditions can beoptimized. The first study optimizes the flow rate of N2 as astripping medium to the H2S concentrator. The second optimiza-tion study focuses on the operating pressure of the HP, MP,and LP flash vessels in the SELEXOL unit.

Optimum Flow Rate of N2 as Stripping Medium. For IGCCplants, N2 from the ASU is often the preferred stripping mediumin the H2S concentrator of the SELEXOL plant. As the flowrate of N2 is increased, Figure 9 shows that the H2S in the Clausfeed decreases and the N2 in the Claus feed increases. It hasbeen mentioned before that the concentration of H2S in the Clausfeed is maintained by a design specification by manipulatingthe pressure of the H2S concentrator. As the flow rate of N2

increases, a lower amount of CO2 gets dissolved in the H2Sabsorber because of a decrease in the partial pressure of CO2.The partial pressure of CO2 in the H2S concentrator decreases

as well. Because of these effects, the pressure of the H2Sconcentrator is increased by the design specification to maintainthe H2S concentration in the Claus feed. As the N2 flow increasesfurther, a substantial quantity of N2 gets dissolved in the solventand the operating pressure of the H2S concentrator is decreasedby the design specification to maintain the same concentrationof H2S in the Claus feed. These effects get reflected in the powerconsumption of the SELEXOL unit as shown in Figure 10. Thepower consumption in the SELEXOL unit initially decreasesbecause of two reasons. First, as a lower amount of CO2 getsdissolved in the H2S absorber, the recycle flow rate decreases,resulting in a decrease in the power consumption in the recyclecompressor. Second, the pressure of the H2S concentratorincreases, resulting in further decrease in the power cosnsump-tion. With further increase in the N2 flow, power consumptionof the recycle compressor increases because of a decrease inthe pressure of the H2S concentrator by the design specification.As the N2 flow is increased, most of it goes to the GT, thusreducing the requirement of diluent N2 flow. The portion of N2

that goes to the Claus unit is also recycled to the SELEXOLunit through the TG. As a result, an insignificant increase inthe power consumption of the N2 compressor is observed dueto an increase in the N2 flow. A negligible decrease in the GTpower is observed because of the increase in the flow of N2.

Table 8. Recyling of the Tail Gas to the SELEXOL Unit vs DirectRouting of the Syngas to Sequestered CO2

tail gasrecycled to

the SELEXOL unittail gas routed

to sequestered CO2

power consumption in theSELEXOL unit (MW)

30.97 31.00

power consumption by theTG compressor (MW)

0.707 0

power consumption by theCO2 compressors (MW)

25.25 25.97

power generated by theGT (MW)

478.35 477.95

net power (MW) 578.28 577.64sequestered CO2 composition

(% mol)H2O 0.008 0.009CO2 98.827 98.508H2 0.684 0.704H2S 0.013 0.047NH3 0.024 0.020 Figure 9. Change in the flow rate of N2 and H2S in the Claus feed because

of change in the flow rate of N2 as the stripping gas.

Figure 10. Change in the power consumption in the SELEXOL unit andnet power generation because of change in the flow rate of N2 as the strippinggas.


As a result, the net power, as shown in Figure 10, is mainlyaffected by the power consumption in the SELEXOL unit. Thestudy shows that there is an optimum flow rate of the strippingN2 (which is about 2270 kmol/h for this plant) beyond whichthe net power from the IGCC plant decreases. This study resultsin an increase of about 1.32 MW in the net power in additionto what was achieved in the first design phase.

