Fuel 89 (2010) 3773–3784

Contents lists available at ScienceDirect

Fuel

journal homepage: www.elsevier .com/locate / fuel

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 2009Received in revised form 22 June 2010Accepted 14 July 2010Available online 30 July 2010

Keywords:Chemical loopingBiomassMoving bed reactor modelingProcess simulation

0016-2361/$ - see front matter � 2010 Elsevier Ltd. Adoi:10.1016/j.fuel.2010.07.018

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

a b s t r a c t

Biomass is a clean and renewable energy source. The efficiency for biomass conversion using conven-tional fuel conversion techniques, however, is constrained by the relatively low energy density and highmoisture content of biomass. This study presents the biomass direct chemical looping (BDCL) process, analternative process, which has the potential to thermochemically convert biomass to hydrogen and/orelectricity with high efficiency. Process simulation and analysis are conducted to illustrate the individualreactor performance and the overall mass and energy management scheme of the BDCL process. A mul-tistage model is developed based on ASPEN Plus� to account for the performance of the moving bed reac-tors considering the reaction equilibriums. The optimum operating conditions for the reactors are alsodetermined. Process simulation utilizing ASPEN Plus� is then performed based on the reactor perfor-mance data obtained from the multistage model. The simulation results indicate that the BDCL processis significantly more efficient than conventional biomass conversion processes. Moreover, concentratedCO2, produced from the BDCL process is readily sequesterable, making the process carbon negative. Sev-eral BDCL configurations are investigated for process optimization purposes. The fates of contaminantsare also examined.

� 2010 Elsevier Ltd. All rights reserved.

1. Introduction

Biomass is defined as ‘‘any organic matter, which is available ona renewable basis, including agricultural crops and agriculturalwastes and residues, wood and wood wastes and residues, animalwastes, municipal wastes, and aquatic plants” [1]. This renewableresource can be converted into electricity, H2, chemicals, and liquidfuels. Compared to fossil fuel, biomass conversion is less carbonand pollutant intensive. Moreover, it is widely distributed andabundantly available. For instance, the annual biomass potentialfrom forest and agricultural resources alone is over 1.3 billion drytonne in the United States, which is sufficient to displace nearly40% of the petroleum consumed in the nation [2]. Therefore, bio-mass has the potential to be a favorable energy source to, at leastpartially, address the increasing environmental and energy sus-tainability concerns.

Most of the carbon in the atmosphere circulates in the form ofCO2. 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, aportion of the solar energy is chemically stored in these organisms.The stored energy can be then transferred into various formsthrough the food chain consumption scheme and anthropogenicactivities.

ll rights reserved.

: +1 614 292 3769.

6nCO2 þ 5nH2O !enzymesðC6H10O5Þn þ 6nO2 þ DH ð1Þ

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 relativelylow energy density and wide geographical distribution requireseffective strategies to ensure the techno-economical viability forbiomass conversion in the modern era. Biomass conversion tech-nology can be classified into either biochemical or thermochemicalprocesses [3]. Biochemical processes involve the use of microor-ganisms to convert biomass into ethanol or other valuable prod-ucts. However, such processes consume a large amount of water[4]. Consequently, energy intensive distillation units are requiredfor product upgrading. Moreover, biochemical techniques havestringent requirements on the properties of the feedstock. Starchis relatively easy to convert; however, starch based fermentationprocesses compete against human food supplies, inflating the costof agricultural products. Cellulosic biomass, although is not edible,is rather difficult to convert through a biochemical approach.

Compared to biochemical processes, thermochemical processeshave less stringent requirements on the feedstock properties andconsume less water. Several thermochemical systems for biomassconversion are currently under development. They include com-bustion, gasification, pyrolysis, and chemical looping. Conventionalcombustion processes fully oxidize the biomass with air to produceheat and/or electricity. Gasification processes partially oxidize thebiomass with pure oxygen or air at high temperatures to produce

Notation

x normalized oxygen content of the reducer product, i.e.the number of transferable oxygen atoms/the numberof Fe atom in reduced oxygen carrier from the reducer

a split ratio, the fraction of the particles that are directlyintroduced to the combustor from the reducer

xFe2O3Transferable oxygen weight

Oxygen carrier weight ¼ 16 3=2�xð Þ56þ16�3=2 ¼ 3�2x

12ðwithout considering inertÞ

M mass percentage of moisture in dried biomass (%)

R the molar ratio between fresh Fe2O3 in oxygen carrierand carbon in biomass introduced to the reducer

C oxygen carrier conversion, 1 � 2x/3E ratio between the reducer heat requirement when the

moisture content in the biomass is M and the heatrequirement when 5 wt% moisture biomass is the feed

3774 F. Li et al. / Fuel 89 (2010) 3773–3784

synthesis gas (syngas). This gaseous fuel can then be processed intoa variety of products including electricity, hydrogen, liquid fuels,and chemicals. Both combustion and gasification processes areexothermic and are relatively inefficient due to the inevitableexergy loss in the combustion or gasification step [5]. Biomasscan also be converted into gases, bio oils, tars and char throughpyrolysis at elevated temperatures in an oxygen free environment.The endothermic nature of the pyrolysis reactions requires effec-tive heat integration strategies, which are yet to be developed.The energy conversion efficiencies of several thermochemical pro-cesses are listed in Table 1 [6]. As can be seen, the conventionaltechnologies are relatively inefficient due to both the technicallimitations and low heating value of the biomass feedstock. Therelatively low process efficiencies for biomass conversion, com-pounded with the high biomass collection and transportationcosts, significantly affect the economic feasibility of these energyconversion systems. Therefore, efficient biomass conversion strat-egies, preferably those applicable at small scales, are desirable.Direct chemical looping provides a promising method to effectivelyutilize the potentials of biomass [7,8].

The chemical looping technology [9], illustrated in Fig. 1, con-verts carbonaceous fuels using metal-oxide based oxygen carrierscirculating between two main reactors, a reducer and an oxidizer.Within the redox loop, oxygen transfers from the metal oxides tofuel in the reducer and from steam or air to reduced metal oxidesin the oxidizer. This strategy avoids direct contact between fuel

Table 1Energy efficiencies of selected thermochemical processes.

