Fuel 92 (2012) 94–106

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Fuel

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Simulations and process analysis of the carbonation–calcination reactionprocess with intermediate hydration

William Wang, Shwetha Ramkumar 1, Danny Wong 2, Liang-Shih Fan ⇑William G. Lowrie Department of Chemical and Biomolecular Engineering, 140 W 19th Avenue, 125 A Koffolt Labs, The Ohio State University, Columbus, OH 43210, United States

a r t i c l e i n f o

Article history:Received 2 February 2011Received in revised form 9 June 2011Accepted 29 June 2011Available online 21 July 2011

Keywords:Coal combustionCO2 captureSulfur captureCalcium sorbentCCR

0016-2361/$ - see front matter � 2011 Elsevier Ltd. Adoi:10.1016/j.fuel.2011.06.059

⇑ Corresponding author. Tel.: +1 614 688 3262; faxE-mail address: fan.1@osu.edu (L.-S. Fan).

1 Currently with ExxonMobil Company.2 Currently with Dow Chemical Company.

a b s t r a c t

Several technologies are currently being developed to separate carbon dioxide from large point sources,such as coal-fired power plants. An emerging technology that shows great potential is a calcium oxide–calcium carbonate cycle. A major drawback is the calcium carbonate decreases in reactivity over multiplecycles. The Ohio State University demonstrated in 2008 the first carbonation–calcination reaction (CCR)process that includes intermediate hydration for sorbent regeneration and its feasibility over multiplecycles at the 120 kWth scale with actual flue gas from coal combustion. The CCR Process utilizes a cal-cium-based sorbent to react with the carbon dioxide and sulfur dioxide in a flue gas stream to form cal-cium carbonate and calcium sulfate, respectively. The carbon dioxide is subsequently released from thecalcium carbonate to produce a high-purity, sequestration-ready carbon dioxide stream while regenerat-ing the calcium oxide sorbent. The sulfur dioxide is fixated as calcium sulfate and removed through apurge stream. An intermediate hydration step restores reactivity to the calcium oxide sorbent. Processanalysis from computer simulations shows the CCR Process to be highly effective and efficient in remov-ing both carbon dioxide and sulfur dioxide at low energy penalties under realistic conditions. A 20–22%decrease in electricity generation efficiency with the CCR Process is expected, compared with aminescrubbing around 27% and oxy-combustion around 25% energy penalty. A 25–28% increase in thermalenergy with the CCR Process is expected to maintain a constant electrical output. Further, the CCR Processconsumes half the oxygen necessary for an oxy-combustion plant and 25% less steam necessary for aminescrubbing.

� 2011 Elsevier Ltd. All rights reserved.

1. Introduction

The capture of carbon dioxide from large point sources, such ascoal-fired power plants, has become an increasingly importantarea of research. Both government and industry have placed signif-icant emphasis on the ability to remove the carbon dioxide that isemitted from existing coal-fired power plants. The flue gas condi-tions provide technological challenges for carbon dioxide capture.The flue gas exiting a power plant possesses a dilute concentrationof carbon dioxide, large volumetric flow, and ambient pressure. Forexample, a typical 500 MWe pulverized coal-fired (PC) power planthas a carbon dioxide (CO2) concentration between 10% and 15%volume and produces 1 million cubic feet per minute [1–3].

As a retrofit option, the main separation technologies focus onabsorption, adsorption, and oxycombustion [1,2,4]. Greenfieldplants with inherent carbon dioxide separation include coal gasifi-cation and chemical looping combustion [1,3,5,6]. Chemical

ll rights reserved.

: +1 614 292 3769.

absorption using an amine-based solvent has received the greatestamount of attention due to its mature technological development[1,2,4,7]. While an amine-based solvent can achieve the Depart-ment of Energy’s target of 90% carbon dioxide removal, the integra-tion of an amine scrubber into an existing PC power plant forcarbon dioxide removal is economically prohibitive and energeti-cally intensive [2]. Physical absorption requires elevated pressuresto effectively remove carbon dioxide and can utilize commercially-available solvents such as Selexol and Rectisol. While carbon diox-ide removal using physical absorption can be effectively integratedinto pressurized systems, such as coal gasification, integration intoa PC power plant provides no clear advantage over chemicalabsorption [3,4,8,9]. Adsorption onto the surface of a high-surfacearea solid is under development. Several types of adsorbents havethe ability to capture CO2. Currently, molecular sieves can retain246 g CO2/kg adsorbent. A bed of high-surface area adsorbent suchas alumina, zeolites, or activated carbon could also be used to ad-sorb the CO2 onto the sorbent [4,10]. Adsorption is a less developedtechnology and not yet suitable for power plant CO2 capture due toits poor capacity and selectivity [3,4,9]. The concept of oxycombus-tion replaces combustion air with a highly pure stream of oxygen.By doing so, the generated flue gas stream is minimally diluted

W. Wang et al. / Fuel 92 (2012) 94–106 95

with nitrogen and a concentrated stream of carbon dioxide is pro-duced [1–4]. The energy required to cryogencially separate oxygenfrom air, dilution of flue gas through air infiltration, and the eco-nomics of oxycombustion still remain challenges [2–4].