Operating Pressure of the Flash Vessels. The powerconsumption for compressing CO2 to the sequestration pressureis about 6% of the total power produced by the GT. Therefore,a reduction in this compression power can augment the totalpower generated in the plant. This can be achieved by operatingthe HP, MP, and LP flash vessels in the SELEXOL unit at anoptimum pressure. The following study to optimize the operatingpressures of the flash vessels was done by keeping the pressureof the HP vessel constant at 2620 kPa. The pressures of the LPand the MP flash vessels were varied. The pressure of the HPflash vessel was kept constant for simplicity and as the objectiveof this study was to find out whether an optimum combinationof the pressures of the LP and MP flash vessels exists for agiven pressure of the HP flash vessel. The compression ratio ineach stage of the LP, MP, and HP CO2 compressors is keptconstant in this study by adding/removing stages with inter-cooling as needed. As mentioned before, the product specifica-tions from the SELEXOL unit are maintained by three designspecifications. Therefore, the following discussion will focusonly on the SELEXOL and CO2 compression unit as the balanceof the plant is hardly affected by the pressures in the flashvessels. Figure 11 shows the total power consumption in theSELEXOL and CO2 compression unit as the pressures of theMP and LP flash vessels are changed. The best combination ofpressure as per this study is 1200 and 414 kPa for the MP andLP flash vessels, respectively, for a given HP flash vesselpressure of 2620 kPa. As the pressure of the LP flash vessel isincreased at a given pressure of the MP flash vessel, the powerconsumption decreases. With further increase in the pressureof the LP flash vessel, the power consumption increases mainlybecause of the increase in the refrigeration load and the pumpingpower for increased solvent flow. On the other hand, if thepressure of the MP flash vessel is increased at a given pressureof the LP flash vessel, the flow rate of the off-gas from the MPflash vessel decreases as its pressure is increased. This resultsin an increase in the off-gas from the LP flash vessel. This causesthe power consumption by the LP compressor to increase andthat by the MP compressor to decrease. As a result, the totalpower consumption by the CO2 compressors first decreases andthen keeps increasing. When the LP flash vessel pressure is heldconstant, there is hardly any change in the power consumptionin the SELEXOL unit. This study results in an increase in the

net power by about 3.52 MW in addition to what was achievedby optimizing the N2 flow rate in the previous study.

Assessment of the Operational Flexibility of the Plant.Assessment of the operational flexibility ensures that a desiredlevel of performance can be achieved in the face of disturbancesand uncertainties typical of a process plant. Therefore, thisassessment must be done before taking the final design decisions.The assessment strongly depends upon the bounds on theuncertain parameters/expected disturbances that one considers.It should be noted that this assessment is done with a steady-state model and, therefore, the effects of the transient responseof the process are not captured. Two studies on the operationalflexibility of the WGS reactors are presented below. The desiredlevel of performance for the WGS reactors is considered to bethe upper limit on the amount of supplemental steam extractionfrom the ST under disturbances and uncertainties. In addition,a study is presented that considers the decision for concentrationof NH3 in the process condensate from the SWS.

Effect of H2O/CO Ratio on CO Conversion in theWGS Reactors. In the first design phase, the optimum COconversion in the WGS reactors is found to be about 87.5%.However, in a flexible CO2 capture scenario, it may be desiredto alter the CO conversion in the WGS reactors. Figure 12 showsthe conversion of CO in the first WGS reactor and the overallconversion in this two stage design as the H2O/CO ratio ischanged. When the H2O/CO ratio is about 2, about 95% overallconversion of CO is achieved. The effect of further increase inthe H2O/CO ratio is small. If the WGS reactors are operated atlower H2O/CO ratio, due consideration of the coke formationproblem is needed as discussed before. One problem withoperating the WGS reactors at high pressure is that the shiftsteam generated in the interstage cooler is much less comparedto the extracted steam from the ST. Therefore, an increase inthe H2O/CO ratio can significantly bring down the powergenerated by the ST. This is seen in Figure 13. The study showsthat if a higher conversion (more than 95%) is envisaged becauseof environmental regulations or otherwise, then this two-stageconfiguration of the WGS reactors will not be able to satisfy ahigher limit on the amount of supplemental steam extractionfrom the ST. A possible option is to consider a third reactor tosatisfy this operational flexibility.

Effect of Change in C/H2 Ratio and CO2/CO Ratio inthe Syngas. The gasifiers have been successfully utilized toprocess biomass, petcoke, and different types of coals in various

Figure 11. Total power consumption in the SELEXOL and CO2 compres-sion unit with change in the pressure of the flash vessels.

