Thermodynamic and Experimental Analyses of the Three-Stage

Jul 20, 2010 - Application of a novel calcium looping process for production of heat and carbon dioxide enrichment of greenhouses. Mohammad Ramezani ,...
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Ind. Eng. Chem. Res. 2010, 49, 7563–7573

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Thermodynamic and Experimental Analyses of the Three-Stage Calcium Looping Process Shwetha Ramkumar and Liang-Shih Fan* William G. Lowrie Department of Chemical and Biomolecular Engineering, 121 Koffolt Laboratories, The Ohio State UniVersity, 140 West 19th AVenue, Columbus, Ohio 43210

Clean-coal technologies that include carbon dioxide and sulfur capture during the production of electric power, liquid fuels, and hydrogen represent a major thrust area. The calcium looping process (CLP) is one such technology that is being developed to convert syngas obtained from coal gasification to hydrogen using a regenerable calcium oxide sorbent. It integrates the water-gas shift reaction with in situ carbon dioxide, sulfur, and halide removal at high temperatures while eliminating the need for a water-gas shift catalyst and reducing the overall footprint of the hydrogen production process. The CLP comprises three reactors: the carbonation reactor, where the thermodynamic constraint of the water-gas shift reaction is overcome by the constant removal of the carbon dioxide product and high-purity hydrogen is produced with contaminant removal; the calciner, where the calcium sorbent is regenerated and a sequestration-ready carbon dioxide stream is produced; and the hydrator, where the calcined sorbent is reactivated to improve its recyclability. In this article, the reaction chemistry occurring in the three calcium looping reactors and the performance of the reactors at various process conditions are presented through thermodynamic and experimental analyses. Highpurity H2 with less than 1 ppm of H2S is obtained in the carbonation stage at a stoichiometric steam-tocarbon ratio at high pressures. Although calcination of the sorbent under realistic conditions causes severe sintering and a loss in reactivity, sorbent reactivation by hydration is effective in restoring sorbent reactivity. 1. Introduction The world energy demand is projected to increase by 40% at a rate of 1.5% per year from 2007 to 2030.1 Although the energy generation from renewable resources is projected to grow, fossil fuels are still projected to contribute a major portion of world energy needs in the near future.1 A growing need for the reduction of anthropogenic carbon dioxide (CO2) emissions has led to a global push toward the development of efficient, economical, and reliable carbon capture and sequestration (CCS) technologies for application to fossil-fuel-based power plants. The implementation of CO2 capture in fossil-fuel-based systems could be through postcombustion capture, oxycombustion, and precombustion capture as illustrated in Figure 1. Postcombustion capture technology involves the combustion of coal or natural gas to produce hot flue gas, which is used to generate steam. The CO2 from the flue gas is then captured. The capture of CO2 from flue gas results in a large increase in parasitic energy and cost of electricity due to the large volumes of flue gas and low concentration of CO2 (13-14% for coal combustion and 3-4% for natural gas combustion). In oxycombustion, the fuel is burnt in oxygen and recycled flue gas, to produce a concentrated stream containing CO2 and steam, which is then dried, compressed, and transported for sequestration. Although oxycombustion obviates the need for a separate CO2 capture stage, it requires an air separation unit (ASU), which is energy-intensive and expensive. Precombustion capture involves the gasification of coal or the reforming of natural gas to produce syngas. The syngas is then cleaned and sent to shift reactors to convert the carbon monoxide (CO) to hydrogen (H2) and CO2 in the presence of steam. For syngas from coal gasification, the water-gas shift reaction can be conducted as a raw syngas (sour) shift or a clean syngas (sweet) shift.2 Commercially, the clean water-gas shift reaction is carried out * To whom correspondence should be addressed. Tel.: (614)-6883262. Fax: (614)-292-3769. E-mail: [email protected].