A relatively new idea for carbon dioxide removal from gasstreams is the reactive separation of carbon dioxide utilizing metalcarbonates. In a cyclical process, a metal carbonate-bearing min-eral is calcined and then exposed to the flue gas stream, wherethe calcined material, metal oxide, is re-carbonated. The processis then repeated. Since the metal carbonate is a solid, the calcina-tion reaction products are carbon dioxide, a gas, and the metaloxide, a solid, which can be separated using current gas–solid sep-aration equipment. The Toshiba Corporation has been developing acyclical process with lithium orthosilicate, Li4SiO4 [10–13]. Thecarbonation reaction occurs between 450 �C and 700 �C, and thecalcination reaction occurs at temperatures greater than 710 �C.The kinetics of lithium orthosilicate carbonation are on the orderof several minutes and requires further improvement [10,13]. Re-search Triangle Institute (RTI) has developed a cyclical process uti-lizing sodium carbonate, Na2CO3. Known as the Dry CarbonateProcess, the solids species cycle between sodium carbonate and so-dium bicarbonate. Occurring at lower temperatures than lithiumorthosilicate, the carbonation reaction occurs between 60 �C and80 �C, and calcination occurs at temperatures greater than 120 �C[7]. The Dry Carbonate Process has been tested for over 200 hand achieved greater than 90% carbon dioxide removal during test-ing with natural gas and coal derived flue gas [14]. The Dry Carbon-ate Process reacts with acid gases present in coal flue gases and isbest suited for power plants with wet Flue Gas Desulfurization forsulfur removal [3,7,14].

The most widely studied reactive sorbent for cyclical activity isbased on calcium carbonate, CaCO3. A cyclic calcium oxide/calciumcarbonate (CaO/CaCO3) process is applicable for both pre-combus-tion and post-combustion carbon dioxide removal [5,6,10,15,16].Based on experimental results, a decay in carbon dioxide removalwhen using natural limestone as the source of calcium carbonateoccurs. To overcome the loss in reactivity, a high Calcium:Carbonmol ratio is necessary along with a high purge stream percentage.Despite the drawbacks, the CaO/CaCO3 process cycle shows greatpromise as a means of carbon dioxide removal based on simula-tions and economics. This is due to the availability of high qualityheat available from both the carbonation and calcination reactionsand the abundance of limestone and dolomite, which provides alow-cost starting raw material. By introducing an intermediateregeneration step between calcination and carbonation, it is possi-ble to reduce the Calcium:Carbon mol ratio, solids circulation, andenergy consumption, all of which improve the process economicsand ability to integrate into a PC power plant [6,15,17,18].

2. CCR Process overview

A variation of the CaO/CaCO3 cycle is the carbonation–calcina-tion reaction (CCR) Process developed at The Ohio State University.The CCR Process has three main components: the Carbonator/Sulf-ator, which will be referred to as the Carbonator, the Calciner, andthe Hydrator. This cyclical process has been successfully demon-strated at The Ohio State University at the 120 kWth scale utilizinga coal and natural gas mixture as the feedstock [6,17–19]. Fig. 1 is adiagram of the CCR Process when incorporated into a coal-firedboiler.

From Fig. 1, the numerous advantages of the CCR Process canbe seen. First, since the entire process occurs at high tempera-tures, a significant amount of high quality heat is generated thatcan be utilized throughout the power plant, either through heatexchange or steam production for electricity generation. Through

intelligent heat integration, a carbon capture process with a lowenergy penalty and favorable economics is achieved. Second, theCCR Process simultaneously removes both carbon dioxide andsulfur dioxide (SO2) in the Carbonator. Although conditions arenot ideal for sulfur removal via solid sorbent, the high Cal-cium:Sulfur mol ratio and the highly reactive calcium hydroxideensure near complete SO2 removal in the CCR Process [19–22].A typical bituminous coal with 75% carbon content at a 1:1 Cal-cium:Carbon mol ratio will possess a 50:1 Calcium:Sulfur mol ra-tio for a 4% sulfur coal. Finally, since the cost of the raw startingmaterial, limestone, is inexpensive, the process will have a fairlylow operating cost.

2.1. Carbonator/sulfator

In the Carbonator, three reactions occur: decomposition of cal-cium hydroxide (Ca(OH)2), carbonation, and sulfation. The decom-position of calcium hydroxide is an endothermic reaction, asshown in (1).

CaðOHÞ2 ! CaOþH2O DH� ¼ þ109 kJ=mol ð1Þ

A relationship between the decomposition temperature of Ca(OH)2

and the partial pressure of water (H2O) exists. Fig. 2 depicts thethermodynamic equilibrium relationship [23]. The dashed line,around 510 �C, is the decomposition temperature of Ca(OH)2 in anatmosphere of steam.

The carbonation reaction has been studied to a significant ex-tent. The irreversible decay in reactivity of CaO with CO2 over mul-tiple cycles has been well-researched [24–27]. The carbonationreaction is exothermic and governed by the thermodynamic equi-librium between temperature and partial pressure of CO2. (2)shows the reaction, while Fig. 3 depicts the thermodynamic equi-librium relationship [23].

CaOþ CO2 ! CaCO3 DH� ¼ �178 kJ=mol ð2Þ

In Fig. 3, the dashed line represents the maximum temperature,around 660 �C, allowable for a 1% CO2 concentration stream. A 1%CO2 concentration is approximately 90% CO2 removal from a coal-fired flue gas stream. Thus, the dashed line represents the maxi-mum temperature the Carbonator can operate to achieve 90% CO2

removal. At temperatures between 660 �C and 775 �C, CO2 removalfrom a coal-fired flue gas stream is possible; however, removals willdecrease exponentially as temperature increases. At temperaturesgreater than 775 �C, the thermodynamic partial pressure of CO2 be-comes greater than the partial pressure of CO2 in a coal-fired fluegas stream, and no removal is expected.