Figure 12. Change in percent conversion of CO in the WGS reactors asthe H2O/CO ratio is changed.


proportions.44 With the CO2 capture option in place, the designneeds to consider the effect of changes in the fuel on the CO2

capture capabilities of the plant. To study the effect of changesin the feed and syngas composition, the C/H2 and CO2/CO ratiosof the syngas to the WGS were changed. The C/H2 ratio iscalculated by taking the ratio of the combined molar flow rateof CO and CO2 to that of H2. The CO2/CO ratio is calculatedon a molar basis. The C/H2 ratio can change widely on the basisof the fuel. The CO2/CO ratio in the syngas can vary for variousreasons such as the amount of oxidant supplied to the gasifier,gasifier temperature, type of gasifier, fuel type, etc. Consideringa flexible CO2 capture scenario, operational flexibility of thedesigned WGS reactors is analyzed for a CO conversion of95.4%. Note that this CO conversion is about 8% higher thanthe optimal design. The H2O/CO ratio is manipulated at the inletof the first reactor. For all the studies, the molar flow rate ofsyngas remained constant. Figure 14 shows that the flow of theshift steam keeps decreasing as the C/H2 ratio decreases. Asthe C/H2 ratio decreases, the extent of reaction decreases. Theamount of CO that enters the WGS reactors decreases as theC/H2 ratio decreases and the CO2/CO ratio increases. Becauseof the decreased heat release, the reactor temperature decreasesfavoring reaction toward the products. This helps to bring downthe shift steam required. However, as the CO2/CO ratio isincreased further, concentration of CO decreases considerably.This requires an increase in the concentration of H2O to achievethe desired conversion. It is seen that the flow of the shift steamkeeps increasing significantly as the CO2/CO ratio increasesbeyond 0.48 for a C/H2 ratio below 1.4. As the total heat released

decreases with a decreased C/H2 ratio and increased CO2/COratio, the steam generation in the interstage steam generatordecreases as seen in Figure 15. The sharper decrease in the steamgeneration is very significant beyond a CO2/CO ratio of 0.48and a C/H2 ratio below 1.4. This can be attributed to a higherflow rate of steam. The flow rate of the extracted steam variesaccordingly as seen in Figure 16. The combined effect causesthe steam extraction to increase significantly as the CO2/COratio becomes more than 0.48 and C/H2 ratio falls below 1.4. Ifan operational flexibility in the feed compositions and operatingconditions that can lead the syngas composition to this regionis desired, then the WGS plant should be designed accordinglyto avoid the violation of the upper limit on the steam extractionfrom the ST. This should be noted that the overdesign value ofthe shift reactors (which is 15% in this study) also affects thefeasible operating window under the conditions of changingsyngas composition.

NH3 Content in the Process Condensate from the SWS.The main purpose of the SWS is to remove NH3 and H2S fromthe process condensate so that the water can be reused and/ordisposed of as wastewater. However, the SWS consumesconsiderable stripping steam and, therefore, a design decisionneeds to be taken considering the steam consumption for a targetremoval of NH3 and H2S. The desired levels are typically below50 ppmw for NH3 and 10 ppmw for H2S,31 to prevent odor. Asthe stripped condensate is used in the RSC sump and thelockhopper, the presence of NH3 makes the wet slag unmarket-able because of the objectionable odor.44 Therefore, in the designof a SWS for an IGCC plant, this issue should be considered.Figure 17 shows the requirement of the steam to the SWSreboiler based on the desired concentration of NH3 and H2S inthe stripped condensate. The figure shows that if NH3 contentis maintained below 50 ppmw, the desired concentration of H2S

Figure 13. Extraction steam from the ST and power generated by the GTas the H2O/CO ratio is changed.

Figure 14. Flow of the shift steam as C/H2 and CO2/CO ratios are changed.

Figure 15. Production of steam in the interstage steam generator as C/H2

and CO2/CO ratios are changed.

Figure 16. Steam extraction from the GT as C/H2 and CO2/CO are changed.


can be maintained much below its desired value. However, onthe basis of the sharp change in the slope of the curve, it is notadvisible to keep a target concentration of NH3 above 30-40ppmw considering the operational flexibility of the SWS. Ifthe plant is designed for more than 30-40 ppmw NH3 in thestripped condensate, a small change in the steam flow to thereboiler due to some external disturbances can result in anundesired rise in the level of NH3 content in the strippedcondensate.

Conclusions

This paper presents a three-phase, top-down, optimization-based approach for designing an IGCC plant with CO2 capturein a commercial process simulator environment. In this ap-proach, first important global design decisions are made on thebasis of optimization studies. At the second design phase,optimum local design decisions are taken. After that, theoperating conditions are optimized. The operational flexibilityof the plant should be assessed before taking the final designdecisions. A number of examples is cited for each phase.