in two stages: in high- and low-temperature shift reactors using iron oxide and copper catalysts, respectively.3 The hightemperature shift (HTS) is conducted to convert the bulk of the CO to H2 because of its fast kinetics. The low-temperature shift (LTS) reaction is carried out because the equilibrium conversion is higher at lower temperatures, but the gas stream has to be cooled to 180-270 °C, which makes the process energy-inefficient.3 The commercial iron oxide catalyst has a sulfur tolerance of about several hundred parts per million, whereas the copper catalyst has a lower tolerance to sulfur and chloride impurities.3 Hence, syngas cleanup is required to remove sulfur, chloride, and other impurities upstream of the shift reactors and CO2 downstream of the shift reactors. The sour gas shift uses a sulfided catalyst that is resistant to high sulfur concentrations in the syngas stream and operates at a temperature of 250-500 °C. By using the raw gas shift, sulfur removal and CO2 removal can be conducted downstream of the shift reactor in an integrated mode.4,5 However, the extent of CO conversion is lower in the raw gas shift than in the clean gas shift. The application of CCS to gasification systems has been found to be more efficient and economical than CCS for postcombustion systems. It has been estimated that, with the implementation of CCS using solvent-based systems, the increase in the cost of electricity (COE) for an integrated gasification combined cycle (IGCC) will be 25-40% whereas that for pulverized coal (PC) boilers will be 60-85%.6 In a carbon-constrained scenario, it has been estimated that the cost of a supercritical PC boiler will be $2140/kWe whereas that of an IGCC will be $1890/ kWe.6 In addition to being more economical and efficient, gasification is also very versatile and capable of producing H2 and liquid fuels in addition to electricity.6 Several options are being investigated for the implementation of CCS on precombustion and postcombustion systems, including use of solvents, sorbents, membranes, and chemical looping processes. The calcium looping process (CLP), which is a calcium-sorbent-based chemical looping process, has the po-

10.1021/ie100846u  2010 American Chemical Society Published on Web 07/20/2010

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Figure 1. Implementation of carbon capture and sequestration (CCS) in fossil-fuel-based power plants.

tential to reduce the cost and increase the efficiency of CCS implementation on postcombustion and precombustion systems.7,8 The concept of utilizing calcium oxide (CaO) for CO2 capture has existed for well over a century. It was first introduced by DuMotay and Marechal in 1869 for enhancing the gasification of coal9 and was tested a century later in a 40 t/day plant by CONSOL’s CO2 Acceptor process.10 A variation of this process, the HyPrRing process,11,12 was developed in Japan for the production of H2 at high pressures. Several other processes have also been developed to enhance H2 production using calciumbased sorbents such as the ZECA,13 Alstom,14 and GE processes.15 Shimizu et al.,16 Gupta and Fan,18 Iyer et al.,17 Wang et al,19 Sun et al.,20,21 Abanades et al.,22,7 and Manovic et al.23 applied this concept to the removal of CO2 from combustion flue gas. Brun-Tsekhovoi et al.,24 Fan et al.,25 Iyer et al,26,27 Ortiz and Harrison,28 Han and Harrison,29 Johnsen et al.,30 Balasubramanian et al.,31 and Hufton et al.32 applied this concept to the removal of CO2 and the production of H2 from syngas through the water-gas shift reaction and from methane (CH4) through the sorption-enhanced steam methane reforming reaction. In the application of calcium-based CO2 capture to postcombustion systems, CaO sorbent is used to react with CO2 in the flue gas thereby removing it. The calcium carbonate (CaCO3) formed in this process is then regenerated in a separate reactor called the calciner to produce a sequestration-ready pure CO2 stream. The details of the implementation of calcium-based simultaneous CO2 and SO2 removal by the carbonation-calcination reaction (CCR) process in a 120 kWth subpilot scale unit at the Ohio State University have been reported elsewhere.19 The CCR process uses a regenerable calcium-based sorbent with sorbent reactivation by hydration as a step in the carbonation-calcination cycle to prevent the decay in sorbent reactivity over multiple cycles.19 The implementation of the CLP in a precombustion CO2 capture scenario is illustrated in Figure 2. The CLP not only aids in curbing CO2 emissions but also improves the yield of H2 and reduces the number of process units by implementing the concept of process intensification. Syngas obtained from coal gasification is sent through a particulate capture device to the carbonation reactor. Steam is added to the carbonation reactor for the water-gas shift reaction. When CaO is injected into the syngas, it reacts with the CO2, sulfur (H2S, COS), and chloride (HCl) to form a mixture containing predominantly CaCO3 and small amounts of calcium