The sulfation reaction, shown in (3), has been commercializedand is the dominant form of sulfur control in power plants[23,28–32].

CaOþ SO2 þ 1=2O2 ! CaSO4 DH� ¼ �506 kJ=mol ð3Þ

Depending on process conditions, the Carbonator may either gener-ate or consume energy. The Calcium:Carbon (Ca:C) mol ratio, fluegas composition, extent of CO2/SO2 removal, operating temperature,and inert solids circulation will determine the Carbonator energystate.

2.2. Calciner

In the calciner, the calcium carbonate formed in the carbonationreaction (2) is decomposed into its constituent components of cal-cium oxide and carbon dioxide. Since the CaO is a solid product andthe CO2 is a gaseous product, a gas–solid separation device can sep-arate the mixture. The CO2 is prepared for sequestration, while theCaO is delivered to the Hydrator to continue the CCR Process cycle.

Fig. 1. Process flow diagram of CCR Process incorporated into coal-fired boiler.

1E-121E-111E-101E-091E-08

0.00000010.0000010.000010.00010.0010.010.1

110

1001000

0 100 200 300 400 500 600 700 800 900 1000 1100 1200

P H2O

(atm

)

Temperature (°C)

Ca(OH)2 CaO + H2O (g)

CaO + H2O (g) Ca(OH)2

Fig. 2. Thermodynamic relationship between Ca(OH)2 and H2O partial pressure.

96 W. Wang et al. / Fuel 92 (2012) 94–106

The calciner presents the greatest obstacle in the CCR Process,as the decomposition of CaCO3 is energy intensive, requires aminimum operating temperature of 900 �C and consumes1500 MJ/ton CaCO3. In addition, the calciner must operate underoxy-fired conditions to generate a high-purity CO2 stream at theexit of the calciner [33–35]. From Fig. 3, the decompositiontemperature is also dictated by the partial pressure of CO2. At low-er CO2 partial pressures, the calcination temperature can belowered. The two current methods of lowering the partial pressureare the addition of a diluent gas and application of a vacuum[36–38]. However, neither method is currently considered suitablefor commercial application.

To further complicate matters, the high-temperature calcina-tion reaction is known to decrease the reactivity of the calcined

CaO [15,25–27,35,39–41]. While a minimum of 900 �C is necessary,current commercial calciners require much greater temperaturesto obtain a high calcium carbonate conversion [42,43]. The CCRProcess overcomes the negative sintering effects of high-tempera-ture calcination by providing an intermediate hydration step,which regenerates the sorbent reactivity.

2.3. Hydrator

The hydration of calcium oxide shows great promise to in-crease reactivity towards CO2 and SO2. The increased reactivityof calcium hydroxide versus calcium oxide derived from calciumcarbonate towards SO2 has been clearly demonstrated [20–22].The Ohio State University has also established the superior

1E-24

1E-22

1E-20

1E-18

1E-16

1E-14

1E-12

1E-10

1E-08

1E-06

0.0001

0.01

1

100

P CO

2(a

tm)

CaO + CO2 CaCO3

CaCO3 CaO + CO2

0 100 200 300 400 500 600 700 800 900 1000 1100 1200

Temperature (°C)

Fig. 3. Thermodynamic relationship between CaCO3 and CO2 partial pressure.

W. Wang et al. / Fuel 92 (2012) 94–106 97

reactivity of calcium hydroxide over calcium oxide with CO2 andhas shown hydration to be an effective means of sorbent regen-eration over multiple carbonation/calcination cycles in its120 kWth facility with co-combustion of coal and natural gas[6,16–19,22]. The use of hydration as an effective method to in-crease sorbent reactivity in a CaO/CaCO3 cycle has been estab-lished [44–48]. By increasing the sorbent reactivity, the Ca:Cmol ratio can be decreased, which will decrease solids circula-tion, calciner energy consumption, and waste disposal rates. Thisultimately leads to an improvement in the economics of the CCRProcess over traditional CaO/CaCO3 cycles. Using a CaO/CaCO3 cy-cle, Ca:C mol ratios between 4 and 8 are typical for 80% CO2

removal [34,49,50]. The Ca:C mol ratio can be greatly decreasedusing Ca(OH)2. Virtually all CO2 is removed at a Ca:C mol ratioof 1.6 [19].

By operating the Hydrator at high temperatures (greater than350 �C), the exothermic heat of reaction can be utilized in theSteam Turbine Cycle. Current commercial hydration occurs atambient temperatures and pressures with excess liquid water toremove the heat liberated. While the ambient hydration will pro-duce a highly reactive sorbent, the inclusion of a low temperaturehydration reaction is detrimental for integrating into a high-tem-perature process and does not generate useful heat that can be uti-lized. At temperatures greater than 510 �C, a pressurized vesselwould be necessary for hydration [23]. While higher temperatureswould be favorable for power plant operations, operating underpressurized conditions would necessarily lead to higher capitaland operating costs. The Hydrator operating conditions will be dic-tated by the economics.

3. Process modeling assumptions

3.1. Aspen software

Computer simulations for integrating the CCR Process into acoal-fired power plant were conducted using Apen Plus� 2004.1software produced by AspenTech. The selected databanks includeCOMBUST, INORGANIC, SOLIDS and PURE13. IDEAL is the propertymethod for conventional components. The properties of the non-conventional components, coal and ash, are calculated based onthe HCOALGEN and DCOALIGT model, respectively.

