Continuous Operation of a 10 kWth Chemical Looping Integrated

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Continuous Operation of a 10 kWth Chemical Looping Integrated Fluidized Bed Reactor for Gasifying Biomass Using an Iron-Based Oxygen Carrier Guoqiang Wei, Fang He,* Zhen Huang, Anqing Zheng, Kun Zhao, and Haibin Li CAS Key Laboratory of Renewable Energy, Guangzhou Institute of Energy Conversion, Chinese Academy of Sciences (CAS), Guangzhou 510640, China ABSTRACT: Chemical looping gasification (CLG) was investigated in a 10 kWth interconnected fluidized bed reactor with Fe2O3/Al2O3 as oxygen carriers (OC) and pine sawdust as fuel. The effects of the operation temperatures and sawdust feeding rate on the gas composition, cold gas efficiency, and carbon conversion rate of biomass were investigated. The fresh and used oxygen carrier particles were characterized by means of XRD, SEM, and BET. The results indicated that the sawdust was partially oxidized to syngas by lattice oxygen from the oxygen carrier. The syngas yield, cold gas efficiency, and carbon conversion increased with increasing operating temperature. Also, the concentrations of CO, H2, and CH4 in the syngas increased at the elevated temperature, while the CO2 fraction decreased. The feeding rate of biomass has a significant impact on the syngas composition and cold gas efficiency. There was an optimal value of feeding rate at 2.24 kg/h corresponding to the maximum cold gas efficiency in the tested reactor system. XRD analysis showed that the oxygen carrier particles were reduced to Fe3O4 from Fe2O3 in the course of the CLG reactions. BET results indicated the surface area, total pore volume, and average pore size of the oxygen carrier particles increased initially and then slightly decreased with the reaction proceeding, due to the interstice and thermal sintering. However, the OC samples were well regenerated and maintained a good crystalline state after 60 h of operation, which illustrated that the synthesized oxygen carrier had a stable reactivity and good resistance to agglomeration.

1. INTRODUCTION Currently, it is encouraging for the energy sector to apply alternative sources for reducing fossil fuel consumption due to the energy crisis and environmental protocols. Synthesis gas production from biomass gasification has been considered as a potential approach to replace fossil fuels, since biomass appears more promising for industrial application among the available sources of renewable energy in the world.1 The synthesis gas production from the partial oxidation of the biomass can be used for power and heat production but also upgraded to other valuable products, such as FT liquids, SNG, hydrogen, or chemicals.2,3 Although the biomass can be converted into syngas by various methods, such as steam gasification, oxygen gasification, air gasification, and so on,4−6 there are shortcomings in these methods that need to be improved. For example, steam gasification and oxygen gasification require water vapor or pure oxygen, the cost of which is expensive. And the syngas from air gasification of biomass was deemed to be of low heating value and high tar content.7−9 Meanwhile, the chemical looping gasification (CLG) of biomass can provide a solution to the issues due to solid fuels being partially oxidized into syngas by the lattice oxygen from the oxygen carriers. Accordingly, traditional gasifying agents such as steam and pure oxygen could be avoided in CLG. Chemical looping combustion (CLC) is a technology with inherent separation of the CO2, which was represented by Ishida et al.10,11 CLC uses oxygen carriers to transfer oxygen from combustion air to fuel, and the system consists of two separate reactors: an air reactor (AR) and a fuel reactor (FR).12 The CLG has the same basic principle as CLC, whereas the © 2014 American Chemical Society

main difference is that the target product of CLG is syngas (H2 and CO) instead of heat. Similar to CLC, the CLG system consists of two separate reactors: an air reactor (AR) and a fuel reactor (FR), as shown in Figure 1. Compared with the

Figure 1. Schematic of chemical looping gasification.

traditional biomass gasification methods, the CLG of biomass is an emerging technology, which can produce biomass-derived syngas by using lattice oxygen of the oxygen carrier instead of molecular oxygen. Since the gasifying agent was avoided in the CLG process, it was more effective to reduce the cost of the gasification. Besides, the tar content in the syngas could be decreased due to the catalytic effect of oxygen carriers.13 Also, the synthesis gas produced by the CLG of biomass theoretically Received: September 23, 2014 Revised: November 25, 2014 Published: December 2, 2014 233

dx.doi.org/10.1021/ef5021457 | Energy Fuels 2015, 29, 233−241

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Table 1. Proximate and Ultimate Analysis of Sawdusta Proximate analysis (wt/%,db) M 8.39 a