Figure 2. Integration of the CLP in a coal gasification system.

sulfide (CaS) and calcium chloride (CaCl2). The in situ removal of CO2 removes the equilibrium limitation of the water-gas shift reaction resulting in the enhanced production of H2. The H2 produced in the carbonation reactor is then separated from the spent sorbent in a particle capture device and could be used for electricity generation or for the synthesis of fuels and chemicals. The CaCO3 in the spent sorbent is regenerated back to CaO in the calciner at a temperature of 800-1000 °C. The CO2 produced in the calciner is then dried, compressed, and transported for sequestration. To improve recyclability of the sorbent, a sorbent reactivation step by hydration is introduced as a step in the carbonation-calcination cycle to prevent the decay in sorbent reactivity over multiple cycles. Reactivation of all or part of the calcined sorbent by hydration has been found to be very effective to reverse the effect of sorbent sintering. This study provides an overview of the CLP for the production of H2 from gasifier syngas. It describes the reaction stages in the CLP through thermodynamic and experimental analyses. 2. Calcium Looping Process Overview A schematic of the CLP is shown in Figure 3. The equipment for the CLP comprises a carbonation reactor, a calciner, and a

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Sulfur capture (COS) CaO + COS S CaS + CO2

(4)

Halide capture (HCl) CaO + 2HCl S CaCl2 + H2O

(5)

2.2. Thermodynamic Analysis of Reactions Occurring in the Carbonation Reactor. The equilibrium constants for the water-gas shift reaction and the combined water-gas shift and carbonation reaction for various temperatures are shown in Figure 4 as a function of temperature. The equilibrium constants were obtained using HSC Chemistry v 5.0 software. The equilibrium constant for the water-gas shift reaction can be defined as Keq1 )

Figure 3. Schematic of the CLP.

hydrator. In the carbonation reactor, high-purity H2 is produced and contaminant removal is achieved; in the calciner, the calcium sorbent is regenerated and a sequestration-ready CO2 stream is produced; and in the hydrator, the sorbent is reactivated. In this work, thermodynamic analyses were conducted for the reactions occurring in each reactor using HSC Chemistry v 5.0 software (Outokumpu Research Oy, Pori, Finland). All of the reactions shown in Figure 3 were found to be thermodynamically spontaneous but reversible, with the extent of each of the reactions depending on the partial pressure of the respective gas species and the reaction temperature. The following sections provide descriptions of the three reactors using thermodynamic and experimental analyses. 2.1. Carbonation Reactor. The carbonation reactor comprises either a fluidized-bed or entrained-flow reactor that operates at pressures ranging from 1 to 30 atm and temperatures of 500-750 °C. The exothermic heat released from the carbonation reactor can be used to generate steam or electricity. In the carbonation reactor, the thermodynamic constraint of the water-gas shift reaction is overcome by the continuous removal of the CO2 product from the reaction mixture, which enhances H2 production and obviates the need for excess steam addition. This is achieved by the concurrent water-gas shift reaction and carbonation reaction of CaO to form CaCO3, thereby removing the CO2 product from the reaction mixture. In addition, the CaO sorbent is also capable of reducing the concentrations of sulfur and halides in the outlet stream to part-per-million levels. Because the combined water-gas shift and carbonation reaction for H2 production occurs at a high temperature of 500-750 °C, the need for a water-gas shift catalyst is eliminated. The reactions occurring in the carbonation reactor are as follows: Water-gas shift reaction CO + H2O S H2 + CO2 (∆H ) -41 kJ/mol)

(1)

Carbonation CaO + CO2 S CaCO3 (∆H ) -178 kJ/mol)

(2)

Sulfur capture (H2S) CaO + H2S S CaS + H2O

(3)

PH2PCO2 PCOPH2O

where PCO2, PH2, PCO, and PH2O are the partial pressures at equilibrium of CO2, H2, CO, and H2O, respectively. The equilibrium constant for the combined water-gas shift and carbonation reaction can then be defined as Combined water-gas shift and carbonation reaction CO + H2O + CaO S H2 + CaCO3 Keq2 )

