Chemical-Looping Combustion of Biomass in a 10 kWth Reactor with

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Chemical-Looping Combustion of Biomass in a 10 kWth Reactor with Iron Oxide As an Oxygen Carrier Laihong Shen,* Jiahua Wu, Jun Xiao, Qilei Song, and Rui Xiao Thermoenergy Engineering Research Institute, Southeast UniVersity, Nanjing 210096, China ReceiVed January 14, 2009. ReVised Manuscript ReceiVed March 1, 2009

Chemical-looping combustion of biomass was carried out in a 10 kWth reactor with iron oxide as an oxygen carrier. A total 30 h of test was achieved with the same batch of iron oxide oxygen carrier. The effect of the fuel reactor temperature on gas composition of the fuel reactor and the air reactor, the proportion of biomass carbon reacting in the fuel reactor, and the conversion of biomass carbon to CO2 in the fuel reactor was experimentally investigated. The results showed that the CO production from biomass gasification with CO2 was more temperature dependent than the CO oxidation with iron oxide in the fuel reactor, and an increase in the fuel reactor temperature produced a higher increase for the CO production from biomass gasification than for the oxidation of CO by iron oxide. Although the conversion of biomass carbon to CO2 in the fuel reactor decreased with the increase of the fuel reactor temperature, there was a substantial increase in the proportion of biomass carbon reacting in the fuel reactor. X-ray diffraction (XRD) and scanning electron microscopy (SEM) were utilized to characterize fresh and reacted oxygen carrier particles. The results showed that the transformation of Fe2O3 to Fe3O4 is the favored step in the process of iron oxide reduction with biomass syngas. The low reactivity of reacted oxygen carrier was mainly ascribed to the sintering grains on the particle surface. To restrain the surface sintering of oxygen carrier particles, an intensive oxidization of reduced oxygen carrier with air in the air reactor should be avoided in the process of oxygen carrier regeneration, and air staging should be adopted for the oxidization of reduced oxygen carrier with air in the air reactor.

1. Introduction Chemical-looping combustion (CLC) is one of the techniques used to combine fuel combustion and CO2 capture. The properties of the oxygen carrier are vital for the practice of the process of CLC, and the following properties should be provided by oxygen carrier: high reactivity, high oxygen transport capacity, high resistance to attrition, and complete fuel conversion to CO2 and H2O. Moreover, it is also an advantage if the metal oxide is cheap and environmentally benign. Different metal oxides, such as Ni, Fe, Cu, Mn, and Co, have been proposed for CLC. Most of them have only been tested in a fluidized bed reactor or thermogravimetric analyzers (TGA) using gaseous fuels such as methane, hydrogen, and syngas from coal gasification.1-8 Copper oxide shows high reactivity in the reduction and regeneration processes in thermobalance tests. There have been difficulties with CuO agglomerating due to its low melting point. A further problem at high temperatures (>870 °C) is that CuO can decompose to Cu2O, which has a lower oxygen capacity for oxidizing fuel. Oxides of Mn and Co * To whom correspondence should be addressed. Telephone: +86-258379 5598. Fax: +86-25-5771 4489. E-mail: [email protected]. (1) Mattisson, T.; Lyngfelt, A.; Cho, P. Fuel 2001, 80 (13), 1953–1962. (2) Jin, H. G.; Ishida, M. Int. J. Hydrogen Energy 2001, 268, 889–894. (3) Ishida, M.; Yamamoto, M.; Ohba, T. Energy ConVers. Manage. 2002, 43 (9-12), 1469–1478. (4) Cho, P.; Mattisson, T.; Lyngfelt, A. Fuel 2004, 83 (9), 1215–1225. (5) de Diego, L. F.; Garcıá-Labiano, F.; Adańez, J.; Gayań, P.; Abad, A.; Corbella, B. M.; Maria Palacios, J. Fuel 2004, 83 (13), 1749–1757. (6) Adańez, J.; de Diego, L. F.; Garcıá-Labiano, F.; Gayań, P.; Abad, A.; Palacios, J. M. Energy Fuels 2004, 18 (2), 371–377. (7) Abad, A.; Mattisson, T.; Lyngfelt, A.; Ryden, M. Fuel 2006, 85 (9), 1174–1185. (8) Son, S. R.; Kim, S. D. Ind. Eng. Chem. Res. 2006, 45 (8), 2689– 2696.

