Calcium Looping Process for Enhanced Catalytic Hydrogen

1716

Ind. Eng. Chem. Res. 2011, 50, 1716–1729

Calcium Looping Process for Enhanced Catalytic Hydrogen Production with Integrated Carbon Dioxide and Sulfur Capture Shwetha Ramkumar, Mahesh V. Iyer, and Liang-Shih Fan* William G. Lowrie Department of Chemical and Biomolecular Engineering, The Ohio State UniVersity, 121 Koffolt Laboratories, 140 West 19th AVenue, Columbus, Ohio 43210

Enhancement in the production of high purity hydrogen (H2) from fuel gas, obtained from coal gasification, is limited by thermodynamics of the water gas shift reaction. However, this constraint can be overcome by the concurrent water gas shift and CaO carbonation reaction to enhance H2 production by incessantly driving the equilibrium-limited water gas shift reaction forward and in situ removing the carbon dioxide (CO2) product from the gas mixture. The in situ removal of CO2 is achieved by using a calcium oxide (CaO) sorbent which also reacts with and removes sulfur and halide contaminants present in the syngas stream. The water gas shift reaction is achieved by the high temperature shift (HTS) iron oxide catalyst while the CO2 capture is achieved using CaO sorbent. The spent sorbent from the system is regenerated by calcining it to produce a pure stream of CO2 and CaO which can be reused. The steam addition for the water gas shift reaction is reduced to a large extent in this process which aids in reducing the parasitic energy consumption. In addition, the extent of sulfur removal by the CaO sorbent is also enhanced by operating at lower steam partial pressures. Experiments conducted in a bench scale facility have revealed that high purity H2 of 99.7% purity can be produced by this calcium looping process. Introduction The total world energy demands for transportation, industry, electricity, etc. are projected to increase from 420 quadrillion BTU in 2003 to well over 720 quadrillion BTU by the year 2030.1 Fossil fuels which include petroleum, natural gas, and coal continue to meet these demands. However with the growing concerns over greenhouse gas emissions, the focus on the development of environmentally benign alternative technologies with high process efficiencies is gaining significance. Clean coal technologies that include CO2 capture is one of the major thrust areas. Coal gasification provides options for the conversion of syngas to liquid fuels, H2, or power via integrated gasification combined cycle routes. In a typical coal gasification system, the coal is fed along with steam and/or oxygen to the gasifier to produce syngas. The syngas is then cooled using a gas cooler or a water quench. The quench system also provides the excess steam required for the water gas shift reaction.2,3 While higher temperatures enhance the kinetics of the water gas shift reaction, the equilibrium limitation of the water gas shift reaction adversely affects H2 production, with the H2 yield falling with rising temperature. Hence, a high steam:carbon monoxide (S:C) ratio is required to enhance CO conversion and the consequent H2 yield. The S:C ratio required at 550 °C can be as high as 50 in a singlestage operation or 7.5 for a more expensive dual-stage process to obtain 99.5% pure H2.4 Numerous research studies have focused on the development of low temperature catalysts to improve H2 production.4 Commercially, the dual-stage sweet water gas shift reaction is carried out in series, with a high temperature shift (HTS; 300-450 °C) stage containing iron oxide catalyst and a low temperature shift (LTS; 180-270 °C) stage containing copper catalyst.5 The commercial iron oxide catalyst has a sulfur tolerance of about several hundred parts per million, while the copper catalyst has a lower tolerance to sulfur and chloride impurities.5 Hence syngas cleanup is required upstream of the shift reactors, which is achieved in conventional scrubbing towers using physical solvents * To whom correspondence should be addressed. Tel.: (614) 6883262. Fax: (614) 292-3769. E-mail: [email protected].

such as selexol or rectisol or chemical solvents such as aminebased solvents. This low temperature syngas cleanup process is energy intensive due to the gas cooling and reheating requirements. In a sour gas shift system, where the sulfur content of synthesis gas is greater than 1000 ppm, a sulfided catalyst is used in a series of reactors at a temperature of 250-500 °C and the desulfurization unit is located downstream of the water gas shift reactors.6,7 After the shift reaction, the syngas is subjected to scrubbing using solvents to remove the CO2 and is sent to the pressure swing absorber (PSA) unit to produce a pure stream of H2. The tail gas from the PSA unit is then used as fuel for power generation. Several methods to enhance the purity of H2 with the simultaneous separation of CO2 have been cited in the literature. A slight advancement in the commercial method of H2 production has been to remove the CO2 from the reaction mixture between the two stages of the shift reaction. However, solvents operate at low temperatures and this method involves severe energy penalties due to cooling and reheating of the reaction gas mixture. An effective technique to shift the water gas shift reaction to the right for enhanced H2 generation has been to remove H2 from the reaction mixture. This concept has led to the development of H2 separation membranes. Kreutz et al. have described the integration of these membranes in a commercial coal gasification unit.8 The syngas produced from the gasifier is shifted at a high temperature over a sulfur tolerant catalyst (STC) followed by a water gas shift H2 membrane reactor which aids in producing more H2 and separating it from the gas mixture.8 However, ceramic membranes have a very low H2 permeability, and the intermediate temperature composites in spite of having a high H2 flux are difficult to fabricate and are very susceptible to poisoning. The cermet membranes are superior to the other two classes of membranes, but again they are susceptible to poisoning and are expensive.9 Ma and Lund have reported the investigation of a Pd membrane reactor system packed with HTS catalyst.10 For optimum performance these reactors require two stages and an S:C ratio of 3. These reactors also suffer from inhibition effects of CO2, which reduces the yield of H2 from 90% to 50%.10 In addition,