(6)

P H2 PCOPH2O

where Keq2 ) Keq1Kcarb and Kcarb is the equilibrium constant of the carbonation reaction. The equilibrium of the water-gas shift reaction decreases with increasing temperature, resulting in low H2 yields at higher temperatures. Hence, in the conventional water-gas shift system, an LTS is used after the HTS to convert the CO slip and to increase the yield of H2 in the presence of an LTS catalyst. The equilibrium constant of the combined water-gas shift and carbonation reaction is significantly higher than the equilibrium constant of the water-gas shift reaction in the desired operating temperature range of 500-750 °C. Hence, the CLP is capable of producing a much higher H2 yield and, hence, purity through its almost complete CO conversion, when compared to the conventional H2 production process.

Figure 4. Thermodynamic data illustrating the equilibrium constants of the water-gas shift reaction and the combined water-gas shift and carbonation reaction.

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Table 1. Typical Fuel Gas Compositions Obtained from Different Gasifiers48

oxidant fuel pressure (atm) CO (mol %) H2 (mol %) CO2 (mol %) H2O (mol %) N2 (mol %) CH4 + HCs (mol %) H2S + COS (mol %)

moving bed, dry

moving bed slagging

fluidized bed

entrained flow, slurry

entrained flow, dry

air subbituminous 20.1 17.4 23.3 14.8 38.5 5.8 0.2

oxygen bituminous 31.6 46 26.4 2.9 16.3 2.8 4.2 1.1

oxygen lignite 9.9 48.2 30.6 8.2 9.1 0.7 2.8 0.4

oxygen bituminous 41.8 41 29.8 10.2 17.1 0.8 0.3 1.1

oxygen bituminous 24.8 60.3 30 1.6 2 4.7 1.3

To estimate the purity of H2 and the extent of carbon capture that can be achieved in a commercial coal-to-H2 plant by the introduction of the CLP, thermodynamic analysis was conducted for various gasifier syngas compositions using the equilibrium constants shown in Figure 4. The typical compositions of syngas from different gasifiers are listed in Table 1. Steam is added to the syngas before it is fed to the carbonation reactor to adjust the steam-to-carbon (S/C) ratio for the water-gas shift reaction. Using the composition of the syngas, the purity of H2 produced from the CLP and the amount of carbon capture obtained in the temperature range of 500-750 °C were determined. The temperature range of 500-750 °C was chosen as the preferred operating range primarily because the kinetics of CO2 capture by CaO is high in this range. It is noted that air was used as the oxidant in the dry moving-bed gasifier whereas oxygen was used in all other gasifiers. Figure 5 illustrates the purities of H2 produced by the carbonation reactor in the CLP at S/C ratios of 1:1 and 3:1. High H2 purities can be obtained at both S/C ratios of 3:1 and 1:1 in all the gasifiers where oxygen is used as the oxidant. In the dry moving-bed gasifier, lower H2 purities were obtained as a result of dilution by nitrogen because air was used as the oxidant in the gasifier. Figure 6 depicts the total equilibrium carbon capture obtained in the CLP at S/C ratios of 1:1 and 3:1. The percentage carbon captured is defined as the ratio of the total number of moles of carbon in the form of CO, CO2 and CH4 removed in the process to the moles of carbon in the syngas feed times 100%. The percentage carbon captured for all syngas compositions and S/C ratios was found to decrease with increasing temperature. This decrease is due to the decrease in not only CO conversion but also CO2 capture by CaO at high temperatures. It can be seen that greater than 95% carbon capture can be obtained from all of the gasifiers by operating in the temperature range of 550-650 °C. Although greater CO conversions could be obtained at temperatures lower than 550 °C, the kinetics of the water-gas shift reaction and CO2 removal by CaO would decrease, resulting in the need for the use of larger reactors. The percentage carbon captured was found to increase with increasing S/C ratio from 1:1 to 3:1 for all compositions of syngas.

