Coal-Direct Chemical Looping Gasification for Hydrogen Production

Article pubs.acs.org/EF

Coal-Direct Chemical Looping Gasification for Hydrogen Production: Reactor Modeling and Process Simulation Liang Zeng,† Feng He,‡ Fanxing Li,‡ and Liang-Shih Fan*,† †

William G. Lowrie Department of chemical and Biomolecular Engineering, The Ohio State University, Columbus, Ohio 43210, United States ‡ Department of Chemical & Biomolecular Engineering, North Carolina State University, Raleigh, North Carolina 27695, United States ABSTRACT: A novel process scheme for hydrogen production from coal with in situ CO2 capture, known as the coal-direct chemical looping (CDCL) gasification process, is discussed in this article. The CDCL process utilizes an iron oxide based oxygen carrier as a chemical looping medium to indirectly gasify coal into separate streams of H2 and CO2. ASPEN Plus reactor simulation models based on both thermodynamic equilibrium limitations and kinetic limitations are developed to analyze individual CDCL reactors. Process simulations are subsequently performed to estimate the performance of the CDCL process under various mass and energy management schemes. Reactor modeling results indicate that a moving bed reducer can effectively convert coal while reducing the oxygen carrier. The reduced oxygen carrier can in turn be oxidized by steam to produce hydrogen in a moving bed oxidizer. The fates of pollutants as well as the effects of various process operating parameters such as carbon and iron oxide conversions are also evaluated. Process simulation indicates that, even under a set of conservative assumptions, the CDCL process has the potential to convert coal to hydrogen at a thermal efficiency of nearly 78% (HHV) while capturing >90% CO2, which is 30% higher, on a relative basis, than conventional hydrogen generation processes.

1. INTRODUCTION Hydrogen, an important feedstock for oil refining and ammonia synthesis, has the potential to be an environmentally friendly fuel for transportation and power generation in the future.1 Hydrogen can be produced from carbonaceous fuels such as natural gas, coal, oil, and/or biomass. Due to its ample availability and relatively low cost, coal has the potential to be an attractive option among these energy sources.2 Hydrogen is already produced from coal through centralized gasification processes, in which coal is first partially oxidized into syngas. The syngas is then cleaned and converted into a concentrated H2 and CO2 mixture through the water−gas shift (WGS) reaction. This is followed by hydrogen separation and purification. Although well-established, the coal gasificationWGS process is more capital and carbon intensive and less efficient compared to the steam methane reforming process. In order to meet increasing demand for H2, innovative coal based hydrogen production processes with high energy conversion efficiency, low capital investment, and minimal pollutant and carbon emissions are highly desirable. The chemical looping strategy shown in Figure 1 offers a promising option to indirectly convert carbonaceous fuels using

a solid oxygen carrier in two stages: In the reduction stage, the carbonaceous fuel reduces the oxygen carrier. At the same time, the fuel is oxidized to a stream containing CO2 and steam. In the oxidation stage, the reduced metal oxide is regenerated with air and/or steam to produce heat and/or H2. A similar concept was applied to produce hydrogen more than 100 years ago in the so-called steam iron process.3 In the 1950s, Lewis and Gilliland adopted the chemical looping strategy to produce CO2 from carbonaceous fuels.4 More recently, chemical looping strategy was proposed to be used for efficient carbonaceous fuel combustion or gasification with integrated CO2 capture.5,6 During the last two decades, extensive research has been conducted with a focus on chemical looping combustion (CLC) using gaseous fuel, such as coal derived syngas or natural gas, especially with interest in oxygen carrier development,7−9 reactor design,10−12 and process synthesis.13,14 Solid fuel, such as coal, biomass, and municipal waste, has been also studied for direct conversion using chemical looping combustion technology.15,16 However, relatively few studies have been conducted on hydrogen production using chemical looping gasification (CLG) strategy. Svoboda et al. investigated the thermochemical properties of various metal oxides and concluded that oxides of iron can potentially be used for hydrogen production through CLG.17,18 Several process configurations have also been developed for cogeneration of hydrogen and electricity via chemical looping using gaseous fuels such as syngas or methane as the feedstock.19−21 These studies confirm the feasibility of using the chemical looping Received: February 29, 2012 Revised: April 24, 2012 Published: April 27, 2012