10.1021/ie100347p  2011 American Chemical Society Published on Web 12/27/2010

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

1717

Figure 1. Schematic of the calcium looping process.

membranes cannot completely remove H2 from the mixture and suffer from a considerable pressure drop across them.9 Any remaining H2 in the main stream would dilute the CO2 and would lead to poor process economics. High temperature CO2 membranes have been developed which operate in the same temperature range as that of the water gas shift reaction. Although polymeric membranes for the removal of CO2 from H2 have been found to have several advantages such as simplicity of operation, high energy efficiency, and lower cost, most polymers have a poor H2/CO selectivity. Hence they are not very effective in shifting the equilibrium of the water gas shift reaction and producing high purity H2.11 An alternative concept to drive the water gas shift reaction forward has been to remove the CO2 from the reaction mixture using solid sorbents which either physisorb or react with the CO2 in the water gas shift reactor. The separation of CO2 from the reaction mixture at high temperatures removes the equilibrium constraint of the water gas shift reaction and enhances H2 production. Sorbents that operate at higher temperatures are beneficial to the process as the water gas shift reaction has superior kinetics due to the high temperature and enhanced thermodynamic extent due to the removal of CO2 from the product gas stream. Faster rates of reaction and larger CO2 capture capacities allow the use of smaller reactors and require a smaller amount of solids circulation through the system. CaO has a high CO2 capture capacity and removes CO2 to parts per million levels at a high temperature of 600 °C, making it one of the most suitable sorbents for this application.12 The concept of utilizing 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 coal13 and followed by CONSOL’s CO2 Acceptor process14 a century later when this concept was tested in a 40 t/day plant. A variation of this process, the HyPrRing process,15,16 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,17 Alstom,18 and GE19 processes. Shimizu et al.,20 Gupta and Fan,22 Iyer et al.,21 Sun et al.,23,24 Abanades et al.,25,26 and Manovic et al.27 have applied this concept

to the removal of CO2 from combustion flue gas. Brun-Tsekhovoi et al.,28 Fan et al.,29 Ortiz and Harrison,30 Han and Harrison,31 Johnsen et al.,32 Balasubramanian et al.,33 Hufton et al.,34 and Akiti et al.35 have 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. Most H2 production processes reported in the literature require a separate sulfur cleanup unit to prevent poisoning of the sorbent used for CO2 capture. Sulfur is present in syngas in the form of hydrogen sulfide (H2S) and carbonyl sulfide (COS). According to equilibrium calculations, at temperatures below 1300 K, all sulfur radicals combine to form predominantly H2S, which is close to 95% of the total sulfur content, and COS, which forms the other 5%.36 There have been studies conducted on the simultaneous calcination and sulfidation of calcium-based sorbents at temperatures higher than 600 °C.37 There have also been studies on the sulfidation of calcium carbonate (CaCO3) in the presence of CO2, but the CO2 was used only to maintain a high enough partial pressure to prevent the calcination of CaCO3.38-41 However, there is no mention of studies conducted on simultaneous CO2 and sulfur capture integrated with H2 production in the literature. This paper introduces the calcium looping process, a high purity H2 production process which combines H2 production with CO2, sulfur, and chloride capture from the syngas stream in a single stage.29,42,43 Calcium Looping Process (CLP). Figure 1 illustrates the reaction schemes in the two reactor modules of the calcium looping process. As shown in Figure 1, the CLP comprises two reactors: the carbonation reactor and the calciner. Detailed explanations of the H2 production and sorbent regeneration sections of the CLP are provided below. The carbonation reactor comprises either a fixed fluidized bed or an 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 electricity or steam. In the carbonation reactor, the thermodynamic constraint of the water gas shift reaction is overcome by the incessant removal of the CO2 product from

1718


Figure 2. Integration of the calcium looping process in a coal gasification system for the production of electricity, H2, and liquid fuels.

