Biomass direct chemical looping process: Process simulation

Fuel 89 (2010) 3773–3784

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Biomass direct chemical looping process: Process simulation Fanxing Li, Liang Zeng, Liang-Shih Fan * William G. Lowrie Department of Chemical and Biomolecular Engineering, The Ohio State University, Columbus, OH 43210, USA

a r t i c l e

i n f o

Article history: Received 10 August 2009 Received in revised form 22 June 2010 Accepted 14 July 2010 Available online 30 July 2010 Keywords: Chemical looping Biomass Moving bed reactor modeling Process simulation

a b s t r a c t Biomass is a clean and renewable energy source. The efficiency for biomass conversion using conventional fuel conversion techniques, however, is constrained by the relatively low energy density and high moisture content of biomass. This study presents the biomass direct chemical looping (BDCL) process, an alternative process, which has the potential to thermochemically convert biomass to hydrogen and/or electricity with high efficiency. Process simulation and analysis are conducted to illustrate the individual reactor performance and the overall mass and energy management scheme of the BDCL process. A multistage model is developed based on ASPEN PlusÒ to account for the performance of the moving bed reactors considering the reaction equilibriums. The optimum operating conditions for the reactors are also determined. Process simulation utilizing ASPEN PlusÒ is then performed based on the reactor performance data obtained from the multistage model. The simulation results indicate that the BDCL process is significantly more efficient than conventional biomass conversion processes. Moreover, concentrated CO2, produced from the BDCL process is readily sequesterable, making the process carbon negative. Several BDCL configurations are investigated for process optimization purposes. The fates of contaminants are also examined. Ó 2010 Elsevier Ltd. All rights reserved.

enzymes

1. Introduction

6nCO2 þ 5nH2 O ! ðC6 H10 O5 Þn þ 6nO2 þ DH

Biomass is defined as ‘‘any organic matter, which is available on a renewable basis, including agricultural crops and agricultural wastes and residues, wood and wood wastes and residues, animal wastes, municipal wastes, and aquatic plants” [1]. This renewable resource can be converted into electricity, H2, chemicals, and liquid fuels. Compared to fossil fuel, biomass conversion is less carbon and pollutant intensive. Moreover, it is widely distributed and abundantly available. For instance, the annual biomass potential from forest and agricultural resources alone is over 1.3 billion dry tonne in the United States, which is sufficient to displace nearly 40% of the petroleum consumed in the nation [2]. Therefore, biomass has the potential to be a favorable energy source to, at least partially, address the increasing environmental and energy sustainability concerns. Most of the carbon in the atmosphere circulates in the form of CO2. Through photosynthesis, as schematically shown in Reaction (1), green plants convert water and carbon dioxide into glucose, starch, cellulose, hemicellulose, lignin, etc. From this reaction, a portion of the solar energy is chemically stored in these organisms. The stored energy can be then transferred into various forms through the food chain consumption scheme and anthropogenic activities.

where n denotes the degree of polymerization. Historically, biomass has long served as the primary source of thermal energy for heating and cooking. However, its relatively low energy density and wide geographical distribution requires effective strategies to ensure the techno-economical viability for biomass conversion in the modern era. Biomass conversion technology can be classified into either biochemical or thermochemical processes [3]. Biochemical processes involve the use of microorganisms to convert biomass into ethanol or other valuable products. However, such processes consume a large amount of water [4]. Consequently, energy intensive distillation units are required for product upgrading. Moreover, biochemical techniques have stringent requirements on the properties of the feedstock. Starch is relatively easy to convert; however, starch based fermentation processes compete against human food supplies, inflating the cost of agricultural products. Cellulosic biomass, although is not edible, is rather difficult to convert through a biochemical approach. Compared to biochemical processes, thermochemical processes have less stringent requirements on the feedstock properties and consume less water. Several thermochemical systems for biomass conversion are currently under development. They include combustion, gasification, pyrolysis, and chemical looping. Conventional combustion processes fully oxidize the biomass with air to produce heat and/or electricity. Gasification processes partially oxidize the biomass with pure oxygen or air at high temperatures to produce

* Corresponding author. Tel.: +1 614 688 3262; fax: +1 614 292 3769. E-mail address: [email protected] (L.-S. Fan). 0016-2361/$ - see front matter Ó 2010 Elsevier Ltd. All rights reserved. doi:10.1016/j.fuel.2010.07.018

ð1Þ

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Notation x

a

xFe2 O3 M

normalized oxygen content of the reducer product, i.e. the number of transferable oxygen atoms/the number of Fe atom in reduced oxygen carrier from the reducer split ratio, the fraction of the particles that are directly introduced to the combustor from the reducer 16ð3=2xÞ Transferable oxygen weight ¼ 56þ163=2 ¼ 32x 12 Oxygen carrier weight ðwithout considering inertÞ mass percentage of moisture in dried biomass (%)

synthesis gas (syngas). This gaseous fuel can then be processed into a variety of products including electricity, hydrogen, liquid fuels, and chemicals. Both combustion and gasification processes are exothermic and are relatively inefficient due to the inevitable exergy loss in the combustion or gasification step [5]. Biomass can also be converted into gases, bio oils, tars and char through pyrolysis at elevated temperatures in an oxygen free environment. The endothermic nature of the pyrolysis reactions requires effective heat integration strategies, which are yet to be developed. The energy conversion efficiencies of several thermochemical processes are listed in Table 1 [6]. As can be seen, the conventional technologies are relatively inefficient due to both the technical limitations and low heating value of the biomass feedstock. The relatively low process efficiencies for biomass conversion, compounded with the high biomass collection and transportation costs, significantly affect the economic feasibility of these energy conversion systems. Therefore, efficient biomass conversion strategies, preferably those applicable at small scales, are desirable. Direct chemical looping provides a promising method to effectively utilize the potentials of biomass [7,8]. The chemical looping technology [9], illustrated in Fig. 1, converts carbonaceous fuels using metal-oxide based oxygen carriers circulating between two main reactors, a reducer and an oxidizer. Within the redox loop, oxygen transfers from the metal oxides to fuel in the reducer and from steam or air to reduced metal oxides in the oxidizer. This strategy avoids direct contact between fuel

Table 1 Energy efficiencies of selected thermochemical processes. Process

Products

Efficiency (%)

Direct combustion Pyrolysis-flash Oxygen-blown gasification

Electricity Pyrolytic oil, gas, char Fuel gas

25 37–45 50–60

CO2, H2O…

H2/N2+Power M

Reducer

Oxidizer MO

Fuel

H2O/Air

Fig. 1. Simplified schematic of the chemical looping concept (MO represents metal oxides, and M is its reduced form).