3.2. Coal combustion

The coal composition and coal feed rate into the Boiler are keptconstant throughout the simulations and follow the values forPittsburgh #8 coal [28]. Table 1 lists the Proximate and UltimateAnalysis. Air consists of 21% volume oxygen (O2) and 79% volumenitrogen (N2). 20% excess air, entering at ambient conditions, issplit into Primary Air and Secondary Air. Primary Air, which is usedto transport and dry the coal, comprises 16.8% of the total inlet air.Secondary Air, which is used for combustion to generate the coalflue gas stream, contains the balance of the total air. The PrimaryAir and Secondary Air are then heat exchanged in the Air Pre-Hea-ter with flue gas exiting the boiler. The Primary Air and SecondaryAir final temperatures are 146 �C and 316 �C, respectively [28]. ThePrimary Air is mixed with the coal feed stream and delivered to theBoiler at a temperature of 93 �C, while the Secondary Air is deliv-ered directly to the Boiler. The combustion of coal was modeledfollowing the guidelines provided by AspenTech.

3.3. Baseline power plant

A 505 MWe net generation power plant is used as the baselinefor all simulations. The pulverized coal-fired boiler combusts205 tons per hour (tph) of coal and generates 1506.8 MWth

(HHV). The boiler output provides 1335.7 MWth of energy, whichyields a boiler efficiency of 88.6% and is typical of coal-fired powerplants currently in operation [51–53]. Given the complexity andproprietary nature of power plant steam turbine cycles, a general42% thermal to electric steam turbine efficiency was assumed[51–53]. Applied to the 1335.7 MWth provided by the flue gas,561.0 MWe gross is generated. In addition, a 10% in-house powerplant electricity consumption was applied, and a net electrical gen-eration of 504.9 MWe is delivered to the grid. The overall efficiencyof the power plant then becomes 33.5% (HHV). Fig. 4 is a processflow diagram of the Baseline Power Plant. Fig. 5 is a diagram ofthe Aspen Model. Table 2 provides the components of the coal-fired flue gas.

3.4. CCR components

In addition to the baseline Boiler, several additional componentsare necessary to complete the CCR Process simulations. Table 3

Table 1Composition of Pittsburgh #8 coal.

Proximate analysis Weight% as-received Weight% dry Ultimate analysis Wt% as-received Weight% dry

Moisture 5.2 Moisture 5.2Fixed carbon 48.1 50.7 Ash 8.6 9.1Volatiles 38.1 40.2 Carbon 70.2 74Ash 8.6 9.1 Hydrogen 4.8 5.1HHV (BTU/lb) 12,540 13,227.8 Nitrogen 1.5 1.6

Chlorine 0 0Sulfur 2.2 2.3Oxygen 7.5 7.9

Fig. 4. Process flow diagram of coal-fired boiler.

98 W. Wang et al. / Fuel 92 (2012) 94–106

provides the remaining Aspen model assumptions. Fig. 6 is a gen-eral process flow diagram of the Aspen simulation. The red linesoutline the CO2-laden streams, the green lines outline the CO2/SO2 lean gas streams, and the blue lines outline the steam path.

3.4.1. BoilerThe outlet flue gas stream from the BOILER, FLUGAS, is cooled to

the temperature necessary such that the Carbonator/Sulfator,CARB, operates at 625 �C with no heat liberation. The heat fromthe flue gas boiler, Q � 1, is high quality heat used in the SteamTurbine Cycle.

3.4.2. PCD-1The flue gas from the boiler, FLUGAS, then enters a Particulate

Capture Device (PCD) to remove 90% of the ash in the FLUGASstream. The Ash PCD, PCD-1, operates at the FLUGAS temperature.The fly ash flow rates in the simulations are higher than those typ-ically found in a PC boiler since no bottom ash is generated. The ef-fect is a lower ash circulation rate in an actual system compared tothe simulations.

3.4.3. CARBThe FLUGAS stream enters CARB at temperatures slightly great-

er than 625 �C to provide the minimal amount of energy necessary

to maintain the Carbonator at 625 �C. The exact inlet flue gas tem-perature is dependent on the Carbonator operating temperature,Hydrator operating temperature, solids circulation, and Ca:C molratio.

For the CCR Process simulations, a Ca:C mol ratio of 1.4 is as-sumed adequate for 90% CO2 and 100% SO2 removal. The Ca:Cmol ratio is based on the total carbon content in the coal and theactive sorbent in the form of calcium hydroxide. For 90% CO2 re-moval and virtually 100% SO2 removal, a 1.3:1 Ca:C mol ratio hasbeen shown to be sufficient [19,22]. The Carbonator is modeledusing an RStoic reactor to remove 90% CO2 and 100% SO2 whilecompletely dehydrating the Ca(OH)2.

3.4.4. PCD-2PCD-2 operates at the identical temperature of the Carbonator,

625 �C. The gas–solid separation efficiency is set to 98%, which canbe achieved with current gas–solid separation devices [56,57]. Thegas stream is then cooled to 350 �C in HX-2 and delivered to the AirPreheater, APH. The solids stream is sent to the Purge, PURGE.

3.4.5. PurgeSince the solids stream will contain a mixture of several inerts,

including deactivated CaO, fly ash, and calcium sulfate, a purgestream is necessary to prevent accumulation of the inerts. The

Fig. 5. Aspen model of 505 MWe coal-fired Boiler.

Table 2Components of flue gas exiting coal-fired boiler.