Process Products Efficiency (%)

Direct combustion Electricity 25Pyrolysis-flash Pyrolytic oil, gas, char 37–45Oxygen-blown gasification Fuel gas 50–60

Reducer Oxidizer

H2O/Air

H2/N2+Power

Fuel

CO2, H2O…

M

MO

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

and oxidants. Hence, a separate CO2 product stream is generatedwith minimal energy consumption. Chemical looping processesusing gaseous feedstock have been extensively studied in the pastdecade. Experimental studies from the lab to small pilot(�120 kW) scales are being performed to demonstrate the feasibil-ity of the chemical looping concept. The ongoing studies focus onoxygen carrier particle performance, reaction chemistry, solidshandling, and process system design and testing [10–16]. Comple-mentary to experimental study, theoretical investigations, such asthermodynamic analysis and process simulation, have been per-formed to evaluate the feasibility of the chemical looping strategyfrom the overall process standpoint. The thermodynamic investiga-tions on chemical looping combustion (CLC) have been performedby Richter [17] and McGlashan [18]. Both studies confirm the po-tential of this processing strategy in increasing the efficienciesfor power generation. For its application in hydrogen production,Xiang [19] and Cleeton [20] suggested several process configura-tions using syngas derived from coal gasification. Other types ofgaseous feedstock such as methane [21] are also proposed as fuelsupplies for the chemical looping process to produce hydrogen.

Compared to gaseous fuels, the direct conversion of solid fuelssuch as biomass using chemical looping strategy is a relativelynew subject. Pan et al. [22] studied the reaction between copperbased oxygen carrier particles and woody biomass using thermo-gravimetric analyzers and concluded that the oxygen carrierenhances the conversion of biomass. However, the application ofcopper based oxygen carrier is limited by its low melting pointand the inability to react with steam for hydrogen generation. Shenet al. [23] and Leion et al. [24] tested biomass conversion using ironbased oxygen carrier particles in circulating fluidized bed reactorsystems. Both the fuel conversion and CO2 concentration in thereducer product gas, however, are significantly lower than thosein gaseous fuels CLC systems. In addition, both systems are limitedto electricity generations. Fan et al. proposed a novel direct chem-ical looping process [8,9,25] that directly converts solid fuels suchas coal and biomass into hydrogen and/or electricity using ironbased oxygen carrier particles and moving bed reactor system.Gnanapragasam et al. performed a process analysis on the directchemical looping system, showing that the process is efficient forhydrogen production [26]. The limitations in gas and solid conver-sions due to thermodynamic equilibrium, however, are not takeninto account in this analysis.

This study conducts a comprehensive process simulation usingbiomass as the feedstock in order to ascertain its DCL processapplication. The paper first provides an overview on the biomassdirect chemical looping (BDCL) process and its mass and energyflow schemes. A multistage model is then developed to study thethermodynamic equilibrium in the BDCL reactors. The multistagemodel assists in predicting the moving bed reactor performanceand in optimizing the operating conditions. It is followed by overallprocess simulation using the results obtained from the multistagereactor modeling. Finally, the feasibility of the BDCL process isevaluated.

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

F. Li et al. / Fuel 89 (2010) 3773–3784 3775

2. BDCL process overview

As illustrated in Fig. 2, the BDCL process converts biomass, gen-eralized as (C6H10O5)n, to electricity and hydrogen using threereactors, i.e. the Reducer, Oxidizer, and Combustor. Each reactorperforms a unique function in the chemical looping scheme. Thethree reactors closely interact with one another via the circulationof iron oxide based oxygen carrier particles, which is an essentialcarrier for both chemical and thermal energy.

2.1. Reducer

The reducer is a moving bed reactor where the oxygen carrierparticles, in Fe2O3 form, react with biomass feed to produce re-duced Fe/FeO solids and a gaseous CO2/H2O stream. Here, theFe2O3 particles are fed from the top while the dried biomass is in-jected into the middle section of the reactor. Additionally, a smallamount of gasifying agents such as steam, CO2, and/or H2 can be

Fig. 3. Gas solid contacting

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:

ðC6H10O5Þn þ24n

3� 2xFeO1:5 !gasifying agents

6nCO2 þ 5nH2O

þ 24n3� 2x

FeOx þ DH1 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 providedin Fig. 3. Initially, thermal decomposition takes place as the bio-mass is exposed to the high temperature environment of the redu-cer generating char, tar, and pyrolysis gases such as CO2, H2O, COand H2. The gasifying agents facilitate interactions between char/tar and the oxygen carrier particles. Specifically, the oxidativeproducts, CO2 and H2O, instigate the char/tar gasification intoreductive gases, such as CO and H2. The gaseous fuels are then oxi-dized by the iron oxides. As the reducer is a moving bed, the lower

pattern of the reducer.

3776 F. Li et al. / Fuel 89 (2010) 3773–3784

section of the reactor is in a reducing environment due to the pres-ence of excess char. To compare, the upper portion where the freshFe2O3 first enters the reducer is under a more oxidizing environ-ment. The gasifying agents, or promoter gases, are injected fromthe bottom of the lower section to gasify the remaining char, pro-ducing gaseous fuels that travel upward through the vessel to reactwith iron oxides (mainly FeO and Fe3O4). The reactions betweenthe fuel and the oxygen carriers generate reduced Fe/FeO particlesand oxidative gases that further ascend to assist in gasifying thesolid fuel above. Prior to exiting this reactor, the gas mixture reactswith Fe2O3/Fe3O4 solids producing a concentrated CO2 and H2Ooutlet stream. As can be seen in Fig. 3, the upper portion of the re-ducer uses the higher oxidation state iron oxides to fully convertthe gaseous phase into CO2 and H2O, while the lower portion (par-tially) oxidizes or gasifies the biomass char with lower oxidationstate iron oxides.

Full biomass conversion is desired to improve the energy con-version and prevent any unconverted solid fuels from contaminat-ing the ensuing reactors. In the BDCL process, all the carbon in thebiomass feedstock leaves the reducer in the form of CO2 as illus-trated in Reaction (2). After heat recovery and gas cleanup, the gas-eous product can be directly vented into the atmosphere with closeto zero net CO2 emission. Alternatively, by condensing the steamfrom the outlet gases, the concentrated CO2 stream produced canbe compressed and sequestered resulting in a carbon negative pro-cess from the life cycle standpoint. Another issue to be consideredis the highly endothermic nature of the reducer. To compensate forthe heat required, the oxygen carrier particles are heated to a high-er temperature in the combustor before entering the reducer. Theoverall heat integration scheme for the BDCL process is given inSection 2.4.