Figure 1. Simplified schematic of the chemical looping strategy for fuel conversion (MO/M represents a redox pair). © 2012 American Chemical Society

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Fe2O3 is recycled from the combustor. The reactions taking place in the reducer, which operates under 850−1050 °C and 1−30 atm, can be represented as

technology for hydrogen production. Detailed reactor and process analyses for CLG processes, especially for the direct conversion of coal, have not been performed to date. The coal-direct chemical looping (CDCL) gasification process is a promising process concept which efficiently converts coal into separate streams of hydrogen and readily sequestrable CO2.22 As shown in Figure 2, the chemical looping

3Fe2O3 + CO ↔ 2Fe3O4 + CO2

(1)

Fe3O4 + CO ↔ 3FeO + CO2

(2)

FeO + C ↔ Fe + CO

(3)

C + CO2 ↔ 2CO

(4)

For illustration purposes, coal is considered as pure carbon in the above reactions. In the bottom stage of the reducer, carbon in coal is first partially oxidized by CO2 and partially reduced iron oxides to form CO. The CO will then move up along the reactor to react with iron oxides to form CO2, which will in turn assist in coal gasification. During this process, Fe2O3 is gradually reduced to the form of FeO and Fe. The top stage of the reactor involves oxidation of gaseous species including both the gasified coal gas (CO and H2) and coal volatiles. To balance the heat required for the highly endothermic reactions 3 and 4, the oxygen carrier will be superheated before entering the reducer by combustion with air. Additionally, a small amount of steam/ CO2/H2 can be injected into the bottom of the reducer to enhance the reaction by promoting the coal/char gasification 4. The gaseous products from this reactor will mainly be CO2 mixed with a small amount of H2O, N2, and pollutants, where water can be removed by condensation to form a pure CO2 stream. The destinies of pollutants are discussed in section 4.1. 2.2. Oxidizer. The moving bed oxidizer produces hydrogen through the steam−iron reaction in a countercurrent contacting mode. The reduced Fe/FeO particles from the reducer are introduced at the top of the oxidizer and steam is injected at the bottom of the oxidizer. The oxidizer operates at around 600− 900 °C and 1−30 atm. Fe and FeO are oxidized to FeO and Fe3O4 while H2O is reduced to H2, as shown in reactions 5 and 6. High purity hydrogen can be obtained by condensing the steam and removing other trace pollutants. These reactions are mildly exothermic.

Figure 2. Simplified flow diagram for the coal-direct chemical looping process.

strategy is utilized to convert coal directly in the CDCL process. Unlike the conventional gasification processes where coal is gasified with air or oxygen, an iron oxide based composite oxygen carrier is used in the CDCL process for coal gasification/combustion. As a result, a concentrated CO2 stream can be obtained, obviating the CO2 separation step in the conventional coal to hydrogen processes. Coal conversion with iron oxide produces iron and iron oxides in reduced form, which can then react with steam to produce hydrogen or with air for heat and/or electricity generation. The present work evaluates the CDCL concept through reactor and process simulations using ASPEN Plus. In the following sections, the CDCL process is first described in detail. Then, simulation of individual CDCL reactors is performed considering both thermodynamic and kinetic limitations. This is followed by integration of these reactors using process simulation models to achieve optimized mass and energy management scheme. The effects of various operational parameters including coal and iron oxide conversions in the reducer, system operating pressure, and heat and power integration scheme to the overall process performance are also evaluated. The present study indicates that the CDCL process has the potential to be an efficient option for hydrogen generation from coal.

Fe + H 2O ↔ FeO + H 2

(5)

3FeO + H 2O ↔ Fe3O4 + H 2

(6)

2.3. Combustor. The combustor is operated at a turbulent fluidization regime. A stream of compressed air is used to regenerate the oxygen carrier particles. The compressed air also entrains these particles to the reducer through a riser, which connects the combustor outlet to the reducer inlet. In the combustor, oxygen carrier is fully reoxidized to Fe2O3 through reactions 7 and 8, which releases a significant amount of heat. As a result, combustor operating temperatures can reach 1050− 1250 °C. The operating pressure can range from 1 to 30 atm based on the type of the pressure changers for the inlet air. The hot solids from the combustor can compensate the heat needed in the reducer, while the high temperature flue gas can be utilized for steam and electricity generation.