R C E

the molar ratio between fresh Fe2O3 in oxygen carrier and carbon in biomass introduced to the reducer oxygen carrier conversion, 1  2x/3 ratio between the reducer heat requirement when the moisture content in the biomass is M and the heat requirement when 5 wt% moisture biomass is the feed

and oxidants. Hence, a separate CO2 product stream is generated with minimal energy consumption. Chemical looping processes using gaseous feedstock have been extensively studied in the past decade. Experimental studies from the lab to small pilot (120 kW) scales are being performed to demonstrate the feasibility of the chemical looping concept. The ongoing studies focus on oxygen carrier particle performance, reaction chemistry, solids handling, and process system design and testing [10–16]. Complementary to experimental study, theoretical investigations, such as thermodynamic analysis and process simulation, have been performed to evaluate the feasibility of the chemical looping strategy from the overall process standpoint. The thermodynamic investigations on chemical looping combustion (CLC) have been performed by Richter [17] and McGlashan [18]. Both studies confirm the potential of this processing strategy in increasing the efficiencies for power generation. For its application in hydrogen production, Xiang [19] and Cleeton [20] suggested several process configurations using syngas derived from coal gasification. Other types of gaseous feedstock such as methane [21] are also proposed as fuel supplies for the chemical looping process to produce hydrogen. Compared to gaseous fuels, the direct conversion of solid fuels such as biomass using chemical looping strategy is a relatively new subject. Pan et al. [22] studied the reaction between copper based oxygen carrier particles and woody biomass using thermogravimetric analyzers and concluded that the oxygen carrier enhances the conversion of biomass. However, the application of copper based oxygen carrier is limited by its low melting point and the inability to react with steam for hydrogen generation. Shen et al. [23] and Leion et al. [24] tested biomass conversion using iron based oxygen carrier particles in circulating fluidized bed reactor systems. Both the fuel conversion and CO2 concentration in the reducer product gas, however, are significantly lower than those in gaseous fuels CLC systems. In addition, both systems are limited to electricity generations. Fan et al. proposed a novel direct chemical looping process [8,9,25] that directly converts solid fuels such as coal and biomass into hydrogen and/or electricity using iron based oxygen carrier particles and moving bed reactor system. Gnanapragasam et al. performed a process analysis on the direct chemical looping system, showing that the process is efficient for hydrogen production [26]. The limitations in gas and solid conversions due to thermodynamic equilibrium, however, are not taken into account in this analysis. This study conducts a comprehensive process simulation using biomass as the feedstock in order to ascertain its DCL process application. The paper first provides an overview on the biomass direct chemical looping (BDCL) process and its mass and energy flow schemes. A multistage model is then developed to study the thermodynamic equilibrium in the BDCL reactors. The multistage model assists in predicting the moving bed reactor performance and in optimizing the operating conditions. It is followed by overall process simulation using the results obtained from the multistage reactor modeling. Finally, the feasibility of the BDCL process is evaluated.

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Fig. 2. Material flow and energy flow in the biomass direct chemical looping (BDCL) process.

2. BDCL process overview As illustrated in Fig. 2, the BDCL process converts biomass, generalized as (C6H10O5)n, to electricity and hydrogen using three reactors, i.e. the Reducer, Oxidizer, and Combustor. Each reactor performs a unique function in the chemical looping scheme. The three reactors closely interact with one another via the circulation of iron oxide based oxygen carrier particles, which is an essential carrier for both chemical and thermal energy.

2.1. Reducer The reducer is a moving bed reactor where the oxygen carrier particles, in Fe2O3 form, react with biomass feed to produce reduced Fe/FeO solids and a gaseous CO2/H2O stream. Here, the Fe2O3 particles are fed from the top while the dried biomass is injected into the middle section of the reactor. Additionally, a small amount of gasifying agents such as steam, CO2, and/or H2 can be

used to enhance the reaction kinetics by gasifying the solid fuel. The reducer operates at 850–950 °C and atmospheric pressure. The overall reaction occurring in this reactor can be represented as:

ðC6 H10 O5 Þn þ

24n gasifying agents FeO1:5 ! 6nCO2 þ 5nH2 O 3  2x 24n FeOx þ DH1 þ 3  2x

ð2Þ

where x denotes the oxidation state of iron oxide in the product, 0 6 x 6 3/2, for Fe2O3, x = 3/2; Fe3O4, x = 4/3; FeO, x = 1; Fe, x = 0. A detailed reaction scheme and the reactor design are provided in Fig. 3. Initially, thermal decomposition takes place as the biomass is exposed to the high temperature environment of the reducer generating char, tar, and pyrolysis gases such as CO2, H2O, CO and H2. The gasifying agents facilitate interactions between char/ tar and the oxygen carrier particles. Specifically, the oxidative products, CO2 and H2O, instigate the char/tar gasification into reductive gases, such as CO and H2. The gaseous fuels are then oxidized by the iron oxides. As the reducer is a moving bed, the lower

Fig. 3. Gas solid contacting pattern of the reducer.