Coal flue gas component Flow (kmol/h) Flow (tons/h) Volume (mol)%

Nitrogen (N2) 57752.53 1783.37 75.2Carbon dioxide (CO2) 10388.25 503.96 13.5Water (H2O) 4955.28 98.40 6.4Oxygen (O2) 2632.11 92.84 3.4Carbon monoxide (CO) 473.68 14.63 0.6Nitrogen oxide (NO) 455.71 15.07 0.6Sulfur dioxide (SO2) 126.39 8.93 0.2Hydrogen (H2) 41.81 0.093 0.1Nitrogen dioxide (NO2) 0.20 0.01 –Sulfur trioxide (SO3) 0.07 6.16 � 10�3 –S (elemental sulfur) 1.31 � 10�3 4.65 � 10�5 –

Total 76826.03 2517.3 100.0Fly ash 17.685

W. Wang et al. / Fuel 92 (2012) 94–106 99

calcium mol content of the purge stream is replaced by an equiva-lent calcium mol content of a fresh feed stream in order to main-tain a constant Ca:C mol ratio. The location of the purge canoccur in many places, such as after the Carbonator PCD, after thecalciner PCD, or intrinsic to the inefficiencies of the PCD them-selves [34,50,58].

The Purge percentage was set at 2% of the total solids exitingthe solids stream of PCD-2. Since the Purge stream exits at625 �C, a cyclic dual gas–solid heat exchanger system is imple-mented to utilize the available thermal energy. In the first heatexchanger, AIR1, the Purge stream heats air from 50 �C to 600 �C,while the Purge cools from 625 �C to 60 �C. In the second heatexchanger, AIR2, the heated air cools from 600 �C to 50 �C whilethe fresh feed stream is pre-heated from 25 �C to its final temper-ature, which is dependent upon the Fresh Feed inlet and Purgeoutlet conditions. The pre-heated Fresh Feed is then mixed withthe Recycle stream prior to entering the calciner. By pre-heatingthe incoming Fresh Feed, the thermal energy required by thecalciner is reduced, useful energy is not wasted, and the overallenergy balance is improved.

3.4.6. CalcinerThe calciner is simulated using a Gibbs reactor with an operat-

ing temperature of 1000 �C. Heat input into the calciner is neces-sary for solids heating and the calcination reaction. Simulationswere performed using three different calciner configurations,shown in Table 4.

The indirect-fired calciner obtains heat from the flue gasstream exiting the boiler. In an indirect-fired calciner, the heatsource does not directly contact the material to be calcined. Anindirectly-heated calciner is the only configuration that allowsfor a pure CO2 stream. However, an indirect-fired calciner is cur-rently used for niche applications where maintaining product pur-ity is necessary and not the main technique for limestonecalcination [43]. For the natural gas and coal-fired calciner, a95% CO2 purity (dry) stream is generated. The natural gas compo-sition is listed in Table 5 and follows the Department of Energyguidelines [59]. The coal composition provided to the calciner isidentical to the coal provided to the boiler, Pittsburgh #8. TheAir Separation Unit (ASU) was not modeled; however, the ener-getics to operate a cryogenic ASU was factored into the energy

Table 3Assumptions used to model CCR Process.

Unit operation Aspen plus model Specification

Ash particulate capture device SSplit 90% ash removal operating T = flue gas carbonator inlet TCarbonator/sulfator RStoic Ca:C mol ratio = 1.4:1, based on active Ca sorbent and total C in coal 90% CO2/100% SO2 removal

Operating T = 625 �CNo energy generation

Post-carbonator particulate capture device SSPlit Gas–solid separation efficiency = 98%Operating T = 625 �C

Purge Fsplit Splits stream based on mass fractionOperating condition = 2%

Recycle Mixer Combines recycle stream and fresh feed stream with respect to mass and energyEnergy generation heat exchangers Heater Unspecified heat exchange medium for heating and coolingHeat exchangers (M)HeatX Heat exchange two streams

Operates with minimum 10 �C approach TCalciner RGibbs Calculates products based on thermodynamic equilibrium

Operating T = 1000 �CCO2 particulate capture device SSplit Gas–solid separation efficiency = 99.5%

Operating T = 1000 �CCO2 purity RGibbs Minimum 95% (mol) dryHydrator Calculates products based on thermodynamic equilibrium

Operating T = 500 �C H2O:Ca mol ratio = 1.3:1CO2 compression Ref. [54] 119 kW h/tonne CO2 to compress to 14 MPaASU Ref. [55] 200 kW h/tonne O2

Steam Ref. [28] Quality = 0.925Pressure = 2 ‘‘Hg, T = 38 �CObtained from outlet of low-pressure turbine

Fig. 6. Diagram of the Aspen simulation.

100 W. Wang et al. / Fuel 92 (2012) 94–106

consumption [55]. Similarly, the CO2 compression train was notmodeled; however its energetics were factored into the energyconsumption [54].

3.4.7. PCD-3PCD-3 separates the gaseous products from the solid products

exiting the calciner. The gas–solid separation efficiency was set

Table 4Calciner operating conditions for Aspen simulations.

Heatingsource

Operatingtemperature

Excess O2

(%)O2 purity(%)

CO2 purity(%)

Indirect 1000 100Natural gas 1000 10 93.5 95Coal 1013 10 99.8 95

Table 5Natural gas composition, as suggested by department of energy.

Component Chemical formula Volume percentage

Methane CH4 93.1Ethane C2H6 3.2Propane C3H8 0.7n-Butane C4H10 0.4Carbon dioxide CO2 1.0Nitrogen N2 1.6

Fig. 7. Mol fraction of active sorbent as a function of cycle number.