The thermodynamics of H2S capture by CaO sorbent in the carbonation reactor was also investigated. For the reversible sulfidation of CaO, the extent of H2S removal depends on the temperature and pressure of water vapor (PH2O) in the carbonation reactor. Figure 7 depicts the equilibrium H2S concentrations in the product H2 stream, in parts per million, for varying moisture concentrations (PH2O) at 30 atm total system pressure. It can be seen that the equilibrium H2S concentration in the product H2 stream increases with increasing PH2O. At a temperature of 600 °C, the H2S concentration is 0.1 ppm for PH2O ) 0.02 atm and 1 ppm for PH2O ) 0.2 atm. By operating the carbonation reactor at near-stoichiometric steam requirement, it is possible to obtain low concentrations of steam in the reactor system, leading to low H2S concentrations of less than 1 ppm in the product stream. It can also be seen that the reactor system favors H2S removal using CaO at ∼500-650 °C, which is a suitable temperature for the carbonation reaction as well. 2.3. Calciner. The spent sorbent at the exit of the carbonation reactor is a mixture consisting of CaCO3, CaO, CaS, and CaCl2. The CaCO3 in the spent sorbent mixture is regenerated back to CaO in the calciner. The calciner is operated at atmospheric pressure in a rotary- or fluidized-bed system. The heat can be supplied directly or indirectly using a mixture of fuel and oxidant. From the thermodynamic curve for CaO and CO2 shown in Figure 8, calcination was found to occur at temperatures above 890 °C in the presence of 1 atm of CO2. Dilution of CO2 in an indirectly fired calciner with steam or oxycombustion of a fuel (syngas, natural gas, coal, etc.) in a directly fired calciner would permit the calcination reaction to be conducted at temperatures lower than 890 °C. The reaction occurring in the calciner is Calcination

CaCO3 f CaO + CO2

(7)

The regenerability of the CaO sorbent over multiple cycles has been the major drawback of high-temperature calcium-based CO2 capture processes. CaO sorbents are prone to sintering during the high-temperature calcination step. There is a decrease

Figure 5. Effect of temperature on equilibrium H2 purity in the presence of CaO at S/C ratios of (a) 1:1 and (b) 3:1.

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Figure 6. Effect of temperature and S/C ratio on the percentage of carbon captured in the CLP using syngas from different gasifiers as the feed. Figure 9. Thermodynamic data for the hydration of CaO.

Figure 7. Thermodynamic data for the sulfidation (H2S) of CaO with varying steam partial pressures.

process outlined in this article, the addition of a sorbent reactivation step by hydration, as part of the carbonationcalcination cycle, was used to reverse the effect of sintering during each cycle and thus maintain the sorbent reactivity.51 2.4. Sorbent Reactivation by Hydration. The calcination process causes sintering of the sorbent, which results in a reduction in its reactivity and, hence, the overall CO2 capture capacity. The hydration process reverses this effect by increasing the pore volume and surface area available for reaction with the gas mixture. Figure 9 shows the partial pressure of steam required for hydration of the sorbent at various temperatures. Hydration occurs at atmospheric pressure at temperatures below 500 °C. At temperatures of 600 °C and above, hydration occurs at steam partial pressures of above 4 atm. Operation of the hydrator at high temperatures reduces the extent of cooling and reheating of the solids required between the calciner and the carbonation reactor. This aids in reducing the parasitic energy consumption of the process. Hydration at higher temperatures also produces high-quality heat that can be used to produce steam or electricity. Depending of the reactivity of the calcined sorbent, a fraction of the calcined sorbent or the entire stream of sorbent could be hydrated. The reactivity of the calcined sorbent depends on a variety of factors, including the type of calciner (direct or indirect), mode of calcination (rotary kiln, fluidized bed, or entrained bed), temperature of calcination, and gas atmosphere within the calciner. The reaction occurring in the hydrator is Hydration

Figure 8. Thermodynamic curve for the carbonation of CaO.