V 84.31

FC 6.88

LHV/ (KJ/kg,db)

Ultimate analysis (wt/%,db) A 0.42

C 46.44

H 6.21

O 47.29

N 0.05

S 0.01

18707

db: dry basis; FC: fixed carbon; LHV: lower heating value; M: moisture; V: volatile matter. was used to analyze the crystal structure of fresh and reacted samples. The samples was scanned at a rate of 2° min−1 between 2θ = 10°−90° with a step of 0.0167°. The surface morphology and characteristics of the samples were performed by scanning electron microscopy (SEM) on a Hitachi S4800 instruments. The BET surface was determined by N2 physisorption using a Micromeritics ASAP 2010 instruments. The samples were degassed under vacuum at 493 K for 6h before measurement. 2.3. Experimental Setup and Procedure. The experiments in the present work were performed in a 10 kWth interconnected circulating fluidized bed (shown in Figure 2). The prototype is

had a higher heating value for the reason that it can be not diluted with N2. In the past a few years, CLC with different oxygen carriers, such as the metal oxides of Ni, Fe, Co, Cu, and Mn, CaSO4, and perovskite, were investigated in either TG or a fixed bed or fluidized bed reactor using various fuels, such as methane, hydrogen, and syngas from coal gasification.14−24 Lyngfelt et al.25 proposed two interconnected fluidized bed designs to be used in CLC, and the results indicated that the process was feasible. Nicolas et al. tested a South African coal in a 10 kWth chemical looping combustor at the temperature 850 °C, finding that the coal conversion from the fuel reactor was in the range 78%−81%.16 Also, Johannes et al. studied the performance of two different Ni-based oxygen carriers in a 120 kW chemical looping pilot.26 The results indicated that high CH4 conversion and CO2 yield were achieved for both oxygen carriers. Tao Song et al. carried out the experimental investigation on hydrogen production from biomass gasification in interconnected fluidized beds.27−29 They found that there was an optimal value of steam/biomass ratio corresponding to maximal hydrogen yield. However, there are a limited number of studies focused on using CLG of biomass to produce syngas in the interconnected fluidized bed reactor in the open literature. Zhen Huang et al.30 conducted biomass direct chemical looping conversion with natural hematite as oxygen carrier in a single stage fluidized bed, which proved the feasibility of CLG. However, the biomass and oxygen carrier cannot realize continuous reaction. To further reveal the reactivity behavior and usefulness of iron oxide as an oxygen carrier in the CLG 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. In this work, a 10 kWth interconnected fluidized bed reactor has been built and the CLG of biomass was investigated in the reactor by using an iron-based oxygen carrier.

Figure 2. Schematic diagram of the interconnected circulating fluidized beds for CLG of biomass. composed of a fast fluidized bed as an air reactor, a bubbling bed as fuel reactor, two cyclones, two loopseals and feeder system. The air reactor was consisted of a bottom bubbling fluidized bed and an upper fast fluidized bed as a riser. The bottom bubbling fluidized bed is a circular column of 100 mm in diameter and 200 mm in height. The upper fast fluidized bed is a circular column with 50 mm in diameter and 1650 mm in height. A conical expansion section with a height of 120 mm was used to connect the two fluidized beds. The fuel reactor has a circular column shape with 150 mm in diameter and 500 mm in height. There are porous conical stainless steel plates as gas distributors with 1% opening ratio at the bottom of both air reactor and fuel reactor. The hole diameter of gas distributors was 1 mm. The taper angles for gas distributors of air reactor and fuel reactor were 47° and 115°. The top of the air reactor gas distributor was fixed at the bottom flange of air reactor with a diameter of 100 mm. The bottom of the air reactor gas distributor was connected to outlet pipe with a diameter of 14 mm. The fuel reactor gas distributor was installed by using a similar connection method with a top diameter of 150 mm and bottom diameter of 14 mm. The two reactors are connected by a rectangular loopseal with a cross section of 40 × 85 mm2 and a height of 280 mm. The outlets of the fast fluidized bed and bubbling bed are linked to the top of the fuel reactor by cyclones and another loopseal. The feeder system was fixed at the right bottom of fuel reactor, which was composed of hopper, screw feeder and propulsive motor. The hoper was including a circular column of 300 mm in inner diameter, a cone shaped bottom of 250