2.2. Oxidizer

A portion (a) of the reduced oxygen carrier particles from thereducer outlet are transferred to the oxidizer for hydrogen produc-tion. The remaining reduced oxygen carrier particles are directlycombusted with air in the combustor for heat generation. The oxi-dizer is a countercurrent moving bed reactor that produces hydro-gen via the steam–iron reaction at 700–900 �C and 30 atm. Asshown in Reaction (3), the amount of hydrogen produced can beadjusted by varying a. In this reactor, FeOx is oxidized to Fe3O4

as it moves downwards. Meanwhile, H2O is reduced to H2 as ittravels upwards. By condensing steam, a high purity hydrogenstream can be produced. The steam–iron reaction is mildlyexothermic.

24an3� 2x

FeOx þ8ð4� 3xÞan

3� 2xH2O! 24an

3� 2xFeO1:33

þ 8ð4� 3xÞan3� 2x

H2 þ DH2 ð3Þ

where a denotes the percentage of FeOx used for hydrogen produc-tion, 0 6 a 6 1.

2.3. Combustor

The combustor is an entrained flow reactor through which theoxygen carrier particles, from either the reducer or the oxidizer,are conveyed back to the reducer using compressed air. During thisstep, the chemical intermediate is re-oxidized to Fe2O3 via Reac-tions (4) and (5) releasing a significant amount of heat. At adiabaticconditions, the reactor can reach above 1100 �C. The temperaturewithin can be moderated by adjusting the amount of excess airor the support content in the oxygen carrier particle. In most cases,the heat carried by the solids from the combustor can fully com-pensate for the thermal energy required in the reducer. The flue

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

24ð1� aÞn3� 2x

FeOx þ 6ð1� aÞnO2 !24ð1� aÞn

3� 2xFeO1:5 þ DH3 ð4Þ

24an3� 2x

FeO1:33 þ24an

3� 2xO2 !

24an3� 2x

FeO1:5 þ DH4 ð5Þ

2.4. Material and energy management in the BDCL system

In the reducer, Fe2O3 is reduced to FeOx. The normalized oxygencontent, x, is inversely related to the effective oxygen carryingcapacity of the particle, x. Larger values of x, or lower values ofx, correspond to higher solids circulation rates (24n/(3 � 2x)). Afluidized bed reducer design can only achieve full fuel conversioninto CO2 and H2O when x is between 4/3 and 3/2. The results trans-late to a x value of less than 0.03. Because Fe3O4 and Fe2O3 arethermodynamically incapable of converting H2O into H2, x shouldbe less than 4/3 if hydrogen is the desired product. A moving bedreducer design is an effective alternative to achieve a higher solidsconversion. The most desirable case is to obtain metallic iron(x = 0) from the reducer minimizing the overall iron oxide circula-tion rate. Further details on the advantages of moving bed reactorsin 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 xand high a value are preferred. Quantitatively, a(8 � 6x)/(9 � 6x)of the transferred oxygen is used for hydrogen production. Itshould be noted that the present analysis for the oxidizer is basedon reaction stoichiometry, which does not consider the equilib-rium limitations. From the thermodynamic equilibrium stand-point, the value of x from the reducer affects both the steam tohydrogen conversion and oxygen utilization. Additionally, a lowoperating temperature thermodynamically favors steam to hydro-gen conversion in the oxidizer. The validity of these analyses is dis-cussed in Section 4.2.

As stated in the previous section, the temperature of the oxygencarrier particles leaving the combustor should be high enough tocompensate for the heat requirement in the reducer. A lower oxi-dizer operating temperature, though thermodynamically favorablefor the reactor, may not be desirable for the overall process since alarger amount of heat needs to be generated in the combustor tooffset the temperature gradient between the oxidizer outlet/com-bustor inlet and combustor outlet/reducer inlet. Further analysisis conducted to determine the optimum operating conditions forthe oxidizer considering all the aforementioned constraints.

Reaction (6) given below, obtained by summation of Reactions(2)–(5), represents the overall reaction of the process.

ðC6H10O5Þnþð10xþ32a�15�24axÞn

3�2xH2O

þð18xþ12ax�12x�16aÞn3�2x

O2 !FeOx 6nCO2þ

8ð4�3xÞan3�2x

H2þDH0

ð6Þ

DH0 ¼X4

i¼1

DHi

The first law of thermodynamics dictates that DH0for a self-sus-taining BDCL system (i.e. no external heat input required) needs tobe smaller than or equal to zero. Fig. 2 schematically summarizesthe material and energy flow in the BDCL process. The high tem-perature oxygen carriers can effectively transfer both oxygen andheat. Furthermore, the heat released from the chemical looping

Table 3Material specification.

Feedstock to DCL systemDry poplar 5% MoistureAir 79% N2, 21% O2 by volume, 16 atmH2O 32 atm, 240 �C (for hydrogen production)Media3-Level steam cycle 124 atm(HP)/30 atm(IP)/2 atm(LP)/0.1 atmOxygen carrier 66.2% Fe2O3, 33.8% SiC by weightOutputCO2 1 atmH2 >99.99% 60 atm, HHV 141.9 MJ/kgFlue gas N2, CO2, NOx (<800 ppm), 1 atmAsh Acting as inertSulfur Recovered by Claus technology

Table 4Operating conditions for the process simulation.

Environmental conditions T = 25 �C, P = 1 atmReaction All reaction reach equilibrium at high

temperaturesParticle attrition/makeup rate 0Exhaust temperature from HRSG 120 �CHeat loss in BDCL system 1% of total thermal inputPower biomass pulverization using

KDS mill80 kW/ton

Thermal energy for biomass drying 0.5% of total thermal inputEnergy consumption for Claus

process0.2% of total thermal input

Other auxiliaries consumptionincluding Pumps

0.5% of total thermal input

Pressure drop in key reactors 2 atmAll pressure changers Mechanical efficiency is 1Expanders Isentropic efficiency is 0.9Steam turbines Isentropic efficiency is 0.86Compressors Four stage with intercooler at 40 �C,

isentropic efficiency is 0.83

F. Li et al. / Fuel 89 (2010) 3773–3784 3777

system can be utilized for supplementary heating and electricitygeneration to compensate for parasitic energy consumptions. Whencombining Reaction (6) with the atmospheric CO2 circulation inReaction (1), Reaction (7) becomes the same as water photolysis,