W. Wang et al. / Fuel 92 (2012) 94–106 101

at 99.5%. This high separation efficiency is necessary as the high-purity CO2 stream will be sequestered and cannot tolerate a signif-icant amount of impurities. The solids were directed towards theHYDRATOR at the calciner operating temperature, while thehigh-purity CO2 stream was cooled to 350 �C and its thermal en-ergy used in the Steam Turbine Cycle. The remaining energy wasused to heat exchange with the oxygen entering the calciner andthe steam entering the hydrator.

3.4.8. HydratorThe HYDRATOR is a Gibbs reactor operating at 500 �C. The heat

liberated from the hydration reaction is sent to the Steam TurbineCycle. The HYDRATOR obtains steam from the exit of the Low-Pres-sure Turbine, which exits under vacuum at approximately 2’’ Hg,wet with a quality of 0.925, and at 38 �C [28]. The steam is heatexchanged with the CO2 stream to increase the steam temperatureto greater than 100 �C, where the pressure can then be raised toatmospheric and remain as steam. Excess steam is deliveredto the HYDRATOR, and a Calcium:Steam mol ratio of 1.3 is assumedto be sufficient for complete hydration. As steam quality increases,the efficiency of the CCR Process integrated into a power plant willincrease since less energy will be consumed by the latent heat ofvaporization in HX-VAC.

Theoretically, a post-Hydrator gas–solid separation device canbe envisioned. By doing so, the excess steam can be recycled backinto the Hydrator and lower the overall steam consumption.However, difficulties arise in reality due to a lack of motive forcebetween the Hydrator, post-Hydrator PCD, and Carbonator. Fur-ther, the calcium hydroxide can theoretically be decomposed inan independent vessel located after the Hydrator. A PCD wouldthen separate the steam from the hydrate with the steam recy-cled into the Hydrator and the oxide delivered to the Carbonator.In this configuration, only an initial charge of steam along withmake-up steam is necessary. However, additional sources of en-ergy would need to be provided for the hydrate decomposition,and the lack of a motive force for gas–solid separation would alsoexist.

3.4.9. Sorbent reactivityThe overall success of the CCR Process for simultaneous CO2 and

SO2 removal from a coal-fired power plant is dependent on theability of Ca(OH)2 to maintain a high reactivity over multiple cy-cles. While the main function of the Purge is to prevent accumula-tion of inerts, it also allows for the introduction of fresh sorbentinto the CCR Process cycle. Overall, the inefficiencies of the PCDs,

the Purge, and the permanent binding of sorbent to SO2 provideroutes for sorbent losses. For combustion of Pittsburgh #8 coalwith complete removal of SO2, a 2% Purge rate, 98% post-Carbona-tor PCD efficiency, and 99.5% post-calciner PCD efficiency, thereplenishment of fresh limestone is 5.23% of the total mols of ac-tive sorbent.

While the Ca:C mol ratio at the Carbonator inlet maintains aconstant value of 1.4, the solids will possess a distribution of vary-ing carbonation-calcination-hydration cycles. The cyclic distribu-tion in conjunction with cyclic reactivity is pertinent to anycyclical process to account for decay in sorbent reactivity. Fig. 7shows the percentage of active sorbent, in mols, as a function of cy-cle number [24]. The initial charge consists entirely of fresh CaCO3,while all subsequent cycles consist of a mixture of fresh and cycledsorbent. An underlying assumption presumes an equal probabilityof removal through previously stated losses independent of cyclenumber, mass fraction, and particle size distribution.

Bench-scale studies using high-calcium limestone were con-ducted over 15 cycles to quantify the effect of hydration on sorbentreactivity over multiple cycles. Using the cyclic mol ratio distribu-tion and the CO2 removal for each respective cycle, the Ca:C molratio for 90% removal over 15 cycles is equivalent to a Ca:C mol ra-tio of 1.44. This validates the assumption of a Ca:C mol ratio of 1.4used in the simulations. The Ca:C mol ratio will be determined bycoal sulfur content, overall sorbent loss, and cyclic reactivity of sor-bent, as variations will exist.

3.4.10. Electricity generationNo deductions in thermal energy from the Boiler to the Steam

Turbine Cycle are taken since the Boiler efficiency is already88.6%. A 90% thermal efficiency is placed on all heat entering theSteam Turbine Cycle from the CCR Process. A general 42% thermalto electric efficiency is employed to model the Steam Turbine Cy-cle. To account for in-house electric consumption, pressure drops,and heat loss, an additional 10% energy penalty is added to thegross electric generation. When appropriate, a 200 kWh/tonne O2

electricity consumption for the ASU and 119 kWh/tonne CO2 forCO2 is deducted to obtain the final value of Net Electricity to Grid[54,55]. (4)–(8) provides the formula for determining Net Electric-ity to Grid.

Thermal Energy from Coal-Fired Boiler ¼ Q � 1þ Q � 2 ð4Þ

Thermal Energy from CCR Process ¼ ð0:9Þ�ðQ �Hydþ Q � 3Þð5Þ

Table 6Energy consumption in the calciner by fuel type.