the reaction mixture, which enhances H2 production. This is achieved by 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 concentration of sulfur and halides in the outlet stream to parts per million levels. The in situ removal of CO2 removes the equilibrium limitation of the water gas shift reaction, thereby obviating the need for excess steam addition. Thermodynamic analysis, presented subsequently, predicts that the removal of H2S using CaO is inhibited by the presence of steam. Since almost all the steam is consumed in the enhanced water gas shift reaction, the removal of H2S is favored in the system. 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)

sulfur capture (COS): CaO + COS S CaS + CO2

(4)

halide capture (HCl): CaO + 2HCl S CaCl2 + H2O

(5)

The spent sorbent, consisting mainly of CaCO3, is regenerated back to CaO in the calciner. The calciner is operated at atmospheric pressure and at a temperature of 800-1000 °C in a rotary or fluidized bed system. The regenerated sorbent produced from the calciner is then conveyed back into the high pressure carbonation reactor through a lock hopper system. A mixture of CO2 and steam is used as the medium of calcination with the steam being condensed out at the exit of the calciner to produce a sequestration-ready CO2 stream. The heat can be

supplied directly or indirectly using a mixture of fuel and oxidant. The reaction occurring in the calciner is calcination: CaCO3 S CaO + CO2 (∆H ) 178 kJ/mol)

(6)

Calcium sorbents have been investigated for gas cleaning applications for several decades because of their ability to capture sulfur and CO2 readily, large availability, and low cost. A thorough investigation has been conducted on the pore structure requirements for achieving high conversions in calciumbased sorbents. Naturally occurring CaO forms two types of carbonates based on the type of space that it fills due to the increase in the molar volume. One type is formed in the micropores surrounding the micrograins, and the other is formed in the macropores as a product layer on the walls of the CaO grains. While the smaller pores surrounding the micrograins form the useful porosity, the macropores limit carbonation due to the product layer diffusion. With the increase in the number of cycles there is a loss in the microporosity and an increase in the macroporosity.44 Therefore, for achieving high conversions and retaining the reactivity of the sorbent over multiple cycles, the physical pore structure could be altered to increase the porosity in the microporous and the mesoporous ranges. We have developed a patented precipitated calcium carbonate (PCC) sorbent that has been found to have a suitable pore structure for the reaction with CO2.45,46 This highly reactive PCC is prepared by bubbling CO2 through a slurry of Ca(OH)2 containing surface modifying agents that lends PCC its unique surface and pore structure. It has been found that this sorbent has a mesoporous structure (15 nm) which prevents pore pluggage and gives a high conversion of 90%.12,46 Multicyclic carbonation and calcination tests conducted show that PCC sorbent attains a capture capacity of 40-36 wt % over 50-100 cycles, which is significantly higher than most of the high temperature sorbents reported.21 Figure 2 shows the integration of the calcium looping process in a typical coal gasification system with the cogeneration of electricity and H2. Coal or biomass is fed to the gasifier to


Keq2 )

Figure 3. Thermodynamic data illustrating the equilibrium constants of the water gas shift reaction and the combined water gas shift and carbonation reaction.

produce syngas. The syngas from the gasifier is then mixed with steam and sent to the carbonation reactor in which the water gas shift reaction goes to completion in the presence of the CaO sorbent and HTS catalyst. The CaO sorbent reacts with the CO2, sulfur, and halide impurities and removes them from the product stream. The product gas stream from the reactor contains predominantly H2 which is either purified further for ultrapure applications (e.g., fuel cells) or is sent to a combined cycle system for the generation of electricity. The H2 product stream may also be used for the production of fuels and chemicals. The spent sorbent from the carbonation reactor is then regenerated in the calciner where a sequestration-ready CO2 stream is produced. When calcination is conducted in the presence of steam, a CO2 stream containing a small concentration of H2S is produced from the calciner, which can then be sequestered as is.47 Since calcium sulfide (CaS) and calcium chloride (CaCl2) cannot be regenerated completely, a portion of the sorbent mixture is purged at the exit of the carbonator. Fresh sorbent makeup is added upstream of the calciner. The amount of purge and makeup will depend on the sulfur and chloride content of the coal syngas and the extent of sintering of the sorbent. The sorbent makeup and purge will result in the production of an H2 stream with constant purity and will prevent the accumulation of inert material (CaCl2 and CaS) in the circulating sorbent mixture. Thermodynamic Analysis of the Reaction Schemes in the Calcium Looping Process. The equilibrium constants for the water gas shift reaction and the combined water gas shift and carbonation reactions for various temperatures are shown in Figure 3. The equilibrium constants were obtained using HSC Chemistry version 5.0 (Outokumpu Research Oy, Finland). The equilibrium constant for the water gas shift reaction may be defined as Keq1 )

1719

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 an increase in the 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 increase the yield of H2 in the presence of an LTS catalyst. The addition of the carbonation reaction to the water gas shift reaction results in a significant increase in the equilibrium constant. Hence, the calcium looping process is capable of producing a much higher H2 yield and, hence, purity due to complete CO conversion, when compared to the conventional H2 production process. Equilibrium curves for the partial pressures of H2O (PH2O), CO2 (PCO2), and H2S (PH2S) as a function of temperature, for the hydration, carbonation, and sulfidation reactions with CaO were also obtained using HSC Chemistry version 5.0 (Outokumpu Research Oy, Finland). The relationship between reaction temperature and equilibrium partial pressure of CO2 and H2O for the carbonation and hydration reaction with CaO sorbent is shown in Figure 4. hydration:

CaO + H2O S Ca(OH)2

(8)

Carbonation and hydration of CaO are reversible reactions which occur depending on the conditions of temperature and partial pressures of CO2 and H2O, respectively. Carbonation of CaO occurs at conditions above the equilibrium PCO2 curve, while calcination of CaCO3 occurs at conditions below the curve. Similarly, hydration of CaO occurs above the PH2O curve while dehydration occurs at conditions below the curve. The typical composition of syngas from different gasifiers is shown in Table 1. Using the composition of the syngas, the feasibility of the simultaneous water gas shift, carbonation, and sulfidation reaction in the temperature range 500-750 °C was 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 and the equilibrium partial pressure of CO2 is low as shown in Figure 4. Steam is added to the syngas before it is fed to the carbonation reactor to adjust

PH2PCO2 PCOPH2O

where PCO2, PH2, PCO, and PH2O are the partial pressures of CO2, H2, CO, and H2O at equilibrium. The equilibrium constant for the combined water gas shift and carbonation reaction may be defined as shown by combined water gas shift and carbonation reaction: CO + H2O + CaO S H2 + CaCO3

(7)

Figure 4. Thermodynamic data for the hydration and carbonation of CaO sorbent.

1720


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

Table 2. Fuel Gas Composition Entering the Water Gas Shift Reactor After Steam Addition (S:C Ratio ) 1:1) (Adapted from Ref 48)

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

moving bed, dry


fluidized bed


entrained flow, dry

air subbituminous 20.1 2.97 3.98 2.53 2.97 6.58 0.99 0.03

oxygen bituminous 31.6 11.24 6.45 0.71 11.24 0.68 1.03 0.27

oxygen lignite 9.9 3.42 2.17 0.58 3.42 0.05 0.20 0.03



Table 3. Fuel Gas Composition Entering the Water Gas Shift Reactor After Steam Addition (S:C Ratio ) 3:1) (Adapted from Ref 48)

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

moving bed, dry


fluidized bed


entrained flow, dry

air subbituminous 20.1 2.29 3.07 1.95 6.88 5.08 0.76 0.03


oxygen lignite 9.9 2.02 1.28 0.34 6.06 0.03 0.12 0.02



the steam to carbon ratios (S:C) for the water gas shift reaction. Tables 2 and 3 list the compositions of the syngas-steam mixture for S:C ratios of 1:1 and 3:1, respectively. Depending on the thermodynamic extent of the water gas shift reaction occurring in the carbonator, the PCO2 in the carbonator can be determined. Figure 5a illustrates the PCO2 in the carbonator after the water gas shift reaction has occurred in the syngas-steam mixture with an S:C ratio of 1:1 from various gasifiers. It also shows the equilibrium PCO2 required for the carbonation reaction with CaO. It can be seen that the PCO2 in the carbonator for various gasifiers is higher than the equilibrium PCO2 for the carbonation of CaO; hence it can be inferred that, in the temperature range 500-750 °C, CO2 removal is achieved by CaO to greater than 90%. Figure 5b illustrates the PCO2 in the shifted syngas-steam mixture for an S:C ratio of 3:1. It can be seen that Figure 5b is very similar to Figure 5a and CO2 removal is achieved by CaO even for an S:C ratio of 3:1. In order to determine whether the CaO sorbent will undergo hydration, the PH2O in the carbonation reactor was determined after the combined water gas shift and carbonation reaction has occurred to equilibrium for S:C ratios of 1:1 and 3:1. Figure 6a shows the PH2O in the carbonator for an S:C ratio of 1:1, and since the PH2O in the carbonator is lower than equilibrium PH2O for the hydration of CaO, hydration will not occur for any syngas composition in the temperature range under consideration. Figure 6b shows the PH2O in the carbonator for an S:C ratio of 3:1, and it can be seen that hydration will occur at lower temperatures. Hydration of CaO will occur at temperatures below 600 °C for fluidized bed and moving bed (dry), 670 °C

for entrained flow (dry), 680 °C for moving bed (slagging), and 700 °C for entrained flow (slurry) gasifier syngas. For the reversible sulfidation of CaO, the extent of H2S removal will depend on the temperature and PH2O in the carbonator. 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 the increase in PH2O. At a temperature of 650 °C, the H2S concentration is 0.1 ppm for a PH2O of 0.01 atm and 1 ppm for a PH2O of 0.1 atm. By operating the carbonation reactor at the 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 will favor H2S removal using CaO at around 500-700 °C, which is a suitable temperature for the carbonation reaction as well. Materials and Methods Chemicals, Sorbents, and Gases. The HTS and STC catalysts were procured from Su¨d-Chemie Inc., Louisville, KY. The HTS catalyst consists of iron(III) oxide supported on chromium oxide, while the STC catalyst consists of cobaltmolybdenum on alumina support. PCC was synthesized from calcium hydroxide (Ca(OH)2) obtained from Fisher Scientific (Pittsburgh, PA). The high surface area PCC (BET analysis; surface area 49.2 m2/g; pore volume 0.17 cm3/g) was synthesized