The results from the first design phase shows that, for 90%C capture, the optimum combination is about 87.5% conversionof CO in the WGS reactors and 98.3% capture of CO2 in theSELEXOL unit. This design decision results in about 1.3%increase in net plant efficiency (HHV basis) compared toCase 2.1

At the second design phase, two studies are done. The firststudy shows that a multistage stripper is considerably moreenergy efficient as the H2S concentrator compared to a singlestage flash vessel. The net power decreases by about 1.09%when a flash vessel is used. The second study shows that routingof the TG to the H2S absorber is more energy efficient comparedto routing directly to the sequestered CO2. It is observed that,if the TG is routed to the sequestered CO2, the composition ofthe sequestered CO2 still remains within the limits of thetechnical specifications. Therefore, this option can be consideredfor added flexibility and may be useful under certain circum-stances such as failure of the TG recycle compressor.

At the third phase, the operating conditions are optimized. Astudy is done to find out the optimum flow rate of the strippingmedium in the H2S concentrator. This study results in anadditional increase in the net power by about 1.32 MWcompared to the first phase. The operating pressures of the flashvessels are also optimized. Considering a constant pressure of

2620 kPa in the HP flash drum, the best combination of pressureis found to be 1200 and 414 kPa for the MP and the LP flashdrums, respectively. This results in an additional increase of3.52 MW in the net power. In this paper, each phase of theproposed top-down design approach is executed once. At theend of the first iteration, the achieved net plant efficiency (HHVbasis) is about 34.1% which is 1.6% higher than Case 2.1 Thenet power increases by 4.9% in comparison to Case 2.1

Before taking the final design decisions, operational flexibilityof the plant should be considered. Two studies are doneassessing the operational flexibility of the WGS reactors. Thefirst study shows that, beyond a H2O/CO ratio of 2, theadditional CO conversion is insignificant. If the overall conver-sion of CO in the WGS reactors is uncertain at this time andcan well exceed 95% in future, then the current two-stage WGSconfiguration will result in a significantly higher steam extractionfrom the ST. Other design options should be considered tosatisfy this operational flexibility. The second study evaluatesthe WGS design when the syngas composition changes.Considering a flexible CO2 capture scenario, the operationalflexibility analysis of the designed WGS reactors is done in awide range of CO2/CO ratios and C/H2 ratios by manipulatingthe H2O/CO ratio at the inlet of the first reactor. When the CO2/CO ratio is more than 0.48 and the C/H2 ratio is below 1.4, asharp rise in steam extraction from the ST is observed thatcauses a significant loss of power. Other design options shouldbe evaluated if operation in this region is required withoutviolating an upper limit on the steam extraction from the ST.A plot between the steam flow to the SWS reboiler and NH3

content in the stripped condensate from the SWS bottom showsa sharp change in the slope of the curve when the NH3 contentin the stripped condensate is above 30-40 ppmw. Consideringthe operational flexibility of the plant, a small disturbance inthe steam flow to the reboiler can cause an undesired risein the NH3 content in the stripped condensate.

Finally, the authors would like to reiterate that this papermainly focuses on maximizing the net power from an IGCCplant with CO2 capture. The specific results presented in thisstudy depend on a number of things such as the modelingassumptions (for example, equilibrium tray calculations in theSELEXOL absorbers), configurations of the plant, choice oftechnology, thermodynamic property methods and models,transport property methods, physical property parameters, etc.The authors strongly believe that separate validations of the unitlevel models and the plantwide model using experimental/industrial data would have helped to improve the modelpredictions. Unfortunately, no IGCC plant with CO2 captureexists today. Furthermore, experimental/industrial data are rarein the open literature for most of the unit operations consideredin this paper in the operating range of interest. Nevertheless,the paper presents an approach to optimizing complex IGCCprocesses with CO2 capture that possess large numbers ofinteractions and trade-offs. In addition, a number of key designvariables is identified for improving the efficiency of an IGCCplant with CO2 capture. It should be noted that the capital costmust be considered for an optimal techno-economic design. Toenhance the search space for optima, topological changes inthe plant and consideration of other available technologiesshould also be considered.

Acknowledgment

This technical effort was performed in support of the NationalEnergy Technology Laboratory’s ongoing research in Process

Figure 17. NH3 and H2S in SWS bottom product with change in the steamflow to the reboiler.


and Dynamic Systems Research under the RDS contract DE-AC26-04NT41817.

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ReceiVed for reView July 13, 2010ReVised manuscript receiVed November 30, 2010

Accepted December 2, 2010

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