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section of the reactor is in a reducing environment due to the presence of excess char. To compare, the upper portion where the fresh Fe2O3 first enters the reducer is under a more oxidizing environment. The gasifying agents, or promoter gases, are injected from the bottom of the lower section to gasify the remaining char, producing gaseous fuels that travel upward through the vessel to react with iron oxides (mainly FeO and Fe3O4). The reactions between the fuel and the oxygen carriers generate reduced Fe/FeO particles and oxidative gases that further ascend to assist in gasifying the solid fuel above. Prior to exiting this reactor, the gas mixture reacts with Fe2O3/Fe3O4 solids producing a concentrated CO2 and H2O outlet stream. As can be seen in Fig. 3, the upper portion of the reducer uses the higher oxidation state iron oxides to fully convert the gaseous phase into CO2 and H2O, while the lower portion (partially) oxidizes or gasifies the biomass char with lower oxidation state iron oxides. Full biomass conversion is desired to improve the energy conversion and prevent any unconverted solid fuels from contaminating the ensuing reactors. In the BDCL process, all the carbon in the biomass feedstock leaves the reducer in the form of CO2 as illustrated in Reaction (2). After heat recovery and gas cleanup, the gaseous product can be directly vented into the atmosphere with close to zero net CO2 emission. Alternatively, by condensing the steam from the outlet gases, the concentrated CO2 stream produced can be compressed and sequestered resulting in a carbon negative process from the life cycle standpoint. Another issue to be considered is the highly endothermic nature of the reducer. To compensate for the heat required, the oxygen carrier particles are heated to a higher temperature in the combustor before entering the reducer. The overall heat integration scheme for the BDCL process is given in Section 2.4. 2.2. Oxidizer A portion (a) of the reduced oxygen carrier particles from the reducer outlet are transferred to the oxidizer for hydrogen production. The remaining reduced oxygen carrier particles are directly combusted with air in the combustor for heat generation. The oxidizer is a countercurrent moving bed reactor that produces hydrogen via the steam–iron reaction at 700–900 °C and 30 atm. As shown in Reaction (3), the amount of hydrogen produced can be adjusted by varying a. In this reactor, FeOx is oxidized to Fe3O4 as it moves downwards. Meanwhile, H2O is reduced to H2 as it travels upwards. By condensing steam, a high purity hydrogen stream can be produced. The steam–iron reaction is mildly exothermic.

24an 8ð4  3xÞan 24an FeOx þ H2 O ! FeO1:33 3  2x 3  2x 3  2x 8ð4  3xÞan þ H2 þ DH2 3  2x

ð3Þ

where a denotes the percentage of FeOx used for hydrogen production, 0 6 a 6 1. 2.3. Combustor The combustor is an entrained flow reactor through which the oxygen carrier particles, from either the reducer or the oxidizer, are conveyed back to the reducer using compressed air. During this step, the chemical intermediate is re-oxidized to Fe2O3 via Reactions (4) and (5) releasing a significant amount of heat. At adiabatic conditions, the reactor can reach above 1100 °C. The temperature within can be moderated by adjusting the amount of excess air or the support content in the oxygen carrier particle. In most cases, the heat carried by the solids from the combustor can fully compensate for the thermal energy required in the reducer. The flue

gas from the combustor, consisting mainly of nitrogen, can also be used to drive expanders for electricity generation.

24ð1  aÞn 24ð1  aÞn FeOx þ 6ð1  aÞnO2 ! FeO1:5 þ DH3 3  2x 3  2x

ð4Þ

24an 24an 24an FeO1:33 þ O2 ! FeO1:5 þ DH4 3  2x 3  2x 3  2x

ð5Þ

2.4. Material and energy management in the BDCL system In the reducer, Fe2O3 is reduced to FeOx. The normalized oxygen content, x, is inversely related to the effective oxygen carrying capacity of the particle, x. Larger values of x, or lower values of x, correspond to higher solids circulation rates (24n/(3  2x)). A fluidized bed reducer design can only achieve full fuel conversion into CO2 and H2O when x is between 4/3 and 3/2. The results translate to a x value of less than 0.03. Because Fe3O4 and Fe2O3 are thermodynamically incapable of converting H2O into H2, x should be less than 4/3 if hydrogen is the desired product. A moving bed reducer design is an effective alternative to achieve a higher solids conversion. The most desirable case is to obtain metallic iron (x = 0) from the reducer minimizing the overall iron oxide circulation rate. Further details on the advantages of moving bed reactors in the BDCL system and the possibility of achieving near full Fe2O3 conversion are discussed in Section 4.1. In the oxidizer, Fe3O4 is the final solid product as depicted in Reaction (3). When maximizing the hydrogen production, a low x and high a value are preferred. Quantitatively, a(8  6x)/(9  6x) of the transferred oxygen is used for hydrogen production. It should be noted that the present analysis for the oxidizer is based on reaction stoichiometry, which does not consider the equilibrium limitations. From the thermodynamic equilibrium standpoint, the value of x from the reducer affects both the steam to hydrogen conversion and oxygen utilization. Additionally, a low operating temperature thermodynamically favors steam to hydrogen conversion in the oxidizer. The validity of these analyses is discussed in Section 4.2. As stated in the previous section, the temperature of the oxygen carrier particles leaving the combustor should be high enough to compensate for the heat requirement in the reducer. A lower oxidizer operating temperature, though thermodynamically favorable for the reactor, may not be desirable for the overall process since a larger amount of heat needs to be generated in the combustor to offset the temperature gradient between the oxidizer outlet/combustor inlet and combustor outlet/reducer inlet. Further analysis is conducted to determine the optimum operating conditions for the oxidizer considering all the aforementioned constraints. Reaction (6) given below, obtained by summation of Reactions (2)–(5), represents the overall reaction of the process.

ð10x þ 32a  15  24axÞn H2 O 3  2x ð18x þ 12ax  12x  16aÞn FeOx 8ð4  3xÞan O2 ! 6nCO2 þ H2 þ DH0 þ 3  2x 3  2x ð6Þ

ðC6 H10 O5 Þn þ

DH 0 ¼

4 X

DHi

i¼1

The first law of thermodynamics dictates that DH0 for a self-sustaining BDCL system (i.e. no external heat input required) needs to be smaller than or equal to zero. Fig. 2 schematically summarizes the material and energy flow in the BDCL process. The high temperature oxygen carriers can effectively transfer both oxygen and heat. Furthermore, the heat released from the chemical looping

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system can be utilized for supplementary heating and electricity generation to compensate for parasitic energy consumptions. When combining Reaction (6) with the atmospheric CO2 circulation in Reaction (1), Reaction (7) becomes the same as water photolysis, CO2 8ð4  3xÞan 8ð4  3xÞan 4ð4  3xÞan H2 O ! H2 þ O2 þ DH þ DH0 3  2x 3  2x 3  2x ð7Þ

From this reaction scheme, the overall thermal input, DH, is the amount of energy stored by biomass from solar energy. The value for DH0 is the amount of heat liberated from the BDCL processing, which can be partially used for electricity generation, and DH + DH0 represents the energy stored in the hydrogen product. The energy conversion efficiency (HHV) from biomass to hydrogen gH2 and electricity gE can be defined by,