in sorbent reactivity even when steam is present in the calcination atmosphere. Over multiple cycles, the percentage of sintered CaO progressively increases, which reduces the CO2 capture capacity of the sorbent.10,17,21,33-39 Because of sintering, higher solids circulation or makeup rates need to be used to maintain a high level of CO2 removal.40 Pretreatment methods have been developed to reduce the decay in reactivity, including hydration of the sorbent,37,41-44 preheating and grinding of the sorbent,45 and synthesis of novel sorbents by physical or chemical modification of the precursor.46,44,47-50 In the CLP

CaO + H2O ) Ca(OH)2

(8)

The calcium hydroxide [Ca(OH)2] from the hydrator is conveyed to the carbonation reactor, where it dehydrates to produce high-reactivity CaO and steam. The steam obtained from the dehydration reaction is consumed in the water-gas shift reaction. The advantage of this reactivation process is that no excess steam is required for hydration. Part or all of the steam required for the water-gas shift reaction is supplied to the hydrator depending on the fraction of the calcined sorbent that is sent to the hydrator for reactivation. 3. Experimental Methods 3.1. Carbonation Reactor: Bench-Scale Investigation of H2 Production with in Situ CO2 and Sulfur Capture. 3.1.1. Chemicals, Sorbents, and Gases. The HTS catalyst was obtained from Su¨d-Chemie Inc., Louisville, KY, and consisted

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Figure 10. Simplified flow sheet of the bench-scale experimental setup.

of iron(III) oxide supported on chromium oxide. The CaO sorbent was obtained from a precipitated calcium carbonate (PCC) precursor that was synthesized from Ca(OH)2 obtained from Fisher Scientific (Pittsburgh, PA). The high-surface-area PCC [BET (Brunauer-Emmett-Teller) analysis: surface area (SA) ) 49.2 m2/g, pore volume (PV) ) 0.17 cm3/g] was synthesized using a dispersant-modified wet precipitation technique. The anionic dispersant used in this process was N40V, supplied by Ciba Specialty Chemicals (Basel, Switzerland). PCC was synthesized by bubbling CO2 through a slurry of hydrated lime. The neutralization of the positive surface charges on the CaCO3 nuclei by negatively charged N40V molecules formed CaCO3 particles characterized by a higher surface area/pore volume ratio and a predominantly mesoporous structure. Details of this synthesis procedure have been reported elsewhere.46,52 The feed gas for all H2 production tests was a mixture of 10% CO and 90% nitrogen (N2). The feed gas for the combined water-gas shift and CO2 and H2S removal tests contained 10% CO, 5000 ppm of H2S, and the balance nitrogen. 3.1.2. Bench-Scale Experimental Setup. Figure 10 shows the integrated experimental setup used for the bench-scale studies of the CLP. The bench-scale reactor was coupled to a set of continuous gas analyzers that detected concentrations of CO, CO2, H2S, CH4, and H2 in the product stream. The reactor setup was capable of handling high pressures and temperatures of up to 21 atm and 900 °C, respectively, which are representative of the conditions in a commercial syngas-to-H2 system. The mixture of gases from the cylinders was regulated and sent into the fixed-bed reactor by means of mass flow controllers capable of handling pressures of about 21 atm. From the mass flow controllers, the reactant gases flowed to the steam generating unit. The steam generating unit was maintained at a temperature of 200 °C and contained a packing of quartz chips that provided a large surface area of contact and mixing between the reactant gases and steam. The steam generating unit also served to preheat the reactant gases entering the reactor. The reactor, which was heated by a tube furnace, was provided with a pressure gauge and a thermocouple to monitor the pressure and temperature within. The reactor consisted of two concentric sections. The inner section, which was 25 in. long, had an inner diameter of 1 in. and was filled with solid particles consisting of only catalyst or only sorbent depending on the type of investigation. The outer section provided a preheating zone for the gases before they came into contact with the bed of solids. The packed-bed section of the reactor was detachable, which enabled easy removal and loading of the solid particles. The