Calcinerfuel

Fuelconsumption

Calcinerenergy(MW t h)

ASUenergy(MWe)

CO2

compression(MWe)

Indirect 681.9 0.0 51.22Natural Gas 67 tons/hour 911.8 49.5 72.74Coal 120 tons/hour 882.0 52.8 89.0

102 W. Wang et al. / Fuel 92 (2012) 94–106

Gross Electricity Generation¼ ð0:42Þ�ðThermal Energy from Coal-Fired Boilerþ Thermal Energy from CCR ProcessÞ ð6Þ

Net Electricity Generation ¼ ð0:9Þ�ðGross Electricity GenerationÞð7Þ

Net Electricity to Grid ¼ Net Electricity Generation

� CO2 Compression� ASU ð8Þ

4. Results

The energy necessary to operate the calciner is the single largestconsumer of energy in the CCR Process. Table 6 breaks down theenergy consuming processes in the calciner. Operating the calcinerin an indirect-fired manner consumes the least amount of energyby a significant amount. Since very little difference exists betweennatural gas and coal, the determination would be based oneconomics.

The solids circulation also varies by calciner type. Althoughquite large, the solids circulation is manageable and similar tothe solids circulation rate found in Fluidized Catalytic Cracking(FCC) Processes. Fig. 8 shows the variation in solids circulation atthe reactor inlet. The solid horizontal line represents the averagesolids circulation in the FCC Process. The coal-fired calciner

Fig. 8. Solids circulation within the CCR

processes more solids since both ash and calcium sulfate are gen-erated in the oxy-combustion process. In addition, an increasedfresh feed rate is necessary with the coal-fired calciner to compen-sate for the calcium oxide being converted to calcium sulfate in thecalciner.

Any carbon removal process will necessarily reduce the electri-cal output of a power plant. With respect to energetics, the CCRProcess provides the lowest percentage increase in thermal energyto generate 1 MWe when compared with competing technologiesof scrubbing with econamine and oxycombustion. The increase inratio of thermal energy to electric energy represents the additionalincrease in fuel necessary to maintain the identical electrical out-put as the baseline boiler without carbon dioxide control. Table 7shows the conversion of thermal energy into electricity for thethree calciners, and Table 8 shows the comparison to competingtechnologies [60].

The base ultrasupercritical plant has a thermal to electric effi-ciency of 44.6%, while the base CCR case has a thermal to electricefficiency of 33.5%. When the CCR Process is installed, a maximumof 28.4% increase in thermal to electric conversion occurs. The bestcase scenario for amine or oxycombustion requires an increase of34.16% in thermal to electric conversion. Starting the base CCR caseat a higher thermal to electric efficiency will only magnify the dis-crepancy in efficiency the CCR Process CO2 control system pos-sesses when integrated into a power plant.

The increase in thermal energy necessary to maintain the iden-tical electrical output is also reflected in the reduction in electricaloutput when the thermal input remains constant. For a 205 ton perhour coal input, the CCR Process operating with a coal-fired calcin-er delivers 130 tons per hour to the boiler and 75 tons per hour tothe calciner. Under identical simulation conditions, 399.5 MWe isdelivered to the grid, which is a 20.9% decrease in electrical outputcompared to the base case and an efficiency of 26.5% (HHV). Table 9compares the energy penalty of the CCR Process compared to othercarbon removal processes [1,60].

The CCR Process is competitive, and typically superior, to themajority of competing technologies with respect to energy penalty.

Process for various calciner fuels.

Table 7Electricity generation based on calciner type.

Calciner Boiler (MWth) CCR Process (MWth) Gross electricity (MWe) In-house (MWe) CO2 compression (MWe) ASU (MWe) Net to grid (MWe)

Indirect 651.3 569.34 512.67 �56.0994 �51.22 0.0 405.35Natural gas 1333.8 659.97 837.38 �83.738 �72.74 �49.5 631.4Coal 1319.2 731.25 861.19 �86.119 �89.0 �52.8 633.3

Table 8Comparison of thermal energy ratio/net electricity to grid for CO2 technologies.

Thermal input (MWth) (HHV) Net electricity (MWe) Ratio of MWth/MWe Percent increase (%)

CCR base 1506.84 504.8946 2.9845CCR + indirect 1506.84 405.35 3.7174 +24.5CCR + natural gas 2418.64 631.4 3.8305 +28.4CCR + coal 2388.84 633.3 3.7721 +26.4Supercritical base 1396.2 550.03 2.5384Supercritical + amine 1944.2 549.97 3.5352 +39.2Supercritical + oxycombustion 1878.6 550.01 3.4155 +34.55Ultrasupercritical base 1233.8 549.99 2.2434Ultrasupercritical + amine 1655.3 549.96 3.0098 +34.16Ultrasupercritical + oxycombustion 1668.7 549.98 3.0342 +35.25

Table 9Decrease in power plant efficiency for various carbon dioxide removal processes.

HHV efficiency(%)

Percent decrease(%)

CCR base 33.5CCR + coal 26.5 20.9Subcritical 34.3Subcritical + MEA 25.1 26.8Supercritical 38.5–39.4Supercritical + amine 28.3–29.3 23.9–28.2Supercritical + oxycombustion 29.3–30.6 20.5–25.6Ultrasupercritical base 43.3–44.6Ultrasupercritical + amine 33.2–34.1 21.2–25.6Ultrasupercritical + oxycombustion 33.0 26.0

W. Wang et al. / Fuel 92 (2012) 94–106 103

Depending on the study used, only the supercritical boiler withoxycombustion has a lower decrease in efficiency and an ultrasu-percritical boiler with amine scrubbing is competitive.