2. CDCL PROCESS CONFIGURATIONS Figure 2 illustrates the schematic of a particular configuration of the CDCL process. There are three main reactors in the CDCL system, i.e. a moving bed reducer, a moving bed oxidizer, and an entrained bed combustor. Iron oxides circulate through the three reactors to participate in the key chemical looping reactions (reactions 1−8). Details regarding individual CDCL reactors are described in the following sections. 2.1. Reducer. The moving bed reducer converts the carbon in coal into concentrated CO2. The coal feedstock is introduced at the intermediate stage of the reducer while a Fe2O3 stream is fed from the top stage of the reducer. The majority of the

4Fe3O4 + O2 → 6Fe2O3

(7)

4FeO + O2 → 2Fe2O3

(8)

2.4. Energy Management in the CDCL System. When the looping system is considered as a single block, the mass inputs are coal, water, and oxygen, while the mass outputs are concentrated CO2 and H2. Within the chemical looping reactors, oxygen, chemical energy, and heat are transferred by 3681

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3. ASSUMPTIONS AND METHODOLOGIES Reactor modeling and process simulation are important for the development of novel energy conversion processes since accurate reactor and process modeling can reveal useful information regarding process performance and optimization strategies prior to cost intensive process demonstrations. In order to ensure the accuracy of the simulation, reasonable assumptions and rational simulation strategies need to be selected. The CDCL reactor system operates at high temperatures. Moreover, the key reactions are carried out in moving bed reactors which allow relatively long residence times for both solids and gas. It is therefore assumed in several simulation cases that the rates of most of the looping reactions are high enough to achieve thermodynamic equilibrium. In other cases, kinetic limitations of the slower reactions such as char gasification reactions are taken into account based on the experimental data obtained from moving bed reactor testing. The use of experimental data for coal and iron oxide conversions can help to obtain a more realistic and conservative estimate of the CDCL process performances. 3.1. ASPEN Plus Model Setup. Illinois #6 coal is used as the feedstock. The coal processing capacity of the plant is specified to be 132.65 tonnes/h (1000 MWth, HHV). Coal is first pulverized and sent to the reducer where it reacts with oxygen carrier particles consisting of iron oxide and an inert support (SiC in this case). Oxygen carrier particles circulate among the three looping reactors. In order to generate steam for the oxidizer, water is compressed by a pump and then vaporized using the process heat. Air to the combustor is supplied by a compressor or blower working at the combustor inlet pressure. The gaseous products from the process, i.e. CO2 from the reducer and H2 from oxidizer, are collected after compression for subsequent storage and transportation. The working fluids for power generation include 3-pressure level steams for the turbine system. Table 1 shows the specifications of the materials involved in the simulation.

the iron oxides. By elimination of iron oxides and summation from reactions 1−8, the overall CDCL reaction can be expressed as coal + x H 2O + yO2 → zCO2 + w H 2 (heat of reaction = ΔH )

(9)

The net heat ΔH of the system can be adjusted, theoretically, to any value by varying the ratio between H2O and O2 (x:y). In practice, the system is often operated to be exothermic so that the overall process is energetically and exergetically selfsustaining. It is also noted that the reducer reactor is highly endothermic. In order to maintain net heat generation from the process (ΔH < 0), a minimum H2O to O2 flow rate ratio (x:y) needs to be satisfied. This is achieved through one or more of the following approaches: (i) injection of substoichiometric amount of steam in the oxidizer; (ii) combusting a portion of the hydrogen product; (iii) partial bypassing of the oxidizer by sending a certain fraction of reduced iron-oxides from the reducer directly to the combustor. Figure 3 illustrates a schematic for the material and energy flows of configuration I in which a substoichiometric amount of steam is used to oxidize reduced iron oxides.

Table 1. Material Specifications in the CDCL process

Figure 3. Material flow and energy flow in the CDCL process.