showed a rather poor reactivity.9 Generally, Ni and its corresponding oxide show relatively higher oxidation and reduction rates than other metals. However, it is unavoidable for some of oxygen carrier particles to get lost with ash due to breakage and fragmentation of oxygen carrier particles in the fluidized bed reactors for CLC of solid fuel, which means that the oxygen carrier particles may have to be frequently replaced. Besides, nickel oxide is toxic and expensive, the replacement of nickel-based oxygen carrier would increase the cost of the CLC process, and its leakage would cause environmental problem.10 Therefore, it is necessary to reduce the use of nickel-based oxygen carrier and develop alternative oxygen carriers that are more suitable for the CLC processes of solid fuels. Iron oxide is cheaper than NiO; the melting points of all involved iron compounds in the two stages of oxidization and reduction are very high.11 Thermodynamic analysis predicted that complete fuel conversion to CO2 and H2O is only possible when Fe2O3 is partially reduced to Fe3O4.12,13 Although the capacity of Fe2O3 for transferring oxygen is only 0.5 mol of O2 delivered per 3 mol of Fe2O3, iron oxide is still an attractive (9) Mattisson, T.; Jardnas, A.; Lyngfelt, A. Energy Fuels 2003, 17 (3), 643–651. (10) Leion, H.; Mattisson, T.; Lyngfelt, A. Int. J. Greenhouse Gas Control 2008, 2 (2), 180–193. (11) Corbella, B. M.; Palacios, J. M. Fuel 2007, 86 (1-2), 113–122. (12) Mattisson, T.; Lyngfelt, A. Capture of CO2 Using ChemicalLooping Combustion. In First Biennial Meeting of the ScandinaVian-Nordic Section of the Combustion Institute, Go¨teborg, Sweden, 2001; pp 163168. (13) Scott, S. A.; Dennis, J. S.; Hayhurst, A. N.; Brown, T. AIChE J. 2006, 529, 3325–3328.

10.1021/ef900033n CCC: $40.75  2009 American Chemical Society Published on Web 03/30/2009

Chemical-Looping Combustion of Biomass

oxygen carrier for the commercial application of CLC because of its easy availability and low price.14 Leion et al. investigated the feasibility of CLC of petroleum coke using Fe2O3 as an oxygen carrier in a fluidized bed reactor. The gasification reaction of petroleum coke was slow compared to the reaction of the metal oxide with the gasification product.15 However, most past research on iron oxide as an oxygen carrier has been performed in thermogravimetric analyzers (TGA) or a laboratory fluidized bed reactor, and these tests of the oxygen carrier reactivity covered only a limited number of cycles, and the reactivity of the oxygen carrier progressively decayed with the number of cycles due to accumulative chemical and thermal stresses. To further reveal the reactivity behavior and usefulness of iron oxide as an oxygen carrier in the CLC process, more tests are needed in a real system where the oxygen carrier particles are continuously circulated between an air reactor and a fuel reactor. There are in the open literature a limited number of studies focused on using oxygen carriers for CLC with solid fuel in a continuous CLC reactor.16-18 Berguerand and Lyngfelt investigated a CLC process of a petroleum coke as the solid fuel and the ilmenite oxygen carrier in a 10 kWth chemicallooping combustor. The effects of particle circulation and carbon stripper operation on solid fuel conversion, conversion of gas from the fuel reactor, and CO2 capture were investigated, and the actual CO2 capture ranged between 60% and 75%. In an earlier study, experiments on chemical-looping combustion of coal with a NiO-based oxygen carrier in a 10 kWth prototype was reported, and no major difficulties in operation were encountered.16 Although a small amount of CO and CH4 was present in the gas outlet of the fuel reactor, a high conversion of coal to CO2 was achieved. In the present work the performance of chemical-looping combustion of biomass was experimentally investigated in the 10 kWth reactor with iron oxide as an oxygen carrier. A total 30 h of test was achieved with the same batch of iron oxide oxygen carrier. The effect of the fuel reactor temperature on gas composition of the fuel reactor and the air reactor, the proportion of biomass carbon reacting in the fuel reactor, and the conversion of biomass carbon to CO2 in the fuel reactor was experimentally investigated. X-ray diffraction (XRD) and scanning electron microscopy (SEM) were utilized to characterize fresh and reacted oxygen carrier particles. 2. Experimental Section The experiments for CLC of biomass with iron oxide as an oxygen carrier were conducted in a 10 kWth continuous reactor of interconnected fluidized beds (Figure 1). The prototype is composed of a fast fluidized bed as air reactor, a cyclone, and a spout-fluid bed as fuel reactor. The fast fluidized bed is a circular column with a 50 mm i.d. and 2000 mm height with a perforated plate as an air distributor. To fulfill a high biomass conversion efficiency, a spout-fluid bed, instead of a bubbling fluidized bed, is adopted for the fuel reactor because of strong solids mixing and long residence time of biomass particles in the spout-fluid bed. Internal circulation of oxygen carrier particles through the spout-fluid bed is aimed at delivering a better (14) Chuang, S. Y.; Dennis, J. S.; Hayhurst, A. N.; Scott, S. A. Combust. Flame 2008, 154 (1-2), 109–121. (15) Leion, H.; Mattisson, T.; Lyngfelt, A. Fuel 2007, 86 (12-13), 1947– 1958. (16) Shen, L.; Wu, J.; Xiao, J. Combust. Flame 2009, 156 (3), 721– 728. (17) Berguerand, N.; Lyngfelt, A. Int. J. Greenhouse Gas Control 2008, 2 (2), 169–179. (18) Berguerand, N.; Lyngfelt, A. Fuel 2008, 87 (12), 2713–2726.