Figure 5. Comparison of PCO2 in the carbonator with equilibrium PCO2 for the carbonation of CaO for S:C ratios of (a) 1:1 and (b) 3:1.

using a dispersant modified wet precipitation technique. The anionic dispersant used in this process was N40 V, 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 N40 V molecules forms CaCO3 particles characterized by a higher surface area/pore volume and a predominantly mesoporous structure. Details of this synthesis procedure have been reported elsewhere.12,49 The feed gas for all the H2 production tests was a mixture of 10% CO and 90% nitrogen (N2). Experimental Setup. Fixed Bed Reactor. Figure 8 shows the integrated experimental setup used for the bench scale studies of the CLP. The bench scale reactor is coupled with a set of continuous gas analyzers which detect concentrations of CO, CO2, H2S, CH4, and H2 in the product stream. The reactor setup is 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 is regulated and sent into the fixed bed reactor by means of mass flow controllers that can handle pressures of about 21 atm. From the mass flow

1721

Figure 6. Comparison of PH2O in the carbonator with equilibrium PH2O for the hydration of CaO for S:C ratios of (a) 1:1 and (b) 3:1.

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

controllers the reactant gases flow to the steam-generating unit. The steam-generating unit is maintained at a temperature of 200 °C and contains a packing of quartz chips which provide a large surface area of contact and mixing between the reactant gases and steam. The steam-generating unit not only facilitates the

1722


Figure 8. Simplified flow sheet of the bench scale experimental setup.

complete evaporation of the water being pumped into the steam generating unit, but it also serves to preheat the reactant gases entering the reactor. The reactor, which is heated by a tube furnace, is provided with a pressure gauge and a thermocouple to monitor the pressure and temperature within. The reactor consists of two concentric sections: the inner section is filled with the catalyst or sorbent-catalyst mixture and the outer section provides a preheating zone for the gases before they come in contact with the bed of solids. The sorbent and catalyst loading section of the reactor is detachable, which enables easy removal and loading of the sorbent. The reactant gases leaving the reactor enter a back-pressure regulator, which builds pressure by regulating the flow rate of the gases and is capable of building pressures of up to 68.9 atm. The back-pressure regulator is very sensitive, and the pressure within the reactor can be changed quickly without any fluctuations. In addition, the back-pressure regulator is also capable of maintaining a constant pressure for a long period of time. The valve seat material of the regulator is made of PEEK, which is corrosion resistant to acidic H2S vapors, which makes it suitable for conducting sulfur removal experiments. As shown in Figure 8, the inlet of the back-pressure regulator is connected to the reactor rod and the outlet is connected to a heat exchanger. Since the entire section of the equipment setup upstream of the back-pressure regulator will be exposed to high pressures, flexible stainless steel lines are used to withstand the pressure and the reactor is constructed from Inconel, which is resistant to corrosion due the high pressure, high temperature steam and H2S gas. The product gas mixture exiting the back-pressure regulator is 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 is dried in a desiccant bed and is sent to a set of continuous analyzers capable of determining the concentrations of CO, CO2, H2S, CH4, and H2 in the gas stream.

Water Gas Shift Reaction Testing. The water gas shift reaction was conducted using the catalysts obtained from Su¨dChemie. These experiments were conducted as baseline experiments to determine the conditions for maximum water gas shift catalytic activity at different ranges of temperatures (450-800 °C), S:C ratios, and pressures, which are beyond the commercial mode of operation but are of interest for the calcium looping process. Catalyst particles were used in a fixed bed reactor setup for all the experiments. The total flow rate of the gases through the reactor was maintained a constant at 725 sccm for all the experiments, and the concentration of carbon monoxide (CO) in the reaction mixture was maintained at 10.3%. A 0.25 g sample 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. Simultaneous Water Gas Shift and Carbonation. The combined water gas shift and carbonation reaction was conducted using a sorbent (CaO) to catalyst ratio of 10:1 by weight. The combined water gas shift and carbonation reaction experiments were conducted at 600, 650, and 700 °C with S:C ratios of 3:1, 2:1, and 1:1 at various pressures ranging from 1 to 21 atm. The CaCO3 sorbent is calcined by heating the sorbent-catalyst mixture to 700 °C in a stream of N2 until the CO2 analyzer confirms the absence of CO2 in the outlet stream. Catalyst Pretreatment. It is imperative to understand the HTS catalyst composition during calcination of the sorbent which occurs in the presence of a CO2 atmosphere at high temperature. Iron oxide occurs in three different phases: hematite (Fe2O3), magnetite (Fe3O4), and wustite (FeO). The active phase of the HTS catalyst is magnetite. However, in the presence of an oxidizing atmosphere, such as CO2 or steam, the magnetite phase is oxidized to hematite, which is likely during the calcination step. This is evident from the iron oxide phase diagram for a CO-CO2 system as illustrated in Figure 9a. Thus, a pretreatment procedure was developed which