Feedstock to the CDCL System Coal (Illinois #6)

The excess heat ΔH from the CDCL system can be recovered from the exhaust gas streams of the three CDCL reactors, while the remainder can be recovered by heat exchangers embedded in the CDCL reactors if desired. A significant amount of the heat will be carried by the flue gas from the combustor. Such high temperature flue gas, if produced at high pressure, can potentially be utilized in an expander for electricity generation. When the combustor is operated at low pressures, the resulting flue gas can be introduced to the heat recovery steam generation (HRSG) along with the two other exhaust gas streams from the reducer and oxidizer. The HRSG generates steam at different temperature and pressure levels. High temperature-high pressure steam goes to the steam turbine system for power generation, while a portion of the steam goes to the oxidizer for hydrogen production. The CDCL process is configured such that the electricity generated from the expander and turbines satisfy, at minimum, the parasitic energy needs. In the CDCL process, the chemical energy in coal is transferred to hydrogen, heat, and electricity products. Therefore, hydrogen yield tends to increase with decreasing net heat generation and decreasing net electricity generation Δe.

11.12% moisture, 9.70% ash, 63.75% C, 4.50% H, 1.25% N, 0.29% Cl, 2.51% S, 6.88% O, 0.0001% Hg by weight. HHV 27.113 MJ/kg 79% N2, 21% O2 by volume Working Media

air 3-level steam cycle oxygen carrier CO2 H2 fluegas ash sulfur chlorine mercury

3.1 MPa(HP)/0.45 MPa(IP)/0.01 MPa(LP) Fe2O3, SiC (inert) Process Output

>90% 15.3 MPa >99.99% 6 MPa, HHV 141.9 MJ/kg N2, CO2, NOx (99.9% coal to CO2 conversion. This corresponds to a maximum iron oxide conversion of 11.1%. To summarize, reactor simulations indicate that coal can

Rltd = 1.87

species

inlet flow (g/s)

outlet flow (g/s)

inlet flow (g/s)

C H O N Hg Cl S ash H2O CO CO2 H2 N2 HCl SO2 H2S Fe2O3 Fe0.947O FE FE0.877S Fe3O4

63.8 4.5 6.9 1.3 10−5 0.3 2.5 9.7 11.1 0.0 0.0 0.0 0.0 0.0 0.0 0.0 1170.0 0.0 0.0 0 0

0.0 0.0 0.0 0.0 10−5 0.0 0.0 9.7 51.0 0.0 234.1 0.0 1.3 0.3 0.3 0.0 0.0 661.4 306.8 5.9 0

63.8 4.5 6.9 1.3 10−5 0.3 2.5 9.7 11.1 0.0 0.0 0.0 0.0 0.0 0.0 0.0 1582.7 0.0 0.0 0 0

outlet flow (g/s) 3.2 0.0 0.0 0.0 10−5 0.0 0.0 9.7 51.3 0.0 221.9 0.0 1.3 0.3 0.3 0.0 0.0 1228.9 159.9 5.9 0

a

Coal feed rate is 100 g/s for this case. It decomposed into elements in the table. bPossible products such as COS with less than 1 × 10−3 g/s are neglected (except for Hg).

such as chlorine, sulfur, and mercury, which enter the reducer along with coal, are also included in Table 5. As can be seen, all the chlorine in coal is converted into hydrogen chlorine and exits from the reactor with the exhaust gas stream. Mercury also exits the reactor with the gaseous stream in elemental form. Although a small amount of sulfur will escape from the reducer in the form of SO2, 90+% of the sulfur at R = 1.38 and at R = 1.87 in coal will react with iron oxide, forming Fe0.877S. This solid sulfur compound will be carried to the oxidizer along with 3685

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Figure 8. Effects of (a) temperature on carbon conversions at 30 atm and (b) pressure on coal conversion at 900 °C with R = 1.38.

atom being removed from fully oxidized oxygen carrier. The results show that the injection of a small amount of CO2 or steam into the reducer does not lead to a drastic decrease in Fe2O3 conversion. It is therefore proposed that as-received coal is used without the additional drying step. As can be seen from the discussions above, the multistage model can assist in optimizing both the reactor design and operating conditions. The countercurrent gas solid flow pattern in the moving bed reducer design is thermodynamically favorable for fuel and iron oxide conversions. As will be illustrated in the next section, higher iron oxide conversion leads to higher hydrogen yield in the oxidizer. 4.2. Simulation of the Moving Bed Oxidizer. Similar to the reducer modeling, a five-stage equilibrium based model is used for the countercurrent moving bed oxidizer. Figure 10