2. EXPERIMENTAL SECTION 2.1. Materials. Sawdust of pine collected from Guangdong province (china), was used as fuel in the tests. The sample was crushed and sieved in to particles with a size range of 0.3−0.45 mm and dried for 8h at 105 °C before experiment. Proximate and ultimate analyses are shown in Table 1. The Fe2O3/Al2O3 oxygen carrier with a mass ratio of Fe2O3/Al2O3 = 7/3 was prepared by mechanical method and calcination. Starting materials Fe2O3 (analytically pure), Al2O3 (analytically pure) were mixed according to the mass ratio 7:3, blended in a Planetary Mixer sufficiently and pressed into a 1 mm diameter cylindrical bar by the screw extrusion machine. The resulting extrudates were dried at 105 °C for 24 h and then calcined at 1100 °C for 6h in the muffle furnace. After grinding and screening, the oxygen carriers with particle size in 60−80 mesh were obtained. The Fe2O3/Al2O3 mass ratio of 7/3 was chosen according to our previous experiments. Enough mechanical strength and specific surface area of oxygen carrier can not be achieved at higher ratio of Fe2O3/ Al2O3. Conversely, the reactivity and oxygen transfer capacity of oxygen carrier will decrease at lower ratio of Fe2O3/Al2O3. 2.2. Characterization of Oxygen Carrier. Powder X-ray diffraction (XRD, X’Pert PRO MPD) using Cu Kα (40 kV, 40 mA) 234


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2.4. Data Evaluation. The gas relative concentration (ci) of each run in the fuel reactor was calculated as x ci = i × 100% ∑ xi (1)

mm in height, a balance gas inlet and mixer motor. The screw feeder was directly under the hopper and the design parameters were 38 mm in screw flight diameter (the outer diameter of the helical screw thread), 25 mm in screw core (the diameter of the central screw shaft), 18 mm in pitch (distance from one thread peak to the next) and 350 mm in length of hopper trough. In order to adjust the feeding rate of biomass, the screw feeder was driven by a variable frequency motor. The feeding rate of biomass can be increased linearly with the increasing frequency of motor. In the air reactor, oxygen carrier particles are entrained to the top of the reactor by the air stream, and transported to the fuel reactor through a cyclone. In the fuel reactor, the oxygen carrier reacts with the fuel, and the reduced oxygen carrier particles are transferred back to the air reactor via a loopseal. This achieves the external circulation of oxygen carrier particles in the interconnected circulating fluidized bed. Similarly, oxygen carrier particles carried in the synthesis gas are separated by a cyclone and then directed back into the fuel reactor, which constitutes the internal circulation in the process of apparatus operation. By the way, the loopseal allows particles unidirectional movement and prevents the contamination of the flue gas between the two reactors. The two reactors are electrically heated in an oven which supplies heat for start-up and compensates heat loss during the operation. Thermocouples and differential pressure transducers were located at different points of the prototype (shown in Figure 2 and Table 2) to

where xi denoted the volume fraction of species i and i indicated the compositions of the flue gas in the fuel reactor. Similarly, the relative concentrations of CO2 and O2 in the flue gas of the air reactor were calculated according to eq 1, which was obtained based on the measured data. Carbon conversion efficiency (ηconversion) in the system was defined as the proportion of the carbon converted into gaseous products from the total carbon in the sawdust fed into the fuel reactor, and it was calculated as in eq 2:

P1

P2

P3

P4

T1

T2

T3

Height (mm)

100

1850

100

400

600

1300

250

22.4(298/273)Mc %

× 100%

(2)

−1

where Gv (m ·kg ) and Mc (%) were the gas yield of the synthesis production and carbon fraction in the biomass, respectively, and VCO, VCO2, and VCH4 (%) represented the volume fractions of CO, CO2, and CH4 in the flue gas of the fuel reactor. The lower heating value (LHV, kJ/Nm3) of the gas products is calculated as in the below equation:27 3

LHV = 126VCO + 108VH2 + 359VCH4 + 635VC2Hm

Table 2. Distance of Measuring Points from the Air Distributors in Both Air Reactor and Fuel Reactor NO.