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Figure 1. Configuration for chemical-looping combustion of biomass in interconnected fluidized beds.

distribution of biomass, countering its tendency to segregate to the bed surface; the elutriation of fine char particles is also reduced. These factors all favor biomass conversion to CO2 and H2O in the fuel reactor. The spout-fluid bed is a rectangular bed, with a cross section of 230 × 40 mm2 and a height of 1500 mm. A 60° conical distributor connected with a tube with a 20 mm i.d. is mounted at the bottom of the spout-fluid bed. The tube is used to introduce the fluidization stream with biomass to the bottom of the spoutfluid bed. There exists a specially designed configuration inside the spout-fluid bed. The configuration causes the spout-fluid bed to have two compartments; the major compartment is termed as the reaction chamber, and the minor one is the inner seal. The reaction chamber allows the combination of biomass gasification and oxygen carrier reduction with biomass syngas to proceed inside the spout-fluid bed. The inner seal with a cross section of 23 × 40 mm2 allows particles (including reactant) movement from the reaction chamber to the fast fluidized bed and prevents the bypassing of the flue gas from the fast fluidized bed to the spout-fluid bed. The heat required for the combined process of biomass gasification and oxygen carrier reduction with biomass syngas is achieved by means of the external circulation of oxygen carrier particles between the fast fluidized bed and the spout-fluid bed. For the process of chemical-looping combustion of biomass in the interconnected fluidized beds, the whole reduction of oxygen carrier with biomass in the spout-fluid bed is an intensive endothermic process, while the oxidization of oxygen carrier with air in the air reactor is exothermic. There is a temperature difference between the fast fluidized bed and the spout-fluid bed. With the external circulation of oxygen carrier particles, oxygen carrier particles are heated up in the fast fluidized bed and then release heat in the spout-fluid bed. Thus, oxygen carrier particles also serve as heat carrier to transfer heat from the fast fluidized bed to the spout-fluid bed. Part of the CO2 stream from the flue gas is extracted from the spout-fluid bed and circulated into its bottom, serving as the

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fluidization medium and biomass gasification agent in the spoutfluid bed. An intensive contact between biomass and hot oxygen carrier particles will occur the moment biomass is pneumatically conveyed into the bottom of the spout-fluid bed with the circulated flue gas, followed by the intense exchange of heat and mass. Fresh biomass is immediately heated up to the bed temperature (750-950 °C), and thereby devolatilization and gasification of biomass take place in the lower zone of the spout-fluid bed. The reactions for biomass gasification in the spout-fluid bed are illustrated as follows

pyrolysis

biomass f char + tar + gases (H2, CO, CO2, CH4, CnH2m) (R1) Boudouard

C + CO2 f 2CO

(R2)

The reduction of iron oxide with biomass is a heterogeneous reaction. The direct contact efficiency of solid-phase reaction in the spout-fluid bed is very low. Comparison to the reduction of iron oxide particles with biomass syngas, the reduction rate for iron oxide particles with biomass is very much slower at the temperature of interest (750-950 °C). Hence, the reduction of iron oxide particles directly by biomass is not practical. This indicates that the syngas of biomass gasification should be the major reductive agents for the reduction of iron oxide particles in the spout-fluid bed. The main reductions of iron oxide particles with biomass syngas in the spout-fluid bed are as follows

CO + 3Fe2O3 f 2Fe3O4 + CO2

(R3)

CO + Fe3O4 f 3FeO + CO2

(R4)

CO + FeO f Fe + CO2

(R5)

H2 + 3Fe2O3 f 2Fe3O4 + H2O

(R6)

H2 + Fe3O4 f 3FeO + H2O

(R7)

H2 + FeO f Fe + H2O

(R8)

CH4 + 12Fe2O3 f 8Fe3O4 + CO2 + 2H2O

(R9)

CH4 + 4Fe3O4 f 12FeO + CO2 + 2H2O

(R10)

CH4 + 4FeO f 4Fe + CO2 + 2H2O

(R11)