the reduced iron oxide particles. No NOx is generated in the reducer. Effects of Temperature and Pressure. The effect of temperature is examined using sensitivity analysis with operating pressure fixed at 30 atm and R at 1.38. Using a similar approach, the effect of pressure is determined by fixing the operating temperature at 900 °C and R at 1.38. The results are shown in Figure 8. A higher reaction temperature and lower operating pressure favor the endothermic coal−Fe2O3 reaction. Figure 8a indicates that at 30 atm carbon cannot be fully gasified when the reducer is operated at temperatures notably below 900 °C. With experimental studies indicating that higher operating temperature leads to drastically improved reaction kinetics,6 it is recommended that the reducer be operated above 900 °C. According to Figure 8b, lower pressures favor coal conversion. An operating pressure below 31 atm can be suitable for the CDCL reducer at R equal to 1.38. Since high operating pressure leads to incomplete carbon conversion and low pressure translates into high energy consumption for CO2 and H2 compression steps, the standard reducer pressure is set at 30 atm in the current study. Effects of Steam and CO2 Enhancers. Steam and CO2 can enhance coal char gasification reactions and can therefore be injected at the bottom of the moving bed reducer to enhance the char conversion.6 However, an excessive amount of steam and/or CO2 can negatively affect the Fe2O3 conversion due to their capability of oxidizing Fe/FeO. Sensitivity analysis is performed at 900 °C and 30 atm with varying the feed ratio between moisture/steam/CO2 and carbon in coal. Figure 9 illustrates the effect of steam and CO2 on oxygen carrier conversion, which is defined by the fraction of active oxygen

Figure 10. Steam to hydrogen conversion for a moving bed oxidizer. Reactor operating conditions: 700 °C, 30 atm.

shows the performance of the oxidizer with pure steam and iron fed at 700 °C and 30 atm. With increasing the feed ratio between steam and iron, iron is gradually oxidized to wustite and then magnetite. At high steam to iron ratios, steam to hydrogen conversion goes down with increasing steam injection. The result shows that, at 700 °C, the moving bed oxidizer can fully convert iron to magnetite with a feed molar ratio of 1.9, and the steam conversion is as high as 70%. As a result of the countercurrent flow pattern, the moving bed oxidizer significantly reduces both the solid circulation rate and steam usage. As a result, the parasitic energy consumption of the process can be reduced. Thus, the moving bed oxidizer is adopted to improve system efficiency. Sensitivity analyses are also performed to study the effect of varying operating conditions such as temperature and compositions of the solid reactants in the moving bed oxidizer.

Figure 9. Effect of steam and CO2 on Fe2O3 conversion at 900 °C, 30 atm, and R = 1.38. 3686

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The results in Figure 11 show that, at 700 °C and 30 atm, the moving bed oxidizer remains its highest steam conversion as

These reactions enhance steam conversion. In the meantime, CO and CO2 produced from these reactions lead to a decreased H2 concentration in the product gas stream. 4.3. Process Simulation. Following the method described in section 3.3, the performance of the three cases under different operating conditions and process configurations are evaluated. Case I operates at 30 atm with all reactions reach their thermodynamic equilibrium except for coal, which is 99% converted. Case II is similar to case I except for the operating pressure, which is at 2 atm. By comparing cases I and II, the effect of operating pressure on the process efficiency can be obtained. Case III, which operates at 30 atm, accounts for potential kinetic limitations by assuming 95% coal conversion and 40% iron oxide conversion in the reducer. Case III evaluates the impact of kinetic limitations on the CDCL process efficiency. The process simulation results of all the three cases are shown in the Table 6. Table 7 summarizes the power balances of these scenarios.

Figure 11. Effect of temperature and feeding composition on steam conversion in the moving bed oxidizer.