reactant gases leaving the reactor entered a back-pressure regulator that built pressure by regulating the flow rate of the gases and was capable of building pressures of up to 68.9 atm. As shown in Figure 10, the inlet of the back-pressure regulator was connected to the reactor rod, and the outlet was connected to a heat exchanger. Because the entire section of the equipment setup upstream of the back-pressure regulator was exposed to high pressures, stainless steel lines were used to withstand the pressure and the reactor was constructed from stainless steel. The product gas mixture exiting the back-pressure regulator was then cooled in a heat exchanger using a chilled ethylene glycol-water mixture to condense the unconverted steam. The product gas at the exit of the heat exchanger was dried in a desiccant bed and sent to a set of continuous analyzers capable of determining the concentrations of CO, CO2, H2S, CH4, and H2 in the gas stream. 3.1.3. Water-Gas Shift Reaction in the Presence of HTS Catalyst. The extent of the water-gas shift reaction was determined in the presence of the HTS catalyst obtained from Su¨d-Chemie Inc. Catalyst particles were used in a fixed-bed reactor setup for all experiments. For each eperiment, 0.25 g of the catalyst was loaded into the reactor, and the pressure, temperature, and gas flow rates were adjusted for each run. The dry gas compositions at the outlet of the reactor were monitored continuously using the CO, CO2, H2S, CH4, and H2 gas analyzers. The total flow rate of the gases through the reactor was maintained constant at 725 sccm (standard cubic centimeters per minute) for all experiments, and the concentration of CO in the reaction mixture was maintained at 10.3%. The weight hourly space velocity (WHSV) used for the catalyst experiments was 174000 cm3 [standard temperature and pressure (STP)] per gram of catalyst per hour. 3.1.4. Simultaneous Water-Gas Shift and Carbonation. The combined water-gas shift and carbonation reaction was conducted using only CaO sorbent without the HTS catalyst. Five grams of CaCO3 sorbent was calcined by heating the sorbent to 700 °C in a stream of N2 until the CO2 analyzer confirmed the absence of CO2 in the outlet stream. The feed gas mixture was then introduced into the reactor, and the dry gas composition of the product gas was monitored continuously using the CO, CO2, H2S, CH4, and H2 gas analyzers. The total flow rate of the gases through the reactor was maintained constant at 725 sccm for all experiments, and the concentration of CO in the feed mixture was maintained at 10.3% with the balance nitrogen. The weight hourly space velocity (WHSV) used for the simultaneous water-gas shift and carbonation testing was 15535 cm3 (STP) per gram of sorbent per hour. 3.1.5. Combined H2 Production with H2S Removal. To study the effect of sulfur on the CLP, 5000 ppm of H2S was mixed with the CO, N2, and steam before the gas was sent to the reactor. The H2 production tests were conducted in the presence CaO sorbent as described in the preceding section. 3.2. Calciner: Bench-Scale Investigation of Effect of Realistic Calcination Conditions on Sorbent Reactivity. The effect of realistic calcination conditions on sorbent reactivity was investigated in a bench-scale rotary-bed calciner. The benchscale rotary-bed calciner is described in detail elsewhere;53 it consisted of a stainless steel reactor tube rotating within a horizontal furnace. The carrier gas, consisting of pure CO2 or a mixture of steam and CO2, was fed to the reactor, and the outlet of the reactor was connected to a CO2 analyzer. The sorbent was loaded in the reactor tube, and the temperature was increased to the calcination temperature. At the end of calcination, the CO2 capture capacity of the sorbent was determined

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Figure 11. CO conversion achieved in the presence of (a) HTS catalyst and (b) CaO sorbent (650 °C) at 21 atm.