Simultaneously, the reduction of iron oxide with biomass syngas remarkably accelerates the process of biomass gasification. The flue gas exhausting from the spout-fluid bed is a nearly nitrogen-free gas mixture with CO2 and possibly CO and CH4. The particles of reduced oxygen carrier in the spout-fluid bed return back to the fast fluidized bed via an overflow pipe. Each of the two beds is electrically heated individually, which supplies heat for start-up and compensated heat loss during a continuous operation. At the beginning of a test, the heating of the two beds is carried out. When the temperature of the two beds is electrically heated to 900 °C, both the fast fluidized bed and the spout-fluid bed are fluidized with air. As the external particles circulation is stably reached, the spout-fluid bed is fluidized with a CO2 stream instead of air. In the present study the CO2 stream is supplied with a commercial CO2 tank instead of the recycling gas extracted from the flue gas of the spout-fluid bed. The oxygen carrier is prepared from iron oxide as powders. To improve the mechanical properties of the iron oxide particles, the original particles of iron oxide are sintered at 1000 °C for 3 h. The calcinated particles of iron oxide are then sieved into a size fraction of 0.3-0.6 mm and then used as the fresh oxygen carrier in this study. The particles of fresh oxygen carrier with a density of 2460 kg/m3 belongs to Group B in the Geldart classification and having a minimum fluidization velocity Umf ) 0.60 m/s under atmospheric

Table 1. Proximate Analysis and Ultimate Analysis of Biomass proximate analysis (wt %)

ultimate analysis (wt %)

moisture 11.89 fixed carbon 14.77 volatile 75.78 ash 1.56 low heating value (MJ/kg) 14.47

C 40.06 H 5.61 O 39.88 N 0.90 S 0.10

pressure. In the experiments, the total weight of iron oxide particles as oxygen carrier employed in the experiments is about 12 kg. The biomass used in the study is pine sawdust from Jiangsu, China. The proximate analysis and ultimate analysis of the biomass are illustrated in Table 1. The Sauter mean diameter of biomass used in the experiments is about 1.5 mm. By means of a variablespeed screw feeder, biomass is pneumatically conveyed to the bottom of the spout-fluid bed with the CO2 stream in a volume flow controller. The flue gases of the two reactors are respectively induced with a suction pump to an ice-water cooler where the steam is condensed and removed. The flue gases are sampled by gas bags for off-line analysis after each experimental run reached relatively steady situation for 2 h, which is found to be sufficiently long to reflect the process of CLC of biomass in the interconnected fluidized bed. The composition of the flue gas is measured using a NGA2000 type gas analyzer (EMERSON Co.).

3. Results and Discussion To investigate the performance of CLC of biomass with iron oxide as an oxygen carrier, a relatively long-term experiment of 30 h was accomplished with the same batch of iron oxide particles, i.e., without adding any fresh oxygen carrier particles. In the experiment the feed rate of biomass was kept at 3.0 kg/ h. The air flow of the fast fluidized bed was kept at 11.0 m3/h. The CO2 stream flow of the spout-fluid bed was kept at 2.0 m3/h. To get rid of the effect of the surface microstructure of the oxygen carrier particles on the performance of CLC of biomass in the interconnected fluidized beds, a preliminary experiment for CLC of biomass was conducted with fresh iron oxide as an oxygen carrier at a fuel reactor temperature of 920 °C, and a continuous period of 10 h was accomplished. Afterward the fuel reactor temperature was varied from 920 to 740 °C, and the effect of the fuel reactor temperature on the performance of CLC of biomass was investigated with the reacted oxygen carrier. There was some gas leakage from the fuel reactor to the air reactor with the external circulation of oxygen carrier particles between the fast fluidized bed and the spout-fluid bed. Therefore, during each run of the fuel reactor temperature the gas leakage experiment was first conducted without biomass fed into the fuel reactor, and the CO2 leakage flux from the fuel reactor to the air reactor was estimated based on the measured data of the gas composition of the exit gas from the air reactor. After the end of the gas leakage experiment, 3.0 kg/h of biomass was fed into the fuel reactor, and a relatively steady process for chemical-looping combustion of biomass was kept for 2 h in the interconnected fluidized beds. 3.1. Gas Composition of Both Fuel Reactor and Air Reactor. The fuel reactor temperature is crucial for chemicallooping combustion of biomass in the interconnected fluidized beds. The fuel reactor temperature was varied from 740 to 920 °C in the present experiment, and results were presented to demonstrate the effect of the fuel reactor temperature on the gas composition of both the fuel reactor and the air reactor. The exit gas composition of the fuel reactor after water condensation is shown as a function of the fuel reactor temperature, as indicated in Figure 2. H2 was not detected in the flue gas of the fuel reactor. The concentration of CH4 in the



Figure 3. Effect of fuel reactor temperature on gas composition of the air reactor.