Thermal energy in hydrogen ¼ Thermal energy in biomass Net power gE ¼ Thermal energy in biomass gtot ¼ gH2 þ gE

gH2 ¼

8ð43xÞan eH2 32x

DH

¼1þ

DH’ DH

where eH2 denotes the higher heating value (HHV) of hydrogen, and eH2 = 141.9 kJ/kg The BDCL process can produce hydrogen, heat/electricity, or any combination of the aforementioned product by varying x and a. Therefore, the process is product flexible. The reaction stoichiometry limits the key variables at 0 6 x 6 1.5 and 0 6 a 6 1. However, by considering the thermodynamic equilibrium limitations, the suitable ranges of x and a will be further restricted. In order to accurately assess the BDCL system performance and its relationship with x and a, thermodynamic analyses based on the reactor modeling and process simulation are developed in the following sections. 3. Methodology In this section, the key assumptions in this study are first listed. The detailed approaches for reactor modeling and process simulation are then discussed. The commercial process simulator, ASPEN PlusÒ, is selected for this study due to its comprehensive physical property database and process analysis functions. 3.1. Key assumptions Hybrid poplar, a fast growing woody biomass source, is considered as the feedstock for the BDCL process throughout this study. Some of its properties are listed in Table 2 for further reference [27]. Given the high operating temperature and adequate residence time, it is assumed that all the reactions in the BDCL system will reach their equilibriums. All the reactors are operated under isobaric conditions. Iron compounds considered include Fe2O3, Fe3O4, Fe0.947O, Fe and Fe0.877S. The assumptions used in the process simulation are consistent with those used in the reactor modeling. The overall thermal input for the process is set at 100 MW, which is equivalent to feeding wet

Table 3 Material specification. Feedstock to DCL system Dry poplar Air H2O Media 3-Level steam cycle Oxygen carrier Output CO2 H2 Flue gas Ash Sulfur

5% Moisture 79% N2, 21% O2 by volume, 16 atm 32 atm, 240 °C (for hydrogen production) 124 atm(HP)/30 atm(IP)/2 atm(LP)/0.1 atm 66.2% Fe2O3, 33.8% SiC by weight 1 atm >99.99% 60 atm, HHV 141.9 MJ/kg N2, CO2, NOx (<800 ppm), 1 atm Acting as inert Recovered by Claus technology

Table 4 Operating conditions for the process simulation. Environmental conditions Reaction Particle attrition/makeup rate Exhaust temperature from HRSG Heat loss in BDCL system Power biomass pulverization using KDS mill Thermal energy for biomass drying Energy consumption for Claus process Other auxiliaries consumption including Pumps Pressure drop in key reactors All pressure changers Expanders Steam turbines Compressors

T = 25 °C, P = 1 atm All reaction reach equilibrium at high temperatures 0 120 °C 1% of total thermal input 80 kW/ton 0.5% of total thermal input 0.2% of total thermal input 0.5% of total thermal input 2 atm Mechanical efficiency is 1 Isentropic efficiency is 0.9 Isentropic efficiency is 0.86 Four stage with intercooler at 40 °C, isentropic efficiency is 0.83

hybrid poplar (50% moisture) at a rate of 35.73 tonne/h. Such a capacity is suitable for a centralized biomass conversion plant considering the limiting factors such as harvesting and transportation. The material specifications for the BDCL process simulation are summarized in Table 3. A multistage air compressor is used to elevate the pressure of air up to 18 atm. The composition of the oxygen carrier (i.e. the ratio between the fresh Fe2O3 and inert SiC support) is adjusted to ensure that the desired operating conditions are achieved for all three reactors. The final gaseous products

CO2, H2O, N2, SO2

Fe2O3

5 RGIBBS

4 RGIBBS

Dried Poplar

Decomp RYIELD

3 Q

RGIBBS

2 Table 2 Properties of hybrid poplar.

RGIBBS

1

Ultimate analysis (wt.%), dry Ash Carbon Hydrogen Moisture (wt.%), as received

0.92 50.88 6.04 50

Nitrogen Sulfur Oxygen HHV (dry, kcal/kg)

0.17 0.09 41.9 4820

RGIBBS

Promoter (H2O/CO2)

FeO, Fe, Ash

Fig. 4. Model setup for the moving bed reducer.

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from the process, i.e. H2 (and CO2 when carbon sequestration is required), are cooled to near-ambient temperature prior to compression. The H2 stream is compressed to 60 atm for delivery while CO2 is either vented to the atmosphere or compressed and sequestered after sulfur recovery. Table 4 summarizes the key assumptions for the operating conditions in the process simulation. The thermodynamic models described below are used as the basis for the BDCL system modeling. 3.2. Moving bed reactor modeling A moving bed reducer and oxidizer are adopted in the BDCL system, which differs from previous studies where fluidized bed reactors are employed [23,24]. An equilibrium analysis is conducted through the technique of Gibbs free energy minimization. ASPEN PlusÒ offers this function within the RGIBBS block. Coupling this method with the comprehensive physical data bank provided in the software creates a quick and accurate reactor model. One RGIBBS block can be used for estimating a well-mixed fluidized bed reactor. There are no standard functions available in ASPEN PlusÒ for mimicking the gas–solid countercurrent contacting pattern within a moving bed reactor. In order to simulate the moving bed reducer, a multistage equilibrium model, consisting of a set of interconnected RGIBBS blocks, is proposed as configured in Fig. 4. For processing the nonconventional biomass in the RGIBBS block, one RYIELD block is first used to decompose the biomass into its constituent elements based on its ultimate analysis as listed in Table 2. Additionally, a heat stream is used for balancing the thermal energy. Afterwards, the decomposed biomass is injected into the middle stage of the model while iron oxide is introduced from the top stage. Gasifying agents such as CO2 or steam can also be introduced from the bottom stage. Throughout the model, all the gases flow upwards and the solids move downwards. In each stage, the gas and solid reactants mix well and reach the thermodynamic equilibrium. It is corroborated experimentally that such a multi-