The Hydrator is similar to an amine scrubber in that steam is re-moved from the Steam Turbine Cycle. However, unlike an aminescrubber, the Hydrator consumes steam from the exit of thelow-pressure steam turbine and does not de-rate the electricitygeneration of the power plant by extracting steam prior to thecompletion of the Steam Turbine Cycle. The amount of steam con-sumed in the Hydrator depends on the Ca:C mol ratio, extent ofhydration, and excess steam requirements. Table 10 shows thevariation in steam consumption assuming complete hydration.

The steam flowrate circulating in a coal-fired power plant isapproximately 3.3 tons per hour per Megawatt electric [28,53,61].For a 505 MWe coal-fired power plant, the total steam flowrate isapproximately 1665 tons per hour of steam. The steam consump-tion by the Hydrator will range between 10% at a minimum and30% at a maximum of the total steam flowrate, depending on steamrequirements by the Hydrator and total steam circulating in the

Table 10Hydrator steam consumption (tons per hour).

Ca:C mol ratio

H2O:Ca mol ratio 0.912 Stoichiometric min 1.0 1.41.0 Stoichiometric min 197 216 3021.1 217 237 3321.3 256 281 3931.5 295 324 453

Steam Turbine Cycle. An amine scrubber consumes approximatelyhalf the steam that would enter the low pressure steam turbine,which translates to approximately 40% of the total steam flowrate[61]. Compared to an amine scrubber, the Hydrator steam con-sumption is significantly lower and extracted at the exit of theSteam Turbine Cycle to lessen the impact on electricity generation.

4.1. Sensitivity analysis

The multitude of parameters that can be varied require exami-nation to determine their effect on the overall CCR Process design.By minimizing the calciner energy requirements, the annual oper-ating cost can be reduced. Fig. 9 shows the effect the Ca:C mol ra-tio, post-carbonator PCD efficiency, and purge percentage has onthe calciner energy requirements. The analysis is performed onan indirectly-heated calciner, but the trends can be extended toother cases. The static conditions are identical to the Aspen simu-lation model-Ash PCD efficiency of 90%, complete calcination, com-plete hydration, a post-calciner PCD efficiency of 99.5%, 625 �Ccarbonation temperature, 1000 �C calcination temperature, 500 �Chydration temperature and 90% CO2/100% SO2 removal at a Ca:Cmol ratio of 1.4.

For any Ca:C mol ratio greater than the minimum, there exists aminimum energy consumption by the calciner due to the differ-ences in specific heat capacity of the various circulating solids.Within each Ca:C mol ratio, the minimum calcination energy isthe same. Increasing the post-Carbonator PCD efficiency increasesthe purge percentage necessary to operate at the minima. Operat-ing at a lower Ca:C mol ratio does not guarantee that less energywill be consumed by the calciner. For example, operating at aCa:C mol ratio of 0.912, the stoichiometric minimum for 90% CO2

and 100% SO2 removal, with a carbonator PCD efficiency of 99.5%and a 2% purge is identical to operating at a Ca:C mol ratio of1.40 with a carbonator PCD efficiency of 95% and 4% purge.

Figs. 10 and 11 are identical in conditions, except for the Purgepercentage. The Ca:C mol ratio is fixed at 1.40, the post-carbonatorPCD efficiency is 98%, calcination and hydration are both complete,and the post-calciner PCD efficiency is 99.5%. The purge percentageis 2% and 6% in Figs. 9 and 10, respectively. The figures show theash circulation as a fraction of the total solids at varying ash PCDefficiencies. From Fig. 10, the difference in solids circulation be-tween not having an ash PCD and a 90% efficient ash PCD is approx-imately 250 tons per hour. From Fig. 11, the difference in solidsbetween not having an ash PCD and a 90% efficient ash PCD is

Fig. 9. Variation of calciner energy consumption with multiple parameters.

Fig. 10. Ash circulation (shaded area) as a function of total solids circulation at varying ash PCD efficiencies at a purge rate of 2%.

104 W. Wang et al. / Fuel 92 (2012) 94–106

approximately 130 tons per hour. At higher purge rates or combus-tion of low ash coal, the ash PCD may not be an essential piece ofequipment. By not installing non-essential pieces of equipment ata specific operating condition and choosing the proper operatingconditions, it is possible to reduce both the costs and energy pen-alty of the CCR Process when fully integrated into a power plant.

5. Conclusion

The CCR Process with intermediate hydration has the ability ofbeing an effective method for simultaneously separating carbon

dioxide and sulfur dioxide from coal-fired power plant flue gas.Based on process simulation results, 90% CO2 removal and com-plete SO2 removal can be achieved with a 28% increase in thermalenergy necessary to maintain a constant electrical output. Thesteam consumed by the CCR Process is significantly lower com-pared to amine scrubbing. While the CCR Process has many processparameters that can be varied, they can be optimized to minimizeenergy consumption and solids circulation. Commercially-developed equipment and an abundant supply of raw startingmaterial reduces the process economics and research necessaryto retrofit the CCR Process into existing power plants. Utilizing a

Fig. 11. Ash circulation (shaded area) as a function of total solids circulation at varying ash PCD efficiencies at a purge rate of 6%.

W. Wang et al. / Fuel 92 (2012) 94–106 105

limestone-based carbon capture process, the CCR Process can haveminimal impact when integrated into a power plant and providesan efficient means of controlling both carbon and sulfur emissions.

Acknowledgements

The financial assistance provided by The Ohio Coal Develop-ment Office (OCDO) of the Ohio Air Quality Development Authority(OAQDA) in support of the CCR Process demonstration is gratefullyacknowledged.

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