8ð4� 3xÞan3� 2x

H2O!CO2 8ð4� 3xÞan3� 2x

H2 þ4ð4� 3xÞan

3� 2xO2 þ DHþ DH0

ð7Þ

From this reaction scheme, the overall thermal input, DH, is theamount of energy stored by biomass from solar energy. The valuefor DH0 is the amount of heat liberated from the BDCL processing,which can be partially used for electricity generation, andDH + DH0 represents the energy stored in the hydrogen product.The energy conversion efficiency (HHV) from biomass to hydrogengH2 and electricity gE can be defined by,

gH2¼ Thermal energy in hydrogen

Thermal energy in biomass¼

8ð4�3xÞan3�2x eH2

DH¼ 1þ DH’

DH

gE ¼Net power

Thermal energy in biomassgtot ¼ gH2

þ gE

where eH2 denotes the higher heating value (HHV) of hydrogen, andeH2 = 141.9 kJ/kg

The BDCL process can produce hydrogen, heat/electricity, or anycombination of the aforementioned product by varying x and a.Therefore, the process is product flexible. The reaction stoichiome-try limits the key variables at 0 6 x 6 1.5 and 0 6 a 6 1. However,by considering the thermodynamic equilibrium limitations, thesuitable ranges of x and a will be further restricted. In order toaccurately assess the BDCL system performance and its relation-ship with x and a, thermodynamic analyses based on the reactormodeling and process simulation are developed in the followingsections.

5RGIBBS

4RGIBBS

3RGIBBS

Dried Poplar DecompRYIELD

Fe2O3CO2, H2O, N2, SO2

Q

3. Methodology

In this section, the key assumptions in this study are first listed.The detailed approaches for reactor modeling and process simula-tion are then discussed. The commercial process simulator, ASPENPlus�, is selected for this study due to its comprehensive physicalproperty database and process analysis functions.

3.1. Key assumptions

Hybrid poplar, a fast growing woody biomass source, is consid-ered 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 residencetime, it is assumed that all the reactions in the BDCL system willreach their equilibriums. All the reactors are operated under iso-baric conditions. Iron compounds considered include Fe2O3,Fe3O4, Fe0.947O, Fe and Fe0.877S.

The assumptions used in the process simulation are consistentwith those used in the reactor modeling. The overall thermal inputfor the process is set at 100 MW, which is equivalent to feeding wet

Table 2Properties of hybrid poplar.

Ultimate analysis (wt.%), dry

Ash 0.92 Nitrogen 0.17Carbon 50.88 Sulfur 0.09Hydrogen 6.04 Oxygen 41.9Moisture (wt.%), as received 50 HHV (dry, kcal/kg) 4820

hybrid poplar (50% moisture) at a rate of 35.73 tonne/h. Such acapacity is suitable for a centralized biomass conversion plant con-sidering the limiting factors such as harvesting and transportation.The material specifications for the BDCL process simulation aresummarized in Table 3. A multistage air compressor is used to ele-vate the pressure of air up to 18 atm. The composition of the oxy-gen carrier (i.e. the ratio between the fresh Fe2O3 and inert SiCsupport) is adjusted to ensure that the desired operating condi-tions are achieved for all three reactors. The final gaseous products

2RGIBBS

1RGIBBS

Promoter (H2O/CO2) FeO, Fe, Ash

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

3778 F. Li et al. / Fuel 89 (2010) 3773–3784

from the process, i.e. H2 (and CO2 when carbon sequestration is re-quired), are cooled to near-ambient temperature prior to compres-sion. The H2 stream is compressed to 60 atm for delivery while CO2

is either vented to the atmosphere or compressed and sequesteredafter sulfur recovery. Table 4 summarizes the key assumptions forthe operating conditions in the process simulation. The thermody-namic models described below are used as the basis for the BDCLsystem modeling.

3.2. Moving bed reactor modeling

A moving bed reducer and oxidizer are adopted in the BDCL sys-tem, which differs from previous studies where fluidized bed reac-tors are employed [23,24]. An equilibrium analysis is conductedthrough the technique of Gibbs free energy minimization. ASPENPlus� offers this function within the RGIBBS block. Coupling thismethod with the comprehensive physical data bank provided inthe software creates a quick and accurate reactor model. OneRGIBBS block can be used for estimating a well-mixed fluidizedbed reactor.

There are no standard functions available in ASPEN Plus� formimicking the gas–solid countercurrent contacting pattern withina moving bed reactor. In order to simulate the moving bed reducer,a multistage equilibrium model, consisting of a set of intercon-nected RGIBBS blocks, is proposed as configured in Fig. 4. For pro-cessing the nonconventional biomass in the RGIBBS block, oneRYIELD block is first used to decompose the biomass into its con-stituent elements based on its ultimate analysis as listed in Table2. Additionally, a heat stream is used for balancing the thermal en-ergy. Afterwards, the decomposed biomass is injected into themiddle stage of the model while iron oxide is introduced fromthe top stage. Gasifying agents such as CO2 or steam can also beintroduced from the bottom stage. Throughout the model, all thegases flow upwards and the solids move downwards. In each stage,the gas and solid reactants mix well and reach the thermodynamicequilibrium. It is corroborated experimentally that such a multi-

Fig. 5. Process flow diagram for the BDC

stage moving bed reducer model, usually with no less than fivestages, can predict the performance of a moving bed reactor witha reasonable accuracy [28]. Through consequential iterations andsensitivity analysis, performance variable x and other relatedparameters can be calculated and analyzed.

A similar multistage reactor model as the reducer configurationcan be adopted for the moving bed oxidizer simulation. For the en-trained bed combustor, where the reactions are not thermodynam-ically limited, one RGIBBS block is adequate to simulate itsperformance. It is noted that even a slight change in the key phys-ical property parameters can considerably alter the results. Thecurrent study directly retrieves the thermodynamic data fromthe database COMBUST, INORGANIC, SOLIDS and PURE.

3.3. Process simulation and configuration

Based on the reactor modeling and the thermodynamic analy-sis, an overall process simulation is conducted using ASPEN Plus�

to quantify the mass and energy balance for the BDCL process. Itis noted that the mass fraction of the inert material in the particlesis a design variable that affects the overall system heat balance. Itsvalue can be properly selected so that the operating temperaturesthat dictate the reactant conversions in the reactors and the heatbalance of the BDCL process can be determined. In the simulation,the reducer and combustor are considered to be operated adiabat-ically. In order to establish a desired reducer temperature, theparticle temperature in the combustor and hence the combustortemperature need to be set. As the reaction in the oxidizer is fa-vored at low temperatures, the desired temperature of the oxidizercan be determined also by considering the reactant conversion andthe heat requirement for the looping system. The temperature con-trol in the oxidizer can be performed using heat exchangers. Thesimulation provides the heat and work requirements for the entireprocess system. The recovered heat from the reactor system is thenused to drive gas and steam turbines for power generation. Thevalue a can be decreased if either the gross power output is less

L process for hydrogen production.