stage moving bed reducer model, usually with no less than five stages, can predict the performance of a moving bed reactor with a reasonable accuracy [28]. Through consequential iterations and sensitivity analysis, performance variable x and other related parameters can be calculated and analyzed. A similar multistage reactor model as the reducer configuration can be adopted for the moving bed oxidizer simulation. For the entrained bed combustor, where the reactions are not thermodynamically limited, one RGIBBS block is adequate to simulate its performance. It is noted that even a slight change in the key physical property parameters can considerably alter the results. The current study directly retrieves the thermodynamic data from the database COMBUST, INORGANIC, SOLIDS and PURE. 3.3. Process simulation and configuration Based on the reactor modeling and the thermodynamic analysis, an overall process simulation is conducted using ASPEN PlusÒ to quantify the mass and energy balance for the BDCL process. It is noted that the mass fraction of the inert material in the particles is a design variable that affects the overall system heat balance. Its value can be properly selected so that the operating temperatures that dictate the reactant conversions in the reactors and the heat balance of the BDCL process can be determined. In the simulation, the reducer and combustor are considered to be operated adiabatically. In order to establish a desired reducer temperature, the particle temperature in the combustor and hence the combustor temperature need to be set. As the reaction in the oxidizer is favored at low temperatures, the desired temperature of the oxidizer can be determined also by considering the reactant conversion and the heat requirement for the looping system. The temperature control in the oxidizer can be performed using heat exchangers. The simulation provides the heat and work requirements for the entire process system. The recovered heat from the reactor system is then used to drive gas and steam turbines for power generation. The value a can be decreased if either the gross power output is less

Fig. 5. Process flow diagram for the BDCL process for hydrogen production.

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than the parasitic energy requirement or more electricity generation is desired. When hydrogen is the only desired product in the looping plant, the highest possible value for a is used with the plant maintained under the self-sustaining operating condition. Fig. 5 illustrates a BDCL process configuration for hydrogen and electricity co-production. As shown in Fig. 5, the process consists mainly of biomass preparation, the chemical looping system, heat recovery and steam generation (HRSG), gas cleanup units and power generation systems. The various systems and units are configured in such a manner that the heat and mass flows are effectively integrated to reduce the process energy loss. 3.4. Biomass preparation The biomass preparation involves both mechanical pulverization and drying. As will be indicated in the reactor modeling, the moisture content will not only increase the energy loss through water vaporization, but also decrease the reducer performance. Therefore, drying the biomass as much as possible before being fed into the system is required. In this case, the poplar is first crushed and pulverized into particles with characteristic diameter of around 5 mm using KDS mills, which also reduces the water content to about 20%. For 100 MW thermal BDCL process, the pulverization consumes about 2.86 MW of electricity from the gross power output. Pulverized poplar is then dried to a 5% moisture content using the warm flue gas from the combustor. 3.5. BDCL system Under a typical BDCL configuration, the dried poplar is introduced to the BDCL system from the reducer, which operates at 900–1100 °C, 1 atm. A portion of the reduced oxygen carriers from the reducer is introduced directly into the combustor, which operates at 16 atm and 100–450 °C higher than the reducer. The oxidizer, which converts the remaining portion of the reduced oxygen carriers from the reducer, operates at around 850 °C and 30 atm. 3.6. Heat recovery and steam generation In the BDCL process, most of the reactions occur at high temperatures, resulting in a number of high temperature streams. The HRSG unit recovers the heat from the gaseous product streams producing both high pressure and low pressure steams. The high pressure steam, at 600 °C and 125 atm, is used to drive steam turbines for power generation. The low pressure steam of 240 °C and 32 atm is used in the oxidizer for hydrogen generation. All the gases are cooled to 120 °C before leaving this unit. 3.7. Power consumption and generation

Fig. 6. Effect of moisture on the moving bed reducer operation.

x, is affected by the feedstock, temperature, pressure and gas–solid contacting pattern. The effect of each item on reducer performance is analyzed. The reducer is configured to fully oxidize the biomass into CO2 and H2O while minimizing the circulation rate of the oxygen carrier. 4.2. Effect of moisture At 900 °C and 1 atm, the multistage model is used to compare the reducer performances at various moisture contents of the feedstock. The results are illustrated in Fig. 6. As shown, for non-drying biomass (M = 50%), the maximum oxygen carrier conversion in a moving bed reducer is 26.8% with a minimal molar ratio Rmin = 2.60 between iron oxide and carbon. As the biomass is further dried, both the molar ratio R and the energy requirement E drop while C increases. The results indicate that a lower moisture content in the feedstock correlates to a lower solids circulation rate and energy consumption. Vaporization of excess moisture consumes a large amount of heat. The resulting steam consequently drives the equilibrium of Reaction (2) backwards. Therefore, more iron oxide is needed to fully oxidize the biomass. A similar effect is found when injecting steam or CO2 as the gasifying agent. From the iron oxide conversion standpoint, it is desirable to dry the biomass before entering the reducer and to limit the amount of gasifying agents injected to its lowest effective levels. Based on the system heat balance, the biomass is able to be reasonably dried to a 5% moisture content to limit excess water, and M = 5% is assumed in the following sections.

The major energy consumptions incurred in the BDCL process are from the gas compressors and the biomass pulverizer. Using the hot flue gas from the combustors, an expander is utilized to generate power to meet the parasitic energy requirements. The superheated high pressure steam from the HRSG can also be used for electricity generation via steam turbines.

4.3. Products distribution

4. Results and discussion

Table 5 Yields from the moving bed reducer with M = 5%, at 900 °C, 1 atm.

4.1. Operating condition in the reducer The results of the reducer modeling are discussed in the following section. The optimum value for x is also addressed. As mentioned above, the reducer performance, indicated by the variable

At 900 °C and 1 atm, the multistage model predicts that R = 1.61 is the minimum value to ensure full fuel conversion in the moving bed reducer. The resulting gas and solid streams are summarized in

Gas stream (mol%)

Solid stream (mol%) (ash free)

CO2 H2O N2 SO2

Fe0.947O Fe Fe0.877S

0.562306 0.436501 0.000806 0.000373

0.812341 0.187659 0

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Fig. 7. Effect of temperature on biomass conversion at 1 atm.

Table 5. The model predicts that biomass is fully converted to CO2 and H2O while iron oxide is reduced by 43.4%, which is equivalent to an x of 0.849. The sulfur will exit the reducer with the gaseous stream in the form of SO2. Due to its low concentration in the given fuel source, no sulfur will react with the oxygen carrier particles. Therefore, all the sulfur can be captured from the reducer exhaust gas stream using existing SOx control devices. With all other conditions held constant, decreasing R will lead to a desirable reduction in x. However, when R is lower than 1.61, the unconverted CO and H2 will exit from the reducer resulting in an energy loss. Thus, an Rmin of 1.61 is the optimized value for the given conditions. Further studies also indicate that when Fe3O4 is used in the reducer feed instead of Fe2O3, a considerable amount of fuel is left unconverted due to the thermodynamic equilibrium limitation regardless of the R value. Therefore, it is necessary to regenerate the iron to Fe2O3 from the combustor before entering the reducer.