in a thermogravimetric analyzer (TGA) apparatus from PerkinElmer. A detailed description of the testing equipment is available elsewhere.17 A small sample of the sorbent (15-20 mg) was placed in a quartz boat suspended from a platinum wire. The sorbent was brought to a reaction temperature of 650 °C in flowing nitrogen. Subsequently, the flow was switched to the reaction gas stream, which contained 10% CO2 with the balance N2. The instrument recorded the increase in sample weight with respect to time, which signifies the CO2 capture by the sorbent. The CO2 capture capacity of the sorbent (as a weight percentage) was then determined as the number of grams of CO2 captured per gram of CaO sorbent times 100%. Although the theoretical maximum capture capacity of pure CaO is 78.5%, this capture capacity is not achieved in actual testing because of the increase in molar volume during the conversion of CaO to CaCO3, leading to pore pluggage. 3.3. Hydrator: Bench-Scale Investigation of Effect of Hydration Conditions on Sorbent Reactivity. The effect on sorbent reactivity of sorbent hydration with low-temperature and atmospheric-pressure steam and with high-temperature and highpressure steam was investigated. Limestone sorbent was calcined under realistic calcination conditions at 1000 °C in a 100% CO2 atmosphere in the bench-scale rotary-bed calciner described earlier.53 At the end of calcination, the sorbent was hydrated. Water hydration was conducted by spraying water at 25 °C onto the sorbent with vigorous stirring. Steam hydration was conducted in the bench-scale reactor shown in Figure 10 in a 20% nitrogen/80% steam atmosphere. At the end of hydration, the CO2 capture capacity of the sorbent was determined by TGA. In the TGA, a small sample of the sorbent (15-20 mg) was placed in the quartz boat suspended from the platinum wire. The hydrated sorbent was brought to a reaction temperature of 650 °C in flowing nitrogen. Complete dehydration of the sorbent occurred by the time the sample had been heated to 650 °C. Subsequently, the flow was switched to the reaction gas stream containing 10% CO2 and the balance N2. 4. Results and Discussion 4.1. Carbonator: Bench-Scale Investigation of H2 Production with in Situ CO2 and Sulfur Capture. 4.1.1. Enhanced H2 Production with CO2 Capture. Figure 11a shows the CO conversion profiles for increasing reaction temperatures and S/C ratios at 21 atm in the presence of the HTS catalyst. The CO conversion was found to decrease with increasing temperature because of the thermodynamic limitations of the exothermic water-gas shift reaction. It can also be seen that, as expected, the conversion increased with increasing S/C ratio for all

Figure 12. Effect of S/C ratio on the (a) extent of H2S removal and (b) purity of H2 produced during the combined water-gas shift, carbonation, and sulfidation reaction in the presence of CaO sorbent (1 atm, 600 °C).

temperatures. Figure 11a shows that the maximum CO conversion achieved in the presence of the HTS catalyst even with the addition of excess steam was only 85%. To enhance CO conversion, the combined water-gas shift and carbonation reaction was carried out in the fixed-bed reactor containing the CaO sorbent alone, without a catalyst. Figure 11b shows the breakthrough curves in which high CO conversion was achieved in the prebreakthrough region through the in situ CO2 removal by the sorbent. As the sorbent was consumed, the breakthrough

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Figure 13. Effect of temperature on the (a) extent of H2S removal and (b) purity of H2 produced during the combined water-gas shift, carbonation, and sulfidation reaction in the presence of CaO sorbent (1 atm, S/C ratio of 1:1).

region occurred, followed by the postbreakthrough region, where all the active sorbent had been converted to CaCO3. Almost 100% CO conversion was achieved by the simultaneous water-gas shift and carbonation reaction conducted in the presence of the CaO sorbent. The in situ removal of CO2 during the water-gas shift reaction removed the thermodynamic constraint of the water-gas shift reaction and allowed 100% CO conversion to be achieved even at stoichiometric S/C ratios. Because the combined water-gas shift and carbonation was conducted at a high temperature of 600 °C, the HTS catalyst could be eliminated in the CLP. 4.1.2. Enhanced H2 Production with CO2 and Sulfur Capture. In the CLP, the CaO sorbent assumes the role of a multipollutant capture sorbent, in addition to enhancing the water-gas shift reaction. Hence, the influence of various process variables such as temperature, S/C ratio, and pressure on the purity of H2 produced and the extent of H2S removed during the combined water-gas shift, carbonation, and sulfidation reaction of CaO was determined. Figure 12 illustrates the effect of varying S/C ratio on the extent of H2S removal and the purity of H2 produced in the combined water-gas shift, carbonation, and sulfidation reaction. The extent of H2S removal by the CaO sorbent was found to increase with decreasing S/C ratio in the carbonation reactor. As shown in Figure 12a, the concentration of H2S in the product H2 stream decreased from 100 to