Table 6. Process Simulation Results of Three Cases

long as the iron concentration exceeds 8% in the reduced iron oxides. The 8% iron concentration is equivalent to an oxygen carrier conversion of 36% in the reducer. When the oxygen carrier conversion from the reducer is less than 36%, the steam conversion will decrease. Hence, the operations of the reducer and the oxidizer are in close connection. The temperature effect on the oxidizer can also be studied from the modeling results. As shown in Figure 11, for the exothermic steam iron reaction (reactions 5 and 6), steam conversion decreases with increasing temperatures. Thus, it is preferable to operate the oxidizer at low temperatures when considering the steam usage. Since the steam−iron reaction has the same molar amount of reactant (steam) and product (hydrogen), the reaction is thermodynamically independent of operating pressure. This is also validated using the oxidizer model. In the case when the coal conversion in reducer is incomplete, a fraction of carbon in coal can be carried over from the reducer to the oxidizer, hence it would be helpful to quantify the effect of carbon carryover on the oxidizer performance. The effect of carbon and iron molar flow rate ratio on the oxidizer performance is illustrated in Figure 12.

Illinois #6 Coal, 1000 KWth, 132.65 tonne/h

Steam conversion is shown to increase with increasing the carbon/Fe ratio. This results from the steam carbon reaction and water−gas shift reaction with the presence of carbon: (10)

CO + H 2O = CO2 + H 2

(11)

case II

case III

3103 50% 288 507 56% 68% 99% 307 19.58 77.20

3103 50% 288 507 56% 68% 99% 307 18.1 71.36

3287 64% 277 535 40% 68% 95% 294 18.84 74.26

net power (MWe) ηe, % ηtot, %

24.9 2.49 79.69

0 0 71.36

29.8 2.98 77.24

Taking case I as an example, Tables 8 and 9 summarize the mass flow rates of the key solids and gaseous streams, respectively. The stream numbers are given in Figure 6. An oxygen carrier conversion of 56% with 99% coal conversion is assumed for the reducer, which operates at 900 °C, 30 atm. In order to obtain satisfactory steam to hydrogen conversion while limiting temperature swing between the reactors, an oxidizer temperature of 750 °C is selected. Under this operating temperature, the steam conversion in the oxidizer is 68%. The combustor is operated at 1170 °C. At such a temperature, the heat carried by the oxygen carrier is sufficient to offset the heat required in the reducer. Under case I, 19.6 tonnes/h of H2 and 307 tonnes/h of sequestrable CO2 are produced. This corresponds to 99% CO2 capture in coal. Fates of key pollutants from coal conversion, i.e. mercury, sulfur, and chlorine, were also analyzed using equilibrium based models. It was determined that all the mercury in coal will exit reducer in elemental form along with the CO2 stream. Elemental mercury can be removed by an activated carbon bed at low temperature. In the reducer, more than 90% sulfur reacts with iron oxide, forming Fe0.877S into solids circulating stream. The remaining sulfur exits the reducer in the form of SO2. Fe0.877S, which is carried over from the reducer to the oxidizer and combustor, are converted back to Fe2O3 by steam and air consecutively. As a result, H2S is generated in the oxidizer at a rate of 0.4 tonnes/

Figure 12. Effects of carbon on the product gas stream composition and steam conversion for the moving bed oxidizer. Reactor operating conditions: pure iron as the feedstock, 700 °C, 30 atm.

C + H 2O = CO + H 2

case I solid circulating rate (tonne/h) iron oxide loading wt % steam feed in oxidizer (tonne/h) air feed in combustor (tonne/h) iron oxide conversion in reducer H2O conversion in oxidizer carbon capture CO2 production rate (tonne/h) H2 production rate (tonne/h) ηH2, %

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Table 7. Power Balance over the CDCL Process input

output

unit operations power (MW)