Fig. 8. Effect of pressure on biomass conversion at 900 °C.

Table 6 Comparisons between the fluidized bed reducer and moving bed reducer. Reactor type

Gas–solid contacting pattern

Rmin

x

C

x

Fluidized bed Moving bed

Mixed Countercurrent

6.28 1.61

0.0333 0.1302

0.111 0.434

1.333 0.849

pressure depresses the desired end product. As shown in Fig. 8 (M = 5%, R = 1.61, 900 °C), low pressures favor biomass conversion, whereas high pressures inhibit carbon gasification. From Fig. 8, an operating pressure between 1 atm and 30 atm is determined as a suitable range for the BDCL reducer operation. However, considering the energy required for CO2 compression, it is desirable to operate the reducer at moderate pressures if CO2 is going to be sequestrated. 4.6. Effect of gas–solid contacting pattern

4.4. Effect of temperature A sensitivity analysis is carried out using the multistage equilibrium model. The carbon distribution in the reaction products with a constant pressure of 1 atm and varied temperature are shown in Fig. 7. From a thermodynamic viewpoint, a higher reaction temperature favors the endothermic biomass–Fe2O3 reaction, as is also predicted by the model. In Fig. 7, using a Fe2O3/biomass ratio identical to the case shown in Table 5 (M = 5%, R = 1.61), the carbon content in the fuel source cannot be fully converted at temperatures below 900 °C unless the R value increases. The product distribution becomes even less favorable below 600 °C caused by the absence of Fe0.947O at these relatively low temperatures. Therefore, the reducer should operate at a higher temperature to maximize both the fuel and oxygen carrier conversions. The operating condition selected must also take into account the material costs for constructing a reactor able to withstand the proposed environment, particle melting point, and the ash softening point. A suitable operating temperature is determined to be approximately 900 °C and is used to estimate the other operating temperatures in the following process simulation. 4.5. Effect of pressure The reactions occurring in the reducer generate gaseous products, adding to the overall system volume. Thus, a higher operating

The flow patterns of the gases and solids play an important role in the reducer operation. In a moving bed reactor, both phases behave in a plug flow manner and interact with each other countercurrently. This design has been studied in the multistage equilibrium model for the BDCL process. In contrast, a well mixed flow pattern occurs in a fluidized bed reactor, which can be studied using a one stage equilibrium model. It should be noted that the actual fluidized bed operation is often worse than the mixed flow. While maintaining the operating conditions, of M = 5%, 1 atm and 900 °C, constant, the two reducer models are compared. The results are illustrated in Table 6. The comparison shows that a fluidized bed reducer demands a 2.9 times greater Fe2O3 flow rate than a moving bed reactor in order to fully convert the same amount of biomass feedstock. As a result, the oxygen carrying capacity and iron oxide conversion in the moving bed reducer are approximately 2.9 times larger than in the alternative design. In the fluidized bed, the value for the performance variable, x, is constrained between 1.333 and 1.5, which is incapable of hydrogen generation due to thermodynamic limitations. On the contrary, a moving bed design enables x to reach values as low as 0.849, suitable for both hydrogen production and electricity generation. Therefore, the moving bed reducer delivers a more promising performance than its fluidized bed counterpart based on the thermodynamic analysis. Similarly, the moving bed oxidizer behaves better than the fluidized bed oxidizer under the same assumptions.

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the same temperature, a lower iron oxidation state (or smaller x) from the reducer outlet results in a higher steam to hydrogen conversion in the oxidizer. The condition of small x is highly desired since a higher steam conversion leads to reduced steam usage for producing unit amount of hydrogen. Hence, the energy loss for the steam generation is decreased. Thus, the operation of the BDCL process requires careful management of the circulating oxygen carrier particles. For the entrained bed combustor, one RGIBBS block is used for this simulation. With these reactor models, thermodynamic limitations for the BDCL system can be further clarified, and the results will be adopted in the later overall process simulation. 4.8. Process performance

Fig. 9. Steam conversion at varying temperatures and x values in the oxidizer at 30 atm.

Thus, by using the multistage model simulation, the operating conditions in the moving bed reducer are analyzed and optimized. The results show that it is suitable to operate the reducer at M = 5%, 1 atm, and 900 °C. This corresponds to a normalized oxygen content x of 0.849.

4.7. Oxidizer performance A similar five stage model is used to analyze the performance in the moving bed oxidizer. Fig. 9 summarizes the simulation results of the oxidizer with an iron oxide feed FeOx from the outlet of the reducer at various temperatures. As noted earlier, x refers to the normalized iron oxidation state from the moving bed reducer. The steam flow rate at the inlet of the oxidizer is taken as the minimum amount that can fully oxidize FeOx to Fe3O4. The modeling shows that the steam to hydrogen conversion decreases with increasing operating temperatures, which is consistent with the exothermic nature of the steam–iron reaction. The performance relation between the reducer and the oxidizer is also illustrated in Fig. 9. As shown, for example at 700 °C, the steam to hydrogen conversion is constant for the values of x between 0 and 0.61, and it starts to decrease when x exceeds 0.61. The steam conversion is near zero when x approaches 1.33, representing the outlet product of the oxidizer, Fe3O4. To summarize, when operated at