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

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

Gas stream (mol%) Solid stream (mol%) (ash free)

CO2 0.562306 Fe0.947O 0.812341H2O 0.436501 Fe 0.187659N2 0.000806 Fe0.877S 0SO2 0.000373

F. Li et al. / Fuel 89 (2010) 3773–3784 3779

than the parasitic energy requirement or more electricity genera-tion is desired. When hydrogen is the only desired product in thelooping plant, the highest possible value for a is used with theplant maintained under the self-sustaining operating condition.

Fig. 5 illustrates a BDCL process configuration for hydrogen andelectricity co-production. As shown in Fig. 5, the process consistsmainly of biomass preparation, the chemical looping system, heatrecovery and steam generation (HRSG), gas cleanup units andpower generation systems. The various systems and units are con-figured in such a manner that the heat and mass flows are effec-tively integrated to reduce the process energy loss.

3.4. Biomass preparation

The biomass preparation involves both mechanical pulveriza-tion and drying. As will be indicated in the reactor modeling, themoisture content will not only increase the energy loss throughwater vaporization, but also decrease the reducer performance.Therefore, drying the biomass as much as possible before beingfed into the system is required. In this case, the poplar is firstcrushed and pulverized into particles with characteristic diameterof around 5 mm using KDS mills, which also reduces the watercontent to about 20%. For 100 MW thermal BDCL process, the pul-verization consumes about 2.86 MW of electricity from the grosspower output. Pulverized poplar is then dried to a 5% moisturecontent using the warm flue gas from the combustor.

3.5. BDCL system

Under a typical BDCL configuration, the dried poplar is intro-duced to the BDCL system from the reducer, which operates at900–1100 �C, 1 atm. A portion of the reduced oxygen carriers fromthe reducer is introduced directly into the combustor, which oper-ates at 16 atm and 100–450 �C higher than the reducer. The oxi-dizer, which converts the remaining portion of the reducedoxygen carriers from the reducer, operates at around 850 �C and30 atm.

3.6. Heat recovery and steam generation

In the BDCL process, most of the reactions occur at high temper-atures, resulting in a number of high temperature streams. TheHRSG unit recovers the heat from the gaseous product streamsproducing both high pressure and low pressure steams. The highpressure steam, at 600 �C and 125 atm, is used to drive steam tur-bines for power generation. The low pressure steam of 240 �C and32 atm is used in the oxidizer for hydrogen generation. All thegases are cooled to 120 �C before leaving this unit.

3.7. Power consumption and generation

The major energy consumptions incurred in the BDCL processare from the gas compressors and the biomass pulverizer. Usingthe hot flue gas from the combustors, an expander is utilized togenerate power to meet the parasitic energy requirements. Thesuperheated high pressure steam from the HRSG can also be usedfor electricity generation via steam turbines.

4. Results and discussion

4.1. Operating condition in the reducer

The results of the reducer modeling are discussed in the follow-ing section. The optimum value for x is also addressed. As men-tioned above, the reducer performance, indicated by the variable

x, is affected by the feedstock, temperature, pressure and gas–solidcontacting pattern. The effect of each item on reducer performanceis analyzed. The reducer is configured to fully oxidize the biomassinto CO2 and H2O while minimizing the circulation rate of the oxy-gen carrier.

4.2. Effect of moisture

At 900 �C and 1 atm, the multistage model is used to comparethe reducer performances at various moisture contents of the feed-stock. The results are illustrated in Fig. 6. As shown, for non-dryingbiomass (M = 50%), the maximum oxygen carrier conversion in amoving bed reducer is 26.8% with a minimal molar ratio Rmin = 2.60between iron oxide and carbon. As the biomass is further dried,both the molar ratio R and the energy requirement E drop whileC increases. The results indicate that a lower moisture content inthe feedstock correlates to a lower solids circulation rate and en-ergy consumption. Vaporization of excess moisture consumes alarge amount of heat. The resulting steam consequently drivesthe equilibrium of Reaction (2) backwards. Therefore, more ironoxide is needed to fully oxidize the biomass. A similar effect isfound when injecting steam or CO2 as the gasifying agent. Fromthe iron oxide conversion standpoint, it is desirable to dry the bio-mass before entering the reducer and to limit the amount of gasi-fying agents injected to its lowest effective levels. Based on thesystem heat balance, the biomass is able to be reasonably driedto a 5% moisture content to limit excess water, and M = 5% is as-sumed in the following sections.

4.3. Products distribution

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

Fig. 7. Effect of temperature on biomass conversion at 1 atm. Fig. 8. Effect of pressure on biomass conversion at 900 �C.

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

Reactor type Gas–solid contactingpattern

Rmin x C x

Fluidized bed Mixed 6.28 0.0333 0.111 1.333Moving bed Countercurrent 1.61 0.1302 0.434 0.849

3780 F. Li et al. / Fuel 89 (2010) 3773–3784

Table 5. The model predicts that biomass is fully converted to CO2

and H2O while iron oxide is reduced by 43.4%, which is equivalentto an x of 0.849. The sulfur will exit the reducer with the gaseousstream in the form of SO2. Due to its low concentration in the givenfuel source, no sulfur will react with the oxygen carrier particles.Therefore, all the sulfur can be captured from the reducer exhaustgas stream using existing SOx control devices.

With all other conditions held constant, decreasing R will leadto a desirable reduction in x. However, when R is lower than1.61, the unconverted CO and H2 will exit from the reducer result-ing in an energy loss. Thus, an Rmin of 1.61 is the optimized valuefor the given conditions. Further studies also indicate that whenFe3O4 is used in the reducer feed instead of Fe2O3, a considerableamount of fuel is left unconverted due to the thermodynamic equi-librium limitation regardless of the R value. Therefore, it is neces-sary to regenerate the iron to Fe2O3 from the combustor beforeentering the reducer.