air compressor

CO2 compressor

H2 compressor

steam turbine

expander/turbine

net power

case I case II case III

61.1 13.4 64.5

9.3 27.5 8.9

12.8 39.1 12.4

−48.1 −64.2 −53.3

−60.0 −15.8 −62.3

−24.9 0 −29.8

contain a significant amount of thermal energy. Combustor exhaust, which is at close to 30 atm and 1170 °C, is first introduced to an expander with a discharge pressure of 5 atm, generating 60 MW of power. A heat exchanger is then utilized to recover the remaining heat from the expander off gas while preheating the air for the combustor to 550 °C. After the heat exchanger, the flue gas, along with the high temperature gas streams from the oxidizer and reducer, is introduced to the HRSG. HRSG generates steam for the oxidizer and a 3-level steam turbine system. Here, 49.9 MW of electricity is generated from the steam turbines. In order to generate sequestrable CO2 and marketable H2 products, CO2 is compressed to 15.3 MPa while purified hydrogen produced from pressure swing adsorption (PSA) units is compressed to 6 MPa. Power consumptions for CO2 compression and H2 compression are 9.3 MW and 12.8 MW, respectively. In the present simulation, pressure drop in PSA is assumed to be 5 atm and H2 recovery ratio is 90%. Tail gas from PSA is recycled back to the combustor for heat generation. Under case I, nearly 77.2% of the energy in coal is transferred to hydrogen. When accounting for the net electric power produced, the overall process efficiency is 79.7%. In comparison, the energy conversion efficiency of the conventional coal to hydrogen process is roughly 58% with 90% CO2 capture.6 As can be seen, the CDCL process has the potential to be much more efficient than conventional coal gasification based approaches, especially under a carbon constrained scenario. Unlike case I which operates at evaluated pressure, case II is operated at a nominal pressure of 2 atm. The low operating pressure significantly reduces the needs for air compression. It also obviates need for an expander. Therefore, the capital investment for case II is expected to be much lower than that for case I. In the absence of the expander, the high temperature flue gas from the combustor is directly introduced to HRSG for reactant preheating and power generation. Even though the lower operating pressure in case II has little effect on the reactor performance, the heat and power integration strategy for case II is quite different from case I. The lower operating pressure of case II leads to low pressure hydrogen and CO2 from the oxidizer and combustor. As a result, more energy is required to drive hydrogen and steam compressors. Although the energy savings from the reduced power requirement in the air compressor (compared to case I) can offset the increased energy consumption for CO2 and H2 compression, the overall power that can be generated from the high temperature exhaust gas streams is significantly lower, resulting from the lower exhaust gas pressure and absence of expander. In order to generate adequate power to satisfy parasitic energy requirements, 1.48 tonnes/h of hydrogen is combusted to supplement the power generated from HRSG. The needs to combust hydrogen products for parasitic power generation lead to the lower H2 yield and hence lower process energy conversion efficiency for case II. Case II’s overall process efficiency is 71.4%, which is 8.3% lower than that for case I. As discussed earlier, the lower operating pressure for case II can lead to

Table 8. Circulating Oxygen Carrier Flows (tonnes/h) in the CDCL System for Case I stream

(1)

(2)

(3)

(4)

Fe Fe0.947O Fe3O4 Fe2O3 Fe0.877S support (SiC) ash temperature (°C) pressure (atm)

0 0 0 1552 0 1552 0 1170 29

402 883 0 0 8 1552 13 900 30

0 548 912 0 7 1552 13 750 31

0 0 0 1552 0 1552 13 1170 29

Table 9. Gaseous Product Flows (tonnes/h) in CDCL System for Case I stream H2O H2 CO2 CO O2 N2 H2S HCl SO2 SO3 NO NO2 temperature (°C) pressure (atm)

(5)

(6)

(7)

(8)

(9)

(10)

68.0 0 306.8 0 0 1.6 0 0.4 0.4 0 0 0 900

288.6 0 0 0 0 0 0 0 0 0 0 0 509

92.6 21.9 0 0 0 0 0.4 0 0 0 0 0 750

0 0 0 0 118.0 388.6 0 0 0 0 0 0 550

19.6 0 0 0 10.6 388.5 0 0 5.5 0 0.2 0 1170

19.6 0 0 0 10.6 388.5 0 0 5.5 0 0.2 0 767

30

31

30.5

31

30

5

h and SO2 is generated in the combustor at 5.6 tonnes/h. SO2 and H2S are removed using wet scrubber and Claus process, respectively. Thermal NO generated from the process is expected to be minimal due to the low combustor temperature and flameless combustion conditions. Near zero fuel NOx emissions are expected since nitrogen in coal is converted to nitrogen gas under the mildly oxidative environment of the reducer. All the chlorine in coal is converted to HCl in the reducer, at a rate of 0.4 tonnes/h. Under the proposed configuration, HCl is removed along with SO2 by the wet scrubber. Heat integration is essential for all the thermochemical processes. In the current study, heat exchangers, heat recovery steam generators (HRSGs), and optional high pressure gas expanders are configured to extract energy from the high temperature (and high pressure for cases I and III) exhaust gas streams produced from the CDCL reactors. The heat recovered is used to (1) preheat reactants such as air for the combustor and steam for the reducer; (2) generate power using steam turbine and/or expander. Again using case I as an example, the three product gas streams coming out of the combustor, reducer, and oxidizer 3688