1723

catalyst was conducted in the bench scale fixed bed reactor to determine the effects of temperature, pressure, and S:C ratio on the extent of reaction. Figure 10a shows the CO conversion profiles for increasing reaction temperatures and S:C ratios at ambient pressures. The CO conversion increases with increasing temperature as it approaches the equilibrium value at an optimal temperature (600-650 °C), beyond which it begins decreasing monotonically. At a pressure of 1 atm and an S:C ratio of 3:1 the conversion increases from 45.8% at 450 °C to 83.2% at 600 °C. Beyond 600 °C, the conversion decreases, and at 800 °C, it is 69.4%. This decrease with the increase in temperature is observed due to the thermodynamic limitation of the water gas shift reaction. Thus at lower temperatures although the equilibrium constant is high, the reaction rate is low. At high temperatures, although the reaction is very fast, the equilibrium constant is low. Consequently, maximum conversion is reached at an optimum temperature at which both the kinetics and the reaction equilibrium are favorable. From Figure 10a, it can also be seen, as expected, that the conversion increases with the increase in the S:C ratio for all temperatures. At a temperature of 650 °C, the conversion is 63.5% for an S:C ratio of 1:1, 71.6% for 2:1, and 80.28% for 3:1. As can be seen in Figure 10b, the effects of reaction temperatures and S:C ratios on CO conversion at 21 atm follow the same trend as that at 1 atm. In addition, below 600-650 °C, the CO conversion at 21 atm is greater than at 1 atm due to an increase in the rate of the reaction with increase in pressure. The observed partial pressure ratios were computed for different S:C ratios, temperatures, and pressures and were compared with the equilibrium values obtained from HSC Chemistry version 5.0 (Outokumpu Research Oy, Finland). The observed partial pressure ratio (Kobs) was computed from the experimental data and is defined as the ratio of the product of partial pressures of the products to that of the reactants as given by

Figure 9. Equilibrium phase diagrams for iron oxide systems for various (a) CO-CO2 and (b) H2-H2O gas compositions and temperatures (adapted from ref 50).

consists of treating the oxidized catalyst in a 20%/80% H2/H2O atmosphere at 600 °C which reduces the hematite to magnetite. Figure 9b shows that in a 20% H2 atmosphere the catalyst is in the magnetite form. The effectiveness of the pretreatment procedure was confirmed by X-ray diffraction analyses of the HTS catalyst before and after the pretreatment procedure. In the commercial deployment of the calcium looping process, pretreatment of the catalyst can be avoided by using a fixed fluidized bed reactor for the carbonation reactor in which the catalyst remains in the carbonation reactor while the CaO sorbent is looped between the carbonation reactor and the calciner. In this configuration the, catalyst is never exposed to oxidizing gases in the calciner. No deactivation of the STC catalyst was observed during calcination. Combined H2 Production with H2S Removal. To study the effect of sulfur on the calcium looping process, 5000 ppm H2S was mixed with CO, N2, and steam before being sent to the reactor. The hydrogen production tests were conducted in the presence of the catalyst and CaO sorbent. Results and Discussion Effect of Process Parameters on the Extent of the Water Gas Shift Reaction Using HTS Catalyst. An investigation of the water gas shift reaction in the presence of a HTS

Kobs )

PH2PCO2 PCOPH2O

As shown in Figure 11, it was found that each value of the observed partial pressure ratio (Kobs) was within the equilibrium value. From Figure 11 it can be seen that the partial pressure ratio increases with an increase in the temperature until it approaches equilibrium and then decreases along the equilibrium curve. Also, as the pressure increases, the system is closer to equilibrium for both S:C ratios of 1:1 and 3:1. This can be explained by the increase in the rate of the reaction with an increase in pressure. Enhancing the Water Gas Shift Reaction by in Situ CO2 Removal (HTS Catalyst and CaO Sorbent). From Figure 12 it can be observed that the CO conversion achieved in the presence of the HTS catalyst is only 80-90% even at a high pressure of 21 atm and a high S:C ratio of 3:1. At atmospheric pressure and a stoichiometric S:C ratio, a low CO conversion of 20-60% is obtained. In order to enhance the H2 yield, CaO sorbent could be introduced into the H2 production reactor for in situ CO2 removal from the reaction zone. This increase in H2 yield can be explained by LeChatelier’s principle, where the simultaneous CO2 removal drives the equilibrium limited water gas shift reaction forward. This concept was demonstrated by conducting the combined water gas shift and carbonation reaction in the presence of the calcined PCC sorbent and HTS catalyst in the fixed bed reactor. Figure 12a illustrates the typical breakthrough curves obtained during the combined water gas

1724


Figure 10. Effect of reaction temperature and S:C ratio on the conversion of CO by the water gas shift reaction in the presence of HTS catalyst at (a) 1 and (b) 21 atm.