Figure 2. Effect of fuel reactor temperature on gas composition of the fuel reactor. (a) CO and CH4 concentrations in the flue gas of fuel reactor. (b) CO2 concentration in the flue gas of fuel reactor.

flue gas of the fuel reactor was kept at about 2.7% in the fuel reactor temperature range from 740 to 920 °C. However, CO concentration of the flue gas of the fuel reactor significantly increased with an increase of the fuel reactor temperature. Correspondingly, CO2 concentration of the flue gas of the fuel reactor was maintained around 75.5-87.3%, as shown in Figure 2b. The effect of the fuel reactor temperature on the flue gas composition of the air reactor was also investigated, as indicated in Figure 3. Oxygen concentration in the flue gas of the air reactor was maintained around 6.9-10.0%. It increased with the fuel reactor temperature. The reason is that CO in the flue gas of the fuel reactor significantly increased with the fuel reactor temperature, as shown in Figure 2a. It meant that the amount of oxygen carrier particles reduced by CO in the fuel reactor decreased with the fuel reactor temperature. The particles of reduced oxygen carrier were transferred from the fuel reactor into the air reactor with the external circulation of oxygen carrier particles. It led to the decrease of oxygen consumed by the reduced oxygen carrier particles in the air reactor and the increase of O2 concentration in the exit gas of the air reactor. There was some CO2 in the flue gas from the air reactor. It was composed of both the bypassing of CO2 gas from the spoutfluid bed to the fast fluidized bed and the product of residual char burnt with air in the fast fluidized bed. The residual char with the external circulation of oxygen carrier particles came into the fast fluidized bed from the spout-fluid bed. CO2 in the

flue gas of the air reactor decreased with an increase of the fuel reactor temperature, as shown in Figure 3. Biomass gasification in the fuel reactor was highly strengthened with an increase of reaction temperature, and the amount of residual char in the fuel reactor decreased with an increase of the fuel reactor temperature. It resulted in the decrease of the residual char coming into the air reactor from the fuel reactor. Therefore, CO2 concentration of the exit gas from the air reactor decreased with the fuel reactor temperature. 3.2. Proportion of Biomass Carbon Reacting in the Fuel Reactor. The proportion of biomass carbon reacting in the fuel reactor (γC_Biomass, FR) is an important parameter to evaluate the performance of interconnected fluidized beds for chemical-looping combustion of biomass. It is defined as the ratio of biomass carbonaceous gas flux leaving the fuel reactor (FC_Biomass, FR) to total biomass carbon flux fed into the fuel reactor (FC_Biomass) γC_Biomass,FR )

FC_Biomass,FR FC_Biomass

FC_Biomass,FR ) FC_Biomass - FC_Biomass,AR

(1)

(2)

where FC_Biomass, AR is the biomass carbonaceous gas flux leaving the air reactor. FC_Biomass,AR can be calculated based on the carbonaceous gas flow leaving the air reactor with biomass feed (FC,AR) and the CO2 leakage flux from the fuel reactor into the air reactor without biomass feed (FCO2_Leakage). FC_Biomass,AR ) FC,AR - FCO2_Leakage

(3)

To give an indication of the carbonaceous gas flow leaving the air reactor with biomass feed, a mass balance is performed. Below is an overview of the measured data and how these can be used to calculate the carbonaceous gas flow leaving the air reactor with biomass feed. The CO2, CO, CH4, and N2 concentrations in the flue gas of the air reactor, noted (CO2)AR, (CO)AR, (CH4)AR, and (N2)AR, are given based on the measured data. As the molar air flux of the air reactor (FAir,AR) is known, the carbonaceous gas flow leaving the air reactor (FC,AR) can be calculated using the known nitrogen flux of the air flow in the air reactor as follows