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bed design is desirable for both the reducer and the oxidizer. It was further determined that the reducer favors higher operating temperatures and lower operating pressures. In contrast, higher steam to hydrogen conversion is expected for the oxidizer at lower temperatures. By integrating the individual reactor models, three CDCL process configurations under different operating conditions are evaluated. It was determined that the CDCL process favors higher operating pressure from an energy conversion standpoint. However, low pressure operation was found to have potential for notable capital cost savings. Current studies also indicate that, even under a kinetically limited scenario of 40% iron oxide conversion and 95% carbon conversion, the CDCL process has the potential to achieve an overall energy conversion efficiency of 77.2% with a 95% CO2 removal. All three cases investigated in the current study show significantly higher efficiencies than the conventional coal to hydrogen process. Moreover, pollutants in coal such as sulfur, chlorine, and mercury can be easily removed from the CDCL process. To conclude, the CDCL process simplifies the conventional coal to hydrogen scheme by eliminating the gasifier, WGS, ASU, and CO2 capture units. The process has the potential to be efficient and cost-effective, especially under a carbon constrained scenario.

significant saving in capital investment cost. At an efficiency of 71.36% (HHV), case II is still quite attractive comparing to traditional processes. Such efficiencies can be further improved if supercritical or ultrasupercritical steam cycles are used for power generation. These more efficient steam generation systems can significantly reduce the needs for combusting H2 product, thereby enhancing the H2 yield. The simulation results for case II further confirm the feasibility of the CDCL process. As in case I, the carbon oxide capture rate is 99% for case II. Case III assumes 40% iron oxide conversion and 95% carbon conversion in the reducer. It is further assumed that 4% of the solid carbon in coal is carried over to the oxidizer, whereas 1% of the carbon is entrained out of the system without getting converted. The carryover carbon is converted by steam to H2, CO, and CO2 in the oxidizer. As a result, the overall fuel gas (CO + H2) production rate from the oxidizer in case III is only slightly higher than the H2 production rate in case I, even though the iron oxide conversion in the reducer is much lower for case III. Lower iron oxide conversion in the reducer also leads to lower steam conversion. Since the hydrogen product from the oxidizer in case III is less pure, a larger amount of tail gas is generated from the PSA unit. The tail gas is combusted for steam/power generation. Overall, the process efficiency for case III is 77.2% corresponding to a roughly 2% energy penalty compared to case I. It can therefore be concluded that the overall CDCL process efficiency is not likely to be significantly affected by low iron oxide and carbon conversion. In fact, incomplete coal conversion and carbon carryover can reduce the heat requirements for the reducer. Furthermore, the gaseous product from the oxidizer still consists of concentrated hydrogen and can be easily purified by PSA. The carbon capture rate for case III is 95%. Figure 13 summarizes the energy conversion efficiencies for all three cases of the CDCL process. The energy conversion

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AUTHOR INFORMATION

Corresponding Author

*Telephone: +1 (614) 688-3262. Fax: +1 (614) 292-3769. Email: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS

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REFERENCES

This work was supported by the Ohio Coal Development Office of the Ohio Air Quality Development Authority (Project CDO/D-08-02), Ohio Department of Development (Project TECH 08-062), U.S. Department of Energy (Project DEFC26-07NT43059), The Ohio State University, and an industrial consortium.

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Figure 13. Efficiency comparisons for the various CDCL cases. The dashed line shows traditional hydrogen production process efficiency.

efficiency for conventional coal to hydrogen process with 90% CO2 capture is also shown in Figure 13. All three CDCL cases show significantly higher efficiencies than the conventional coal to hydrogen process. The higher process efficiencies of the CDCL process results from intensified and more advantageous energy integration schemes.

5. CONCLUSIONS The present study evaluates the feasibility and performance of a novel coal direct chemical looping process. Reactor models using ASPEN Plus are developed for reactor simulation and optimization. Reactor simulation results indicate that a moving 3689

dx.doi.org/10.1021/ef3003685 | Energy Fuels 2012, 26, 3680−3690

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dx.doi.org/10.1021/ef3003685 | Energy Fuels 2012, 26, 3680−3690