12(VCO2 + VCO + VCH4)GV

ηconversion =

(3)

where VCO, VH2, VCH4, and VC2Hm were the volume fractions of CO, H2, CH4, and C2Hm in the flue gas, respectively. Cold gas efficiency (η) is defined as the ratio of the heating value of the gas products from unit mass biomass gasification to the total heating value of the unit mass of biomass. It is calculated as

η= display the operating conditions and monitor the cycling stability of the fluidized bed in real time. The outlet gases from the air reactor and fuel reactor were induced with suction pump to an ice−water cooler where the steam was condensed and removed. The produced gases were collected with gas bags and analyzed by an offline gas chromatograph. In the case of the tests of temperature effect, the bed mass added into the prototype system was 5.8 kg. The tested reactors were heated from room temperature to 900 °C under N2 atmosphere. And then, the oxygen carriers were adjusted to circulating balanced state. Correspondingly, the N2 gas velocity in air reactor, fuel reactor and loopseal were 1.7 m/s, 0.58 and 0.3 m/s, respectively. The solid circulation rate of circulating balanced state of oxygen carriers was 0.017 kg/s, which was estimated based on the bed pressure drop and the velocity of particles from previous cold experiment data. After that, biomass sample was continuously fed into fuel reactor from the hopper by screw feeder, where 0.01m/s of N2 was used as balance gas. The feeding rate of biomass was keeping a constant feeding rate of 1.2 kg/ h. Outlet gas was collected from the tested reactor at temperature of 670, 700, 750, 800, 850, and 900 °C, respectively. Data collected from each tested temperature was repeated 3 times with an interval of 5 min for average. Similarly, the biomass feeding rate experiments were performed at two reactors temperature of 850 °C under N2 atmosphere and circulating balanced state. Outlet gas was collected from fuel reactor and air reactor at the feeding rate of 0.72, 1.52, 2.24, 2.93, and 3.66 kg/h (corresponding to 5HZ changes of screw feeder motor) and every data was repeated 3 times with an interval of 5 min for average. The gas yield was calculated according to the N2 balance between the inlet and outlet gas of the prototype system. The concentration of produced gas was calculated on the basis of nitrogen free. The fractions of the four components (H2, CH4, CO2, CO) were analyzed by gas chromatograph. While, the amount of other products (e.g., light hydrocarbons) was relatively low enough to be neglected. The oxygen carrier from fuel reactor and air reactor was cooled in N2 and air atmosphere, respectively and then collected for analysis.

LHV × GP × 100% Qb

(4)

3

where GP (Nm /kg) was the gas yield of the synthesis products under standard state and LHV (kJ/Nm3) and Qb (kJ/kg) were the lower heating value and total heating value of biomass at room temperature, respectively.

3. RESULTS AND DISCUSSION 3.1. Typical Bed Pressure. A test of 60 h on CLG of sawdust with the synthetic oxygen carrier was investigated in the apparatus shown in Figure 2. The typical bed pressure as a function of reaction time in both the air reactor (P1, P2) and fuel reactor (P3, P4) was displayed in Figure 3. It is observed

Figure 3. Typical bed pressure as a function of reaction time in both fuel reactor and air reactor. 235


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that the pressure signal fluctuated around 1.09, 0.45, 0.34, and 0.23 kPa, respectively, which indicated that the system was in a stable operation status and the oxygen carrier circulated in the experiment between the fast fluidized bed reactor and the bubbling fluidized reactor.31 3.2. Analysis of Temperature. 3.2.1. Gas Composition of Both Fuel Reactor and Air Reactor. The temperature of the fuel reactor is crucial for CLG of biomass in the interconnected fluidized bed. In the present work, the effects of reaction temperature on gas composition of both fuel reactor and air reactor were investigated, keeping the biomass feeding rate at 1.2 kg/h. The operation temperature was obtained from the thermal analysis of CLG of biomass and previous experiments. The CLG reactions cannot occur at temperatures below 650 °C, and the sintering of oxygen carrier was prone to occur when the temperature exceeded 1000 °C. The concentrations of outlet gas of the two reactors after water condensation as a function of temperature are shown in Figures 4 and 5, respectively.