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FC,AR ) (CO2)AR + (CO)AR + (CH4)AR × (FAir,AR × 0.79)(kmol/h) (N2)AR (4) Also, the CO2 leakage flux from the fuel reactor into the air reactor without biomass feed (FCO2_Leakage) can be calculated in the same way based on the measured data of the gas concentration of the air reactor in the gas leakage experiment. The above calculation assumes that the carbonaceous gas leaving the air reactor only contains CO2, CO, and CH4. Therefore, the proportion of biomass carbon reacting in the fuel reactor (γC_Biomass,FR) can be calculated with eqs 2-4. The proportion of biomass carbon reacting in the fuel reactor (γC_Biomass,FR) versus the fuel reactor temperature is indicated in Figure 4. There was a substantial increase in the proportion of biomass carbon reacting in the fuel reactor for the fuel reactor temperature ranges of 740-920 °C. The proportion of biomass carbon reacting in the fuel reactor was mainly attributed to the combined process of biomass gasification and oxygen carrier reduction with biomass syngas in the fuel reactor. Reaction R2 of biomass gasification is intensively endothermic and is highly strengthened by an increase of reaction temperature. The higher the fuel reactor temperature, the more biomass char is converted into gaseous product in the fuel reactor. Simultaneously, the reduction of iron oxide with biomass syngas remarkably accelerated the process of biomass gasification. Therefore, it resulted in a clear decrease of residual char entered into the air reactor from the fuel reactor with the external circulation of oxygen carrier particles and the increase of the proportion of biomass carbon reacting in the fuel reactor, as shown in Figure 4. However, the proportion of biomass carbon reacting in the fuel reactor only reached 95% at a relatively high temperature of the fuel reactor 920 °C. It indicated that there was an inherent residual char of biomass with the external circulation of oxygen carrier particles to enter into the air reactor from the fuel reactor, which was inevitable in this kind of CLC reactors with biomass fed into the bottom of the spout-fluid bed. 3.3. Conversion of Biomass Carbon to CO2 in the Fuel Reactor. The conversion of biomass carbon to CO2 in the fuel reactor (ηCO2) is related to actual CO2 capture efficiency for chemical-looping combustion of biomass in the interconnected fluidized beds. It is calculated based on the total carbonaceous gas flux leaving the fuel reactor (FC,FR), the CO2 concentration of the exit gas from the fuel reactor (CO2)FR, the CO2 stream flux fed into the fuel reactor (FCO2_Stream), and the CO2 leakage flux from the fuel reactor into the air reactor without biomass feed (FCO2_Leakage). The total carbonaceous gas flux leaving the fuel reactor (FC,FR) is the sum of biomass carbonaceous gas flux leaving the fuel reactor (FC_Biomass,FR) and the difference between the CO2 stream flux fed into the fuel reactor (FCO2_Stream) and the CO2 leakage flux from the fuel reactor into the air reactor without biomass feed (FCO2_Leakage). FC,FR ) FC_Biomass,FR + (FCO2_Stream - FCO2_Leakage) (5) Therefore, the conversion of biomass carbon to CO2 in the fuel reactor (ηCO2) can be calculated with eqs 2-6 ηCO2 )

FC,FR × (CO2)FR - (FCO2_Stream - FCO2_Leakage) FC_Biomass

(6)

Figure 4. Effect of fuel reactor temperature on the proportion of biomass carbon reacting in the fuel reactor.

Figure 5. Effect of fuel reactor temperature on the conversion of biomass carbon to CO2 in the fuel reactor.

The conversion of biomass carbon to CO2 in the fuel reactor ηCO2 versus the fuel reactor temperature is indicated in Figure 5. It was maintained around 53.7-65.1% at the fuel reactor temperature range of 740-920 °C, and it decreased with the increase of the fuel reactor temperature. The explanation could be that reactions R2 of biomass gasification in the spout-fluid bed was significantly strengthened with a temperature increase, and the CO production from biomass gasification was more temperature dependent than the CO oxidation with iron oxide in the spout-fluid bed. Thus, an increase in the fuel reactor temperature produced a higher increase for the CO production from biomass gasification than for the consumption of CO oxidation to CO2. It resulted in the decrease of the conversion of biomass carbon to CO2 in the fuel reactor with an increase of the fuel reactor temperature. 3.4. Phase Characterization of Oxygen Carriers. At the end of the whole experiment, the air flow of the fast fluidized bed, CO2 stream of the spout-fluid bed, and biomass feed were stopped. The two beds were kept in a closed and quiescent state. Afterward the particles in the two beds were cooled to room temperature, and typical particles were sampled and sealed for characterization analysis. Several samples of the reacted oxygen carrier were regenerated with pure air for 1 h in a muffle oven of 1000 °C. The regeneration of reacted oxygen carrier in the highly oxidative atmosphere was regarded as a complete oxidization process. The BET surface area and total pore volume of fresh oxygen carrier and the regeneration of reacted oxygen carrier were


Figure 6. SEM images for fresh and reacted oxygen carriers: (a) Fresh and (b) regenerated in a muffle oven.

determined by a Micromeritics NOVA1000e. Comparison to fresh oxygen carrier, the porous properties of reacted oxygen carrier particles had a significant change after 30 h tests. The BET surface area of fresh oxygen carrier and the regeneration of reacted oxygen carrier were 0.5465 and 0.1663 × 103 m2/ kg, respectively. The total pore volume of the oxygen carrier decreased from 1.536 to 0.831 × 10-4 m3/kg over the relatively long-term experiment. The remarkable decrease of the BET surface area and total pore volume of reacted oxygen carrier might be ascribed to the sintering surface of the particles. The shape and morphological features of fresh and reacted oxygen carriers were characterized by scanning electron microscopy (SEM, Sirion 200). The magnification of 3000 times was selected to analyze the surface micrographs of fresh oxygen carrier and the regeneration of reacted oxygen carrier, as shown in Figure 6. The surface of fresh oxygen carrier particles was loosely covered with grains of a size around 2-3 µm and had a relatively porous structure. The porous surface facilitated the diffusion of reactant gases into the core of oxygen carrier particles, enhancing the reactions between reactant gases and oxygen carrier particles. After the experiment of 30 h in the 10 kWth CLC reactor with biomass, significant physical change occurred on the surface of reacted oxygen carrier particles and the grains were completely converted into small ones (about 0.5 µm). The more uniform porous surface of reacted oxygen carriers was formed as a consequence of the accumulative effects