4.4. Effect of temperature

A sensitivity analysis is carried out using the multistage equilib-rium model. The carbon distribution in the reaction products witha constant pressure of 1 atm and varied temperature are shown inFig. 7. From a thermodynamic viewpoint, a higher reaction temper-ature favors the endothermic biomass–Fe2O3 reaction, as is alsopredicted by the model. In Fig. 7, using a Fe2O3/biomass ratioidentical to the case shown in Table 5 (M = 5%, R = 1.61), the carboncontent in the fuel source cannot be fully converted at tempera-tures below 900 �C unless the R value increases. The product distri-bution becomes even less favorable below 600 �C caused by theabsence of Fe0.947O at these relatively low temperatures. Therefore,the reducer should operate at a higher temperature to maximizeboth the fuel and oxygen carrier conversions. The operating condi-tion selected must also take into account the material costs forconstructing a reactor able to withstand the proposed environ-ment, particle melting point, and the ash softening point. A suit-able operating temperature is determined to be approximately900 �C and is used to estimate the other operating temperaturesin the following process simulation.

4.5. Effect of pressure

The reactions occurring in the reducer generate gaseous prod-ucts, adding to the overall system volume. Thus, a higher operating

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, anoperating pressure between 1 atm and 30 atm is determined as asuitable range for the BDCL reducer operation. However, consider-ing the energy required for CO2 compression, it is desirable to oper-ate the reducer at moderate pressures if CO2 is going to besequestrated.

4.6. Effect of gas–solid contacting pattern

The flow patterns of the gases and solids play an important rolein the reducer operation. In a moving bed reactor, both phases be-have in a plug flow manner and interact with each other counter-currently. This design has been studied in the multistage equilib-rium model for the BDCL process. In contrast, a well mixed flowpattern occurs in a fluidized bed reactor, which can be studiedusing a one stage equilibrium model. It should be noted that theactual fluidized bed operation is often worse than the mixed flow.While maintaining the operating conditions, of M = 5%, 1 atm and900 �C, constant, the two reducer models are compared. The resultsare illustrated in Table 6. The comparison shows that a fluidizedbed reducer demands a 2.9 times greater Fe2O3 flow rate than amoving bed reactor in order to fully convert the same amount ofbiomass feedstock. As a result, the oxygen carrying capacity andiron oxide conversion in the moving bed reducer are approxi-mately 2.9 times larger than in the alternative design. In the fluid-ized bed, the value for the performance variable, x, is constrainedbetween 1.333 and 1.5, which is incapable of hydrogen generationdue to thermodynamic limitations. On the contrary, a moving beddesign enables x to reach values as low as 0.849, suitable for bothhydrogen production and electricity generation. Therefore, themoving bed reducer delivers a more promising performance thanits fluidized bed counterpart based on the thermodynamic analy-sis. Similarly, the moving bed oxidizer behaves better than the flu-idized bed oxidizer under the same assumptions.

Fig. 9. Steam conversion at varying temperatures and x values in the oxidizer at30 atm.

F. Li et al. / Fuel 89 (2010) 3773–3784 3781

Thus, by using the multistage model simulation, the operatingconditions 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 con-tent x of 0.849.

4.7. Oxidizer performance

A similar five stage model is used to analyze the performance inthe moving bed oxidizer. Fig. 9 summarizes the simulation resultsof the oxidizer with an iron oxide feed FeOx from the outlet of thereducer at various temperatures. As noted earlier, x refers to thenormalized iron oxidation state from the moving bed reducer.The steam flow rate at the inlet of the oxidizer is taken as the min-imum amount that can fully oxidize FeOx to Fe3O4. The modelingshows that the steam to hydrogen conversion decreases withincreasing operating temperatures, which is consistent with theexothermic nature of the steam–iron reaction. The performancerelation between the reducer and the oxidizer is also illustratedin Fig. 9. As shown, for example at 700 �C, the steam to hydrogenconversion is constant for the values of x between 0 and 0.61,and it starts to decrease when x exceeds 0.61. The steam conver-sion is near zero when x approaches 1.33, representing the outletproduct of the oxidizer, Fe3O4. To summarize, when operated at

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

Reactor Feedstock condition Operating pressure, Operating temperature

Reducer Dried poplar 1 atm 900 �C25 �C, 1 atm

Oxidizer Steam 30 atm 850 �C240 �C, 32 atm

Combustor Compressed air 16 atm 1324 �C126 �C, 18 atm

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

Unit operations Input Outpu

Air compressor H2 compressor Pulverizer Other HP st

Power, MW 5.52 0.99 2.86 2.21 �1.74

the same temperature, a lower iron oxidation state (or smaller x)from the reducer outlet results in a higher steam to hydrogen con-version in the oxidizer. The condition of small x is highly desiredsince a higher steam conversion leads to reduced steam usage forproducing unit amount of hydrogen. Hence, the energy loss forthe steam generation is decreased. Thus, the operation of the BDCLprocess requires careful management of the circulating oxygencarrier particles.

For the entrained bed combustor, one RGIBBS block is used forthis simulation. With these reactor models, thermodynamic limita-tions for the BDCL system can be further clarified, and the resultswill be adopted in the later overall process simulation.

4.8. Process performance

The BDCL process can convert the energy in biomass into hydro-gen and/or electricity. In this section, the mass and energy flowsthroughout the process are studied for three cases. The reactormodeling results from the previous sections are integrated intothe BDCL process simulation to demonstrate the optimum opera-tion of the BDCL system.

4.9. Hydrogen and electricity coproduction

Using the process configuration discussed previously, the BDCLprocess is capable of cogenerating any combination of electricityand hydrogen while maintaining a self-sustained operation. Forthe case when a = 0.657, the simulation results are listed in Table7. The power balance for the overall process is summarized in Ta-ble 8. Under this scheme, the oxygen carrier particles comprising66.2% Fe2O3 and 33.8% inert SiC, are introduced from the top ofthe reducer at a flow rate of 290.84 tonne/h. The solid oxygencarriers exit the unit in the form of Fe and Fe0.947O (wüstite) whilethe biomass is fully converted. Any sulfur within the fuel feedstockis removed with the gas stream as SO2. To maintain the overall heatbalance for this case, the performance variable, a, is taken as 0.657.This value for a implies that, among the reduced iron particlesleaving the reducer, 65.7% is directed to the oxidizer for hydrogenproduction, while the remaining 34.3% is sent to the combustor di-rectly for heat generation. The steam–iron reaction in the oxidizeris carried out at 850 �C and 30 atm. The steam flow rate is set at25.52 tonne/h; it yields a steam conversion of 54.7%, as predictedby the multistage oxidizer model. The combustor regenerates theiron particles to Fe2O3 at 1324 �C, 16 atm. Using the chemical loop-ing scheme, the poplar is able to be fully converted into CO2 fromthe reducer (consisting of mainly 44% CO2 and 56% steam) whilehigh purity H2 is produced from the oxidizer (about 55% H2 and