The BDCL process can convert the energy in biomass into hydrogen and/or electricity. In this section, the mass and energy flows throughout the process are studied for three cases. The reactor modeling results from the previous sections are integrated into the BDCL process simulation to demonstrate the optimum operation of the BDCL system. 4.9. Hydrogen and electricity coproduction Using the process configuration discussed previously, the BDCL process is capable of cogenerating any combination of electricity and hydrogen while maintaining a self-sustained operation. For the case when a = 0.657, the simulation results are listed in Table 7. The power balance for the overall process is summarized in Table 8. Under this scheme, the oxygen carrier particles comprising 66.2% Fe2O3 and 33.8% inert SiC, are introduced from the top of the reducer at a flow rate of 290.84 tonne/h. The solid oxygen carriers exit the unit in the form of Fe and Fe0.947O (wüstite) while the biomass is fully converted. Any sulfur within the fuel feedstock is removed with the gas stream as SO2. To maintain the overall heat balance for this case, the performance variable, a, is taken as 0.657. This value for a implies that, among the reduced iron particles leaving the reducer, 65.7% is directed to the oxidizer for hydrogen production, while the remaining 34.3% is sent to the combustor directly for heat generation. The steam–iron reaction in the oxidizer is carried out at 850 °C and 30 atm. The steam flow rate is set at 25.52 tonne/h; it yields a steam conversion of 54.7%, as predicted by the multistage oxidizer model. The combustor regenerates the iron particles to Fe2O3 at 1324 °C, 16 atm. Using the chemical looping scheme, the poplar is able to be fully converted into CO2 from the reducer (consisting of mainly 44% CO2 and 56% steam) while high purity H2 is produced from the oxidizer (about 55% H2 and

Table 7 Simulation results in the BDCL system coproducing hydrogen and electricity. Reactor

Feedstock condition

Operating pressure,

Operating temperature

Mole fraction in products (excluding ash)

Conversion

Reducer

Dried poplar 25 °C, 1 atm Steam 240 °C, 32 atm Compressed air 126 °C, 18 atm

1 atm

900 °C

30 atm

850 °C

16 atm

1324 °C

Gas: 0.4365 H2O, 0.5623 CO2, 806 ppm N2, 373 ppm SO2 Solid: 0.097 Fe, 0.410 Fe0.947O, 0.493 SiC Gas: 0.453 H2O, 0.547 H2 Solid: 0.247 Fe3O4, 0.753 SiC Gas: 0.01 O2, 0.989 N2, 509 ppm NO Solid: 0.33 Fe2O3, 0.67 SiC

Biomass: 100% Fe2O3: 43.4% Steam: 54.7% To Fe3O4: 100% O2: 96.2% To Fe2O3: 100%

Oxidizer Combustor

Table 8 Power balance in the BDCL process coproducing hydrogen and electricity. Unit operations

Power, MW

Input

Output

Net power

Air compressor

H2 compressor

Pulverizer

Other

HP steam turbine

IP steam turbine

LP steam turbine

Expander

5.52

0.99

2.86

2.21

1.74

3.09

1.99

10.31

5.55

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45% steam).Based on this conservative simulation, 1.56 tonne/h of compressed hydrogen is produced from the BDCL process equating to a hydrogen HHV efficiency,gH2, of 61.60%. After removing a portion of the energy to satisfy the plant operation requirements, there is still a net electricity amount of 5.55 MW available for the grid. Thus, the BDCL process, with an a of 0.657, can coproduce hydrogen and electricity with a combined efficiency (gH2 + gE) as high as 67.15%.

4.10. Maximized hydrogen production For maximum hydrogen production from the BDCL process, the heat released for the steam generation and power production needs to be minimized. One effective approach is to preheat the air to 800 °C before entering the combustor using the HRSG unit. By doing so, the amount of solids required for direct combustion from the reducer can be reduced to 20.9%. The remainder can then be used for hydrogen production in the oxidizer. This configuration generates 5.71 tonne/h of high pressure steam for steam turbine

Fig. 10. Process efficiency as a linear function of the split ratio a.

utilization, with a net electricity yield decreased to 0.15 MW. 1.88 tonne/h of hydrogen is produced from this process at 60 atm, corresponding to a total HHV hydrogen production efficiency of 74.2%. 4.11. Maximized power generation By sending all the reduced iron oxides directly to the combustor, implying that the oxidizer is eliminated from the process, the power generation is maximized. With this process scheme, the operating temperature within the BDCL system is raised by 200 °C. Also, 72.5% excess air is used to maintain the suitable operating temperature. Another method to maintain temperature can be achieved by recycling the depleted air back to the combustor as a coolant. The overall electricity generation efficiency using this method can reach approximately 38.1%. For power generation, the combustor can be also integrated with supercritical or ultrasupercritical (USC) boiler technology, which eliminates the expander after the combustor. Other schemes could include producing hydrogen first and then generating electricity from hydrogen either through combustion or commercial fuel cells. As deduced from the three cases, a performance curve relative to the variable a is plotted in Fig. 10. At a given biomass feedstock and suitable operating conditions, the performance variable a linearly determines the end product distribution and corresponding process efficiency. In the presented process scheme discussed throughout the article, the range of a is limited between 0 and 0.79 by the heat and work balance. As a decreases, i.e. more reduced iron particles flow to the combustor, more heat will be released for electricity generation with less hydrogen yield. The flexibility between hydrogen production and electricity generation makes the BDCL process versatile in product selection when compared to other biochemical technologies. Compared to the current thermochemical processes given in Table 1, the BDCL process efficiency for biomass conversion is greater in both hydrogen production and electricity generation cases. The increased efficiencies result from an effective energy management system and a unique looping reaction scheme. The high grade heat released from the circulating oxygen carriers significantly reduce the inefficient generation of low pressure steam. Additionally, the BDCL system simplifies the traditional biomass gasification conversion processes by removing the gasifier, air separation unit (ASU) and water gas shift reactor for hydrogen production. Further, elaborate CO2 separation steps are not required. Therefore, through

Fig. 11. Power generation efficiency and operating temperature range as a function of inert content.