Figure 11. Effects of reaction temperature and pressure on observed partial pressure ratio for the water gas shift reaction in the presence of HTS catalyst at S:C ratios of (a) 1:1 and (b) 3:1.

Figure 12. Typical curves for the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst depicting (a) gas composition (mol %) and (b) CO conversion (650 °C, 1 atm, S:C ratio of 3:1).

shift reaction and carbonation reaction for the N2 free dry product gas compositions. High purity H2 is produced in the prebreakthrough region due to in situ CO2 removal by the

sorbent. As the sorbent becomes exhausted, the breakthrough region occurs followed by the postbreakthrough region in which all the sorbent has been converted to CaCO3 and H2 production


1725

Figure 13. Effect of pressure on purity of H2 produced during the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst at S:C ratios of (a) 3:1 and (b) 1:1 (650 °C).

occurs in the presence of the HTS catalyst. The concentrations of CO and CO2 in the product gas mixture are very low in the prebreakthrough region and increase in the breakthrough region due to the depletion of the sorbent. Figure 12b illustrates the typical breakthrough curve obtained for CO conversion. Effect of Pressure. The effect of pressure on the combined water gas shift and carbonation reaction for S:C ratios of 3:1 and 1:1 are shown in Figure 13. The purity of H2 produced in the prebreakthrough region of the curves increases with the increase in pressure. From Figure 13a it can be observed that during the initial prebreakthrough period for a S:C ratio of 3:1, 95.6% H2 is produced at 1 atm, 99.7% pure H2 is obtained at 11 atm, and 99.8% pure H2 is produced at 21 atm. The extent of the prebreakthrough region, which signifies the extent of conversion of the CaO sorbent, also increases with the increase in pressure. A similar observation is made for a lower S:C ratio of 1:1 as shown in Figure 13b. It can be inferred that higher pressure results in increased partial pressure of CO2 which enhances the extent and rate of carbonation due to higher driving force. This consequently results in enhanced CO conversion, sorbent conversion, and H2 yield. Effect of S:C Ratio. The effect of S:C ratio on the CO conversion and H2 purity for the combined water gas shift and carbonation reaction is shown in Figure 14. The effect of S:C ratio at atmospheric pressure is shown in Figure 14a,b, while that at 21 atm is shown in Figure 14c,d. It can be seen that at atmospheric pressure a reduction in the S:C ratio results in a decrease in the CO conversion and associated H2 purity. However, at a higher pressure of 21 atm, almost 100% CO conversion and H2 purity is achieved for all three S:C ratios in the prebreakthrough region. This can again be attributed to the higher partial pressure of CO2 contributing to enhanced carbonation kinetics, which plays a key role in driving the water gas shift reaction to completion. Also, from a process design and cost perspective, operation at high pressures clearly illustrates the benefit of using a smaller amount of steam for a high CO conversion, resulting in cost savings. Effect of Temperature. Parts a and b of Figure 15 illustrate the effect of temperature on the combined reactions at 21 atm and S:C ratios of 3:1 and 1:1, respectively. At a high S:C ratio of 3:1, there is almost no change in the CO conversion with the change in temperature, as can be seen in Figure 15a. On decreasing the S:C ratio to the stoichiometric amount, it is observed in Figure 15b that temperature plays a significant role in the extent of CO conversion and a temperature of 600 °C is