of long-term alternating reduction and oxidation as well as the attrition of oxygen carrier particles in the fast fluidized bed. For the process of chemical-looping combustion of biomass in the interconnected fluidized beds, the whole reduction of oxygen carrier with biomass in the fuel reactor was an endothermic process, while the oxidization of reduced oxygen carrier with air in the air reactor is intensively exothermic. Due to a relatively high oxygen concentration in the bottom zone of the fast fluidized bed, the oxidization of reduced oxygen carrier with air was very fierce; the heat released in the oxidization process led to the surface sintering of oxygen carrier particles and then to the limitation of reactant gas diffusion into the core of oxygen carrier particles. Therefore, to restrain the surface sintering of oxygen carrier particles, an intensive oxidization of reduced oxygen carrier with air in the air reactor should be avoided in the process of oxygen carrier regeneration, and air staging should be adopted for the oxidization of reduced oxygen carrier with air in the air reactor. An X-ray diffractometer (XRD, SHIMADZU) using Cu KR radiation (40 kV, 30 mA) was used to analyze the samples of fresh and reacted oxygen carriers. The samples were scanned in a step-scan mode with a step size of 0.02° over the angular 2θ range of 10-90°. Figure 7 shows the XRD spectra of fresh oxygen carrier, the regeneration of reacted oxygen carrier, the oxidized oxygen carrier of the air reactor, and the reduced oxygen carrier of the fuel reactor, respectively. Comparison to fresh oxygen carrier, there was no chemical changes occurred in the regeneration of reacted oxygen carrier and the oxidized oxygen carrier of the air reactor, and there were two phases, Fe2O3 and Fe3O4, in the three samples, as indicated in Figure 7a-c. However, there were no peaks found for Fe2O3 or FeO or Fe phase in the reduced oxygen carrier of the fuel reactor; as shown in Figure 7d, it indicated that the Fe2O3 phase of the oxygen carrier was completely reduced to Fe3O4 phase by biomass syngas in the fuel reactor. The contents of Fe2O3 and Fe3O4 in the samples can be characterized by the relative intensity of the major peak. The most intense reflection of Fe2O3 phase, IFe2O3, is located at 2θ ) 33.152°, while the most intense reflection of Fe3O4 phase, IFe3O4 is located at 2θ ) 35.422°. On the basis of the XRD spectra in Figure 7a-c, the IFe3O4/IFe2O3ratios of fresh oxygen carrier, the regeneration of reacted oxygen carrier, and the oxidized oxygen carrier of the air reactor are 0.22, 0.22, and 9.07, respectively. Due to the sintering surface of reacted oxygen carrier particles, the diffusion and penetration of reactant gases into the oxygen carrier particles was blocked up. The reduced oxygen carrier could not be sufficiently oxidized with air in the fast fluidized bed. It led to a relatively large IFe3O4/IFe2O3 ratio of the oxidized oxygen carrier of the air reactor. 3.5. Thermodynamic Discussion. Chemical reaction thermodynamics is important for the understanding of the reaction mechanism, the product composition, as well as the design of technical parameters in chemical-looping combustion of biomass. With the standard Gibbs free energy change, the equilibrium constants can be calculated for metal oxide reductions with CO in a wide range of operating temperature. The relationship between the equilibrium constant, Kp ) PCO2/PCO, and temperature T is expressed as R ln Kp ) -

∆HTθ + ∆STθ T

(7)

where ∆HθT and ∆SθT are the standard formation heat and the standard formation entropy at the corresponding reaction temperature, respectively, and R is the gas constant. With related

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Figure 7. XRD spectra of fresh and reacted oxygen carriers: (a) Fresh, (b) regenerated in a muffle oven, (c) oxidized in the air reactor, and (d) reduced in the fuel reactor.