F. Li et al. / Fuel 89 (2010) 3773–3784

process intensification, the BDCL process has a lower capital cost and parasitic energy requirements. 4.12. Effect of the inert The inert is used to moderate the operating temperature of the BDCL system. Once the process inputs and outputs are fixed, the overall heat of reaction is determined. By increasing the amount of inert, the temperature differences among reactors can potentially be reduced since the heat capacity of the inert material can help decrease the required temperature gradients. In the present case, all the reduced oxygen carriers are directly combusted (a = 0) under the BDCL-USC configuration. Under this configuration, the combustor is operated at a temperature that ensures enough heat to be carried over to the reducer for the subsequent biomass–oxygen carrier reaction. The excess heat from the BDCL system is recovered by an integrated USC boiler. The steam specification of the USC boiler is 650 °C and 300 atm. Simulation results show that the adiabatic combustor temperatures and process efficiency vary notably with the mass percentage of inert SiC. As shown in Fig. 11, the operating temperature range decreases with increasing inert content whereas the process efficiency shows an opposite trend. The operating temperature can be further lowered by introducing excess air into the entrained bed combustor. Based on the modeling results, 30–70% inert content is a suitable range for the BDCL process. Additionally, the inert compositions can also serve as structural support to minimize attrition and promote reaction kinetics. 4.13. Pollutants control in the BDCL process By capturing CO2 from the reducer, the BDCL process becomes carbon negative. Compressing CO2 to 150 atm for sequestration requires 3% of the total thermal energy input. The ash in biomass can be separated from the oxygen carrier based on size differences using a cyclone. Biomass may also contain sulfur. Through reactor modeling and process simulation, the fate of sulfur can be determined. Poplar has a sulfur content of less than 0.21 wt.% (dry basis). With such low sulfur content, all the sulfur in the biomass will be converted to SO2 in the reducer. For biomass sources with higher sulfur contents, Fe0.877S may form in the reducer. The presence of Fe0.877S will lead to H2S contamination to the hydrogen stream from the oxidizer. Using conventional sulfur control devices, SO2 and H2S can be captured and converted to elemental sulfur. NOx can also be formed in the combustor. The flameless combustion environment and the absence of fuel nitrogen in the combustor may lead to NOx level significantly lower than that in a conventional coal fired boiler. 5. Concluding remarks The biomass direct chemical looping (BDCL) process is simulated, analyzed, and discussed in this study. The process consists of three chemical looping reactors, i.e. the reducer, oxidizer, and combustor, with an iron based particle circulating among the reactors. The iron based particle is used as the carrier for oxygen, chemical energy, and thermal energy, facilitating efficient conversion of biomass to hydrogen and/or electricity. The four oxidation states of iron and their distinct thermodynamic properties give rise to a large degree of freedom in the process design and operating parameters. By using a thermodynamic analysis tool, i.e. ASPEN PlusÒ, the suitable reactor design, operating conditions, and process configuration for the BDCL process are determined. A multistage model is developed in this study to simulate the performance of individual BDCL reactor based on ASPEN PlusÒ.

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The simulation results indicate that a countercurrent gas–solid flow pattern such as that provided by a countercurrent moving bed reactor maximizes the conversions of both solids and gas. They also indicate that a higher moisture content in the biomass feedstock leads to a lower oxygen carrier conversion in the reducer, lower steam to hydrogen conversion in the oxidizer, higher solids circulation rate, and lower process energy conversion efficiency. Thus, maintaining the moisture content in the biomass feedstock at 5 wt.% or lower is desirable prior to the chemical looping reactions. The conversions of the biomass and iron based particles in the reducer are also favored at higher temperatures and lower pressures. Considering various practical factors such as reactor materials and gas compression, the suitable operating temperature for the reducer is 900 °C with pressures between 1 and 30 atm. The steam to hydrogen conversion in the oxidizer is more favored at lower temperatures. The overall energy balance of the process, however, requires the oxidizer to be operated at above 600 °C. To compensate for the heat required in the reducer, the combustor is operated at temperatures 100–450 °C above the reducer. The process simulation based on the reactor modeling results further reflects that the BDCL process can produce hydrogen and/ or electricity at any ratio. Compared to conventional biomass combustion and gasification processes, the BDCL process is 10–25% more efficient. The sulfur and NOx pollutants from the process can be readily removed using commercial pollutant control devices. Additionally, the CO2 stream generated from the BDCL process is of a high concentration. With the CO2 sequestration, the BDCL becomes a carbon negative process from the life cycle standpoint. The unique energy management and conversion scheme of the BDCL process, thus, render it potentially viable for clean and efficient biomass conversion applications. Acknowledgement The authors would like to acknowledge Dr. Nobusuke Kobayashi, Mr. Andrew Tong, and Mr. Eric Sacia for their helpful comments on the manuscript. References [1] Gera D. Biofuels and bioenergy. Encylopedia of chemical processing. CRC Press; 2005. [2] Perlack RD, Wright LL, Turhollow AF, Graham RL, Stokes BJ, Erbach DC. Biomass as feedstock for a bioenergy and bioproducts industry: the technical feasability of a billion-ton annual supply. US DOE and USDA; 2005. p. 16–8. [3] Sofer SS, Zaborsky O. Biomass conversion processes for energy and fuels. Plenum Press; 1981. [4] Gerbens-Leenes W, Hoekstra AY, van der Meer TH. The water footprint of bioenergy. Proc Nat Acad Sci 2009;106(25):10219–23. [5] Turner J, Sverdrup G, Mann MK, Maness P, Kroposki B, Ghirardi M, et al. Renewable hydrogen production. Int J Energy Res 2008;32(5):379–407. [6] Lewis CW. Efficiency improvements in bioenergy conversion systems. Energ Convers 1979;19(3):125–31. [7] Fan L-S, Li F, Clean coal. Phys World 2007;20(7):37–41. [8] Li F, Fan L-S. Clean coal conversion processes – progress and challenges. Energy Environ Sci 2008;1(2):248–67. [9] Fan L-S, Li F, Ramkumar S. Utilization of chemical looping strategy in coal gasification processes. Particuology 2008;6(3):131–42. [10] Li F, Kim HR, Sridhar D, Wang F, Zeng L, Chen J, et al. Syngas chemical looping gasification process: oxygen carrier particle selection and performance. Energ Fuels 2009;23(8):4182–9. [11] Gupta P, Velazquez-Vargas LG, Fan L-S. Syngas redox (SGR) process to produce hydrogen from coal derived syngas. Energ Fuels 2007;21:2900–8. [12] Thomas T, Fan L-S, Gupta P, Velazquez-Vargas LG. Combustionlooping using composite oxygen carriers, US patent appl publ 2005; 2005-175533-A1. [13] Cho P, Mattisson T, Lyngfelt A. Comparison of iron-, nickel-, copper- and manganese-based oxygen carriers for chemical-looping combustion. Fuel 2004;83(9):1215. [14] Adanez J, Garcia-Labiano F, de Diego LF, Gayan P, Abad A, Celaya J. Development of oxygen carriers for chemical-looping combustion. Carbon dioxide capture storage deep geologic formations-results CO2 capture Project 2005;1:587. [15] Adanez J, Garcia-Labiano F, de Diego LF, Plata A, Celaya J, Gayan P, et al. Optimizing the fuel reactor for chemical looping combustion. Proc Int Conf Fluidized Bed Combust 2003;17:173.

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