optimum for achieving high CO conversions of 99.7%. Thus, from a process design perspective this defines the operating temperature for achieving high CO conversions and H2 yield while maintaining low steam requirements. Simultaneous Water Gas Shift, Carbonation, and Sulfidation Reaction Testing. Since syngas obtained from the gasifier contains 0.5-4% sulfur mostly in the form of H2S, the effect of sulfur and the extent of its removal by the CaO sorbent were determined on the combined water gas shift and carbonation reaction. Integrated H2 production and CO2 and H2S removal using calcium sorbent and HTS catalyst were investigated by the addition of 5000 ppm H2S to the fixed bed reactor feed. The calcium sorbent was used to simultaneously capture H2S and CO2 while enhancing H2 production in the presence of the HTS catalyst. As illustrated in Figure 16a, it was found that H2S concentration in the outlet H2 stream is reduced to a few parts per million in the prebreakthrough region by the reaction of H2S with the CaO sorbent. In the thermodynamics section of the sulfidation of CaO, illustrated in Figure 7, it was observed that the extent of H2S removal is inhibited by the presence of a high partial pressure of steam in the system. This concept is demonstrated in the experimental results depicted in Figures 16a and 17. Figure 16a illustrates the entire breakthrough curve of H2S concentration in the product H2 stream with the prebreakthrough and breakthrough regions. Figure 17 is a magnified image of the prebreakthrough region in Figure 16a, and it shows that with the increase in S:C ratio the H2S concentration in the H2 product increases. At a lower S:C ratio of 1:1, the H2S in the outlet stream is lower than 1 ppm, while at an S:C ratio of 3:1 the H2S concentration increases from 2 to 30 ppm in 750 s during the prebreakthrough region. At a S:C ratio of 1:1, in the prebreakthrough region, the carbonation reaction enhances the water gas shift reaction, which results in the consumption of most of the steam. Hence H2S removal by the calcium sorbent is enhanced and the H2S composition in the outlet stream is low. As the reaction proceeds, the CaO sorbent is consumed to form CaCO3 and CaS, resulting in the breakthrough curve seen in Figure 16a. Since the steam composition in the system is higher for an S:C ratio of 3:1, the H2S concentration in the product stream is higher. During the breakthrough region, H2S reacts with both CaO and CaCO3. The postbreakthrough region is not visible in Figure 16a as the H2S concentration in the product will keep increasing with time until all the CaCO3 is

1726


Figure 14. Effect of S:C ratio on the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst at 650 °C: (a) CO conversion at 1 atm, (b) H2 gas composition at 1 atm, (c) CO conversion at 21 atm, and (d) H2 gas composition at 21 atm.

Figure 15. Effect of temperature on CO conversion by the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst at 21 atm and S:C ratios of (a) 3:1 and (b) 1:1.

also converted to CaS. In the postbreakthrough region the H2S concentration in the product will be equal to the H2S concentration in the feed stream. Figure 16b illustrates the change in CO conversion with respect to time for S:C ratios of 3:1 and 1:1. In the prebreakthrough region, the CO conversion for an S:C ratio of 3:1 is slightly higher than that for 1:1. Effect of HTS and STC Catalysts on the Water Gas Shift Reaction. An STC catalyst procured from Su¨d-Chemie

was also tested for its suitability in the calcium looping process. Figure 18 depicts the comparison in CO conversion achieved at atmospheric pressure in the presence and absence of H2S in the inlet gas steam. It was found that, at 650 °C, the CO conversion decreases in the presence of H2S for both the HTS catalyst and the STC catalyst. It has been shown in the literature that the HTS catalyst still retains half its original activity in its sulfided form,51 and the same inference is obtained from Figure 18. Although the decrease in the conversion obtained in the


1727

Figure 16. Effect of S:C ratio on (a) the composition of H2S in the H2 stream and (b) CO conversion in the presence of the catalyst and sorbent during the simultaneous water gas shift, carbonation, and sulfidation reaction (600 °C, 1 atm).

Figure 17. Effect of S:C ratio on the composition of H2S in the H2 stream during the combined water gas shift, carbonation, and sulfidation reaction in the presence of CaO sorbent and HTS catalyst (600 °C, 1 atm).

presence of the STC catalyst is very low when compared to that in the HTS catalyst, it was found that even in the presence of H2S the HTS catalyst shows higher CO conversion at a temperature of 650 °C.

Figure 19. Comparison in the CO conversion obtained at different S:C ratios for different sorbent and catalyst mixtures (650 °C, 1 atm).

The enhancement in CO conversion on the addition of CaO sorbent to the STC catalyst is illustrated in Figure 19. At both S:C ratios of 3:1 and 1:1, the CO conversion was found to be the highest in the presence of the HTS catalyst and CaO sorbent. Although the CO conversion is increased by the addition of CaO to the STC, it is still lower than the conversion obtained in the presence of the mixture of HTS catalyst and CaO sorbent. Conclusions

Figure 18. Effect of type of catalyst and presence of H2S on CO conversion during the water gas shift reaction (650 °C, 1 atm).

Enhancement in the production of high purity H2 from syngas obtained from coal gasification systems can be achieved using CaO sorbent that can drive the equilibrium limited water gas shift forward by in situ removal of CO2. Detailed thermodynamic analyses for different gasifier systems show that the operating temperature window of 500-750 °C is suitable for the production of pure H2, for steam to carbon ratios of 1:1 to 3:1. However, operating at near-stoichiometric steam conditions is advantageous for simultaneous sulfur removal to low levels in the product H2 stream. Bench scale experimental data demonstrate that greater than 99% pure H2 can be produced at high temperatures and pressures. For near-stoichiometric conditions, high CO conversion and hydrogen purity can be obtained at high pressures and an optimal temperature of 600 °C. This

1728


operating temperature was also found to be favorable for simultaneous H2S removal to

Calcium Looping Process for Enhanced Catalytic Hydrogen

Recommend Documents