Figure 8. Equilibrium constants for reduction of iron oxides by CO.

thermodynamic parameters,19 the equilibrium constants Kp can be calculated. Figure 8 shows the equilibrium constants (Kp ) PCO2/PCO) as a function of temperature for the reduction of iron oxides (Fe2O3, Fe3O4, and FeO) with CO. It shows that in the temperature range of 700-1100 °C the equilibrium constant Kp for the reduction of Fe2O3 - Fe3O4 varies from 105 to 104.5, Kp for Fe3O4 - FeO is about 2-5, while Kp for FeO - Fe is less than 1. It is apparent that Fe2O3 - Fe3O4 has a greater tendency to react with CO as compared to Fe3O4 - FeO and FeO - Fe in the temperature range of 700-1100 °C. To interpret the chemical stability of Fe2O3/Fe3O4/FeO/ Fe species in the fuel reactor, the stability diagram of solid phases Fe2O3/Fe3O4/FeO/Fe is indicated in Figure 9. It is calculated from the equilibrium constants of reactions Fe2O3-CO, Fe3O4-CO, and FeO-CO, respectively. Within the interesting temperature range of 700-1100 °C, the phase diagram clearly shows that Fe is stable only at strongly reducing conditions of a high CO concentration (>50%), (19) Ye, D. L.; Hu, J. H. Thermochemical Properties of Inorganic Substances; Metallurgy Industrial Press: Beijing, 2002.

Figure 9. Phase diagrams of iron species for reductions of iron oxides by CO.

whereas Fe3O4 is the stable species at a relatively low CO concentration and temperature. The intermediate range is FeO species at intermediate CO concentration. The range for FeO stability is increasing with temperature and CO concentration. Therefore, it is clear that it is not possible to reduce FeO or Fe further than Fe3O4 in the condition of the fuel reactor due to the thermodynamic limitation. As shown in Figure 2a, the CO concentration of the flue gas of the fuel reactor evidently increased with the fuel reactor temperature. An explanation for this behavior could be that biomass gasification with CO2 in the spout-fluid bed was significantly strengthened with a temperature increase and the CO production from biomass gasification was more temperature dependent than the CO oxidation with iron oxide in the spoutfluid bed. Thus, an increase in the fuel reactor temperature produced a higher increase for the CO production from biomass gasification than for the consumption of CO oxidation to CO2. It resulted in the increase of CO with an increase of the fuel reactor temperature. In the process of iron oxide reduction with biomass syngas the transformation of Fe2O3 to Fe3O4 is the


favored step in the spout-fluid bed. It was verified that there was only Fe3O4 phase in the particles of reduced oxygen carrier in the spout-fluid bed, as shown in Figure 7d. Although high up to 10% of CO existed in the flue gas of the fuel reactor at a fuel reactor temperature of 740 °C in the present system for chemical-looping combustion of biomass with iron oxide as an oxygen carrier, it is possible to reach a high conversion of biomass to CO2 in the fuel reactor by a supplement of a more reactive oxygen carrier of higher oxygen ratio, such as nickel-based oxygen carrier. The oxygen ratios for Fe2O3/ Fe3O4 and NiO/Ni oxygen carriers are 0.03 and 0.21, respectively. The higher oxygen ratio of nickel oxide can make the preparation of a mixed oxygen carrier with a lower NiO content. Therefore, the mixed oxygen carriers of iron oxide and NiO are favorable to achieve a high conversion of biomass to CO2. In future work the synergy of the mixed oxygen carriers of iron oxide and NiO will be carried out in the process for chemicallooping combustion of biomass. 4. Conclusions The performance of chemical-looping combustion of biomass was investigated in a 10 kWth CLC prototype with iron oxide as an oxygen carrier. A total 30 h of test was achieved with the same batch of iron oxide oxygen carrier, i.e., without adding any fresh iron oxide particles. The effect of the fuel reactor temperature on gas composition of the two reactors, the


proportion of biomass carbon reacting in the fuel reactor, and the conversion of biomass carbon to CO2 in the fuel reactor was experimentally investigated. The results showed that the CO production from biomass gasification with CO2 was more temperature dependent than the CO oxidation with iron oxide in the fuel reactor. An increase in the fuel reactor temperature produced a higher increase for the CO production from biomass gasification than for the consumption of CO oxidation to CO2. Although the conversion of biomass carbon to CO2 in the fuel reactor decreased with the increase of the fuel reactor temperature, there was a substantial increase in the proportion of biomass carbon reacting in the fuel reactor. The transformation of Fe2O3 to Fe3O4 is the favored step in the process of iron oxide reduction with biomass syngas. There was only Fe3O4 phase in the particles of reduced oxygen carrier in the fuel reactor. It is possible to reach a high conversion of biomass to CO2 in the fuel reactor by a supplement of a more reactive oxygen carrier of higher oxygen ratio, such as nickelbased oxygen carrier. Acknowledgment. This work was supported by the National Natural Science Foundation of China (90610016) and the HighTech Research and Development Program of China (2006AA05Z318). EF900033N

Chemical-Looping Combustion of Biomass in a 10 kWth Reactor with

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