Mole fraction in products (excluding ash) Conversion

Gas: 0.4365 H2O, 0.5623 CO2, 806 ppm N2, 373 ppm SO2 Biomass: 100%Solid: 0.097 Fe, 0.410 Fe0.947O, 0.493 SiC Fe2O3: 43.4%Gas: 0.453 H2O, 0.547 H2 Steam: 54.7%Solid: 0.247 Fe3O4, 0.753 SiC To Fe3O4: 100%Gas: 0.01 O2, 0.989 N2, 509 ppm NO O2: 96.2%Solid: 0.33 Fe2O3, 0.67 SiC To Fe2O3: 100%

t Net power

eam turbine IP steam turbine LP steam turbine Expander

�3.09 �1.99 �10.31 �5.55

3782 F. Li et al. / Fuel 89 (2010) 3773–3784

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

4.10. Maximized hydrogen production

For maximum hydrogen production from the BDCL process, theheat released for the steam generation and power productionneeds to be minimized. One effective approach is to preheat theair to 800 �C before entering the combustor using the HRSG unit.By doing so, the amount of solids required for direct combustionfrom the reducer can be reduced to 20.9%. The remainder can thenbe used for hydrogen production in the oxidizer. This configurationgenerates 5.71 tonne/h of high pressure steam for steam turbine

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

Fig. 11. Power generation efficiency and operating te

utilization, with a net electricity yield decreased to 0.15 MW.1.88 tonne/h of hydrogen is produced from this process at60 atm, corresponding to a total HHV hydrogen production effi-ciency of 74.2%.

4.11. Maximized power generation

By sending all the reduced iron oxides directly to the combus-tor, implying that the oxidizer is eliminated from the process, thepower generation is maximized. With this process scheme, theoperating temperature within the BDCL system is raised by�200 �C. Also, 72.5% excess air is used to maintain the suitableoperating temperature. Another method to maintain temperaturecan be achieved by recycling the depleted air back to the combus-tor as a coolant. The overall electricity generation efficiency usingthis method can reach approximately 38.1%. For power generation,the combustor can be also integrated with supercritical or ultra-supercritical (USC) boiler technology, which eliminates the expan-der after the combustor. Other schemes could include producinghydrogen first and then generating electricity from hydrogeneither through combustion or commercial fuel cells.

As deduced from the three cases, a performance curve relativeto the variable a is plotted in Fig. 10. At a given biomass feedstockand suitable operating conditions, the performance variable a lin-early determines the end product distribution and correspondingprocess efficiency. In the presented process scheme discussedthroughout the article, the range of a is limited between 0 and0.79 by the heat and work balance. As a decreases, i.e. morereduced iron particles flow to the combustor, more heat will bereleased for electricity generation with less hydrogen yield. Theflexibility between hydrogen production and electricity generationmakes the BDCL process versatile in product selection when com-pared to other biochemical technologies.

Compared to the current thermochemical processes given in Ta-ble 1, the BDCL process efficiency for biomass conversion is greaterin both hydrogen production and electricity generation cases. Theincreased efficiencies result from an effective energy managementsystem and a unique looping reaction scheme. The high grade heatreleased from the circulating oxygen carriers significantly reducethe inefficient generation of low pressure steam. Additionally, theBDCL system simplifies the traditional biomass gasification conver-sion processes by removing the gasifier, air separation unit (ASU)and water gas shift reactor for hydrogen production. Further, elab-orate CO2 separation steps are not required. Therefore, through

mperature range as a function of inert content.

F. Li et al. / Fuel 89 (2010) 3773–3784 3783

process intensification, the BDCL process has a lower capital costand parasitic energy requirements.

4.12. Effect of the inert

The inert is used to moderate the operating temperature of theBDCL system. Once the process inputs and outputs are fixed, theoverall heat of reaction is determined. By increasing the amountof inert, the temperature differences among reactors can poten-tially be reduced since the heat capacity of the inert material canhelp decrease the required temperature gradients. In the presentcase, all the reduced oxygen carriers are directly combusted(a = 0) under the BDCL-USC configuration. Under this configura-tion, the combustor is operated at a temperature that ensures en-ough heat to be carried over to the reducer for the subsequentbiomass–oxygen carrier reaction. The excess heat from the BDCLsystem is recovered by an integrated USC boiler. The steam speci-fication of the USC boiler is 650 �C and 300 atm. Simulation resultsshow that the adiabatic combustor temperatures and process effi-ciency vary notably with the mass percentage of inert SiC. Asshown in Fig. 11, the operating temperature range decreases withincreasing inert content whereas the process efficiency shows anopposite trend. The operating temperature can be further loweredby introducing excess air into the entrained bed combustor. Basedon the modeling results, 30–70% inert content is a suitable rangefor the BDCL process. Additionally, the inert compositions can alsoserve as structural support to minimize attrition and promote reac-tion kinetics.

4.13. Pollutants control in the BDCL process

By capturing CO2 from the reducer, the BDCL process becomescarbon negative. Compressing CO2 to 150 atm for sequestration re-quires �3% of the total thermal energy input. The ash in biomasscan be separated from the oxygen carrier based on size differencesusing a cyclone. Biomass may also contain sulfur. Through reactormodeling and process simulation, the fate of sulfur can be deter-mined. Poplar has a sulfur content of less than 0.21 wt.% (dry ba-sis). With such low sulfur content, all the sulfur in the biomasswill be converted to SO2 in the reducer. For biomass sources withhigher sulfur contents, Fe0.877S may form in the reducer. The pres-ence of Fe0.877S will lead to H2S contamination to the hydrogenstream from the oxidizer. Using conventional sulfur control de-vices, SO2 and H2S can be captured and converted to elemental sul-fur. NOx can also be formed in the combustor. The flamelesscombustion environment and the absence of fuel nitrogen in thecombustor may lead to NOx level significantly lower than that ina conventional coal fired boiler.

5. Concluding remarks

The biomass direct chemical looping (BDCL) process is simu-lated, analyzed, and discussed in this study. The process consistsof three chemical looping reactors, i.e. the reducer, oxidizer, andcombustor, with an iron based particle circulating among the reac-tors. The iron based particle is used as the carrier for oxygen, chem-ical energy, and thermal energy, facilitating efficient conversion ofbiomass to hydrogen and/or electricity. The four oxidation states ofiron and their distinct thermodynamic properties give rise to alarge degree of freedom in the process design and operatingparameters. By using a thermodynamic analysis tool, i.e. ASPENPlus�, the suitable reactor design, operating conditions, and pro-cess configuration for the BDCL process are determined.

A multistage model is developed in this study to simulate theperformance of individual BDCL reactor based on ASPEN Plus�.

The simulation results indicate that a countercurrent gas–solidflow pattern such as that provided by a countercurrent movingbed reactor maximizes the conversions of both solids and gas. Theyalso indicate that a higher moisture content in the biomass feed-stock leads to a lower oxygen carrier conversion in the reducer,lower steam to hydrogen conversion in the oxidizer, higher solidscirculation rate, and lower process energy conversion efficiency.Thus, maintaining the moisture content in the biomass feedstockat 5 wt.% or lower is desirable prior to the chemical looping reac-tions. The conversions of the biomass and iron based particles inthe reducer are also favored at higher temperatures and lowerpressures. Considering various practical factors such as reactormaterials and gas compression, the suitable operating temperaturefor the reducer is �900 �C with pressures between 1 and 30 atm.The steam to hydrogen conversion in the oxidizer is more favoredat lower temperatures. The overall energy balance of the process,however, requires the oxidizer to be operated at above 600 �C. Tocompensate for the heat required in the reducer, the combustoris operated at temperatures 100–450 �C above the reducer.

The process simulation based on the reactor modeling resultsfurther reflects that the BDCL process can produce hydrogen and/or electricity at any ratio. Compared to conventional biomass com-bustion and gasification processes, the BDCL process is 10–25%more efficient. The sulfur and NOx pollutants from the processcan be readily removed using commercial pollutant control de-vices. Additionally, the CO2 stream generated from the BDCL pro-cess is of a high concentration. With the CO2 sequestration, theBDCL becomes a carbon negative process from the life cycle stand-point. The unique energy management and conversion scheme ofthe BDCL process, thus, render it potentially viable for clean andefficient biomass conversion applications.

Acknowledgement

The authors would like to acknowledge Dr. Nobusuke Kobayashi,Mr. Andrew Tong, and Mr. Eric Sacia for their helpful comments onthe 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. Biomassas feedstock for a bioenergy and bioproducts industry: the technical feasabilityof a billion-ton annual supply. US DOE and USDA; 2005. p. 16–8.

[3] Sofer SS, Zaborsky O. Biomass conversion processes for energy andfuels. Plenum Press; 1981.

[4] Gerbens-Leenes W, Hoekstra AY, van der Meer TH. The water footprint ofbioenergy. 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. EnergConvers 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. EnergFuels 2009;23(8):4182–9.

[11] Gupta P, Velazquez-Vargas LG, Fan L-S. Syngas redox (SGR) process to producehydrogen from coal derived syngas. Energ Fuels 2007;21:2900–8.

[12] Thomas T, Fan L-S, Gupta P, Velazquez-Vargas LG. Combustionlooping usingcomposite oxygen carriers, US patent appl publ 2005; 2005-175533-A1.

[13] Cho P, Mattisson T, Lyngfelt A. Comparison of iron-, nickel-, copper- andmanganese-based oxygen carriers for chemical-looping combustion. Fuel2004;83(9):1215.

[14] Adanez J, Garcia-Labiano F, de Diego LF, Gayan P, Abad A, Celaya J. Developmentof oxygen carriers for chemical-looping combustion. Carbon dioxide capturestorage 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 ConfFluidized Bed Combust 2003;17:173.

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

[16] Kronberger B, Loeffler G, Hofbauer H. Chemical-looping combustion – reactorfluidization studies and scale-up criteria. Carbon dioxide capture storage deepgeologic formations–results CO2 capture project 2005;1:605.

[17] Richter HJ, Knoche KF. Reversibility of combustion processes. ACS Symp Ser1983;235:71.

[18] McGlashan NR. Chemical-looping combustion – a thermodynamic study. Proc IMech Eng Part C J Mech Eng Sci 2008;222(C6):1005–19.

[19] Xiang W, Chen Y. Hydrogen and electricity from coal with carbon dioxideseparation using chemical looping reactors. Energ Fuels 2007;21(4):2272.

[20] Cleeton JPE, Bohn CD, Mueller CR, Dennis JS, Scott SA. Clean hydrogenproduction and electricity from coal via chemical looping: identifying asuitable operating regime. Int J Hydrogen Energy 2009;34(1):1–12.

[21] Chiesa P, Lozza G, Malandrino A, Romano M, Piccolo V. Three-reactorschemical looping process for hydrogen production. Int J Hydrogen Energy2008;33(9):2233–45.

[22] Cao Y, Casenas B, Pan W-P. Investigation of chemical looping combustion bysolid fuels. 2. Redox reaction kinetics and product characterization with coal,biomass, and solid waste as solid fuels and CuO as an oxygen carrier. EnergFuels 2009;23(4):2307–15.

[23] Shen L, Wu J, Xiao J, Song Q, Xiao R. Chemical-looping combustion of biomassin a 10 kW h reactor with iron oxide as an oxygen carrier. Energ Fuels2009;23(5):2498–505.

[24] Leion H, Mattisson T, Lyngfelt A. CO2 capture from directcombustion of solid fuelswith chemical-looping combustion. Proceedings of the 33rd internationaltechnical conference on coal utilization and fuel systems, vol. 1; 2008: p. 221–32.

[25] Gupta P, Velazquez-Vargas LG, Li F, Fan L-S. Chemical looping reformingprocess for the production of hydrogen from coal. Proceedings of 23rd annualinternational Pittsburgh coal conference, 2006; p. 37 3/1–3/8.

[26] Gnanapragasam NV, Reddy BV, Rosen MA. Hydrogen production from coalusing coal direct chemical looping and syngas chemical looping combustionsystems: assessment of system operation and resource requirements. Int JHydrogen Energy 2009;34(6):2606–15.

[27] Phillips S, Aden A, Jechura J, Dayton D, Eggeman T. Thermochemical ethanolvia indirect gasification and mixed alcohol synthesis of lignocellulosicbiomass. NREL 2007. p. 19.

[28] Li F, Zeng L, Velazquez-Vargas LG, Yoscovits Z, Fan L-S. Syngas chemicallooping gasification process: bench scale studies and reactor simulations.AICHE J 2009;56(8):2186–99.