gas−solid reactions. It is known that the solid−solid reaction is much slower than the gas−solid reaction. Therefore, the pyrolysis intermediates of biomass should be the major reductive agents for the reduction of oxygen carriers in a fuel reactor. The major reactions for CLG of biomass under atmospheric pressure are as follows:32 Biomass (pyrolysis) → synthesis gas (H 2 , CO, CH4 , Cx Hy , etc. ) + char + tar (5)

ΔH > 0 C + CO2 = 2CO

ΔH1073K = 40.603 kcal·mol−1 (6)

C + H 2O(g) = CO + H 2 ΔH1073K = 32.449 kcal·mol−1

(7)

CO + H 2O(g) = CO2 + H 2 ΔH1073K = −8.154 kcal·mol−1

(8)

CH4 + H 2O(g) = CO + 3H 2 ΔH1073K = 53.828 kcal·mol−1

(9)

CO + 3Fe2O3 = 2Fe3O4 + CO2 ΔH1073K = −9.768 kcal·mol−1

(10)

H 2 + 3Fe2O3 = 2Fe3O4 + H 2O ΔH1073K = −1.614 kcal·mol−1

(11)

CH4 + 9Fe2O3 = 6Fe3O4 + CO + 2H 2O ΔH1073K = 48.986 kcal·mol−1

(12)

H 2 + Fe2O3 = 2FeO + H 2O

Figure 4. Effect of temperature on gas composition of fuel reactor.

ΔH1073K = 11.163 kcal·mol−1

(13)

CH4 + 12Fe2O3 = 8Fe3O4 + CO2 + 2H 2O ΔH1073K = 39.219 kcal·mol−1

(14)

CO + Fe2O3 = 2FeO + CO2 ΔH1073K = −1.25 kcal·mol−1 FexOy

Tar ⎯⎯⎯⎯→ CO + H 2 + Cx Hy

(15)

ΔH > 0

(16)

As shown in Figure 4, the concentrations of CO and H2 increased slowly with increasing temperature at the lower temperature of the fuel reactor and then ascended rapidly at the temperature range of 800 to 900 °C, while the content of CH4 remained stable at first and then decreased with the increasing of temperature. These changes can be explained by the following two aspects. On the one hand, high temperature benefited the products in endothermic reactions and restrained the exothermic reactions as well. On the other hand, the oxygen carrier can act as a catalyst strengthening hydrocarbon cracking at high temperature.13,33 Taking into account the thermodynamics of CLG for biomass, the principle generation reactions of CO involved

Figure 5. Effect of reaction temperature on gas composition of air reactor.

The direct reduction of iron oxygen carrier with biomass was a solid−solid reaction,while the reactions of iron-based oxygen carrier with biomass pyrolysis gas at high temperature were 236


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interconnected fluidized beds, as indicated in Table 3. Both carbon conversion and cold gas efficiency increased with the

reactions 5−7, 9, 12, and 16, which were intensive endothermic processes, whereas the consumption of CO with reactions 8, 10, and 15 was an exothermic process. Consequently, more CO was prone to generate with the temperature increasing. The syngas contained an amount of CH4 as a consequence of pyrolysis of biomass determined by reaction 5, which was beneficial at the elevated temperature, whereas the CH4 oxidization reactions 12 and 14 and the steaming reform reaction 9 were all intensive endothermic reactions, which can offset the effect of reaction 5 at an early stage and caused the fraction of CH4 to decrease at high temperature. As for the H2, it was mainly generated by endothermic reactions 5, 7, 9, and 16 and consumed by the exothermic process (eq 11). In spite of the fact that the H2 consumption reaction, eq 13, was an endothermic reaction, which can be enhanced by increasing temperature, it was not the principal reaction in the CLG of biomass verified by the XRD of products, so the concentration of H2 increased at the elevated temperature. In the case of CO2, although the water gas shift reaction 8 and oxidization reaction 15 were attributed to the generation of CO2, the fraction of CO2 still decreased at high temperature for the reason that the consumption rate of CO2 by the Boudouard reaction 6 was much faster due to the lattice supply being restricted by the oxygen carrier. Also, due to the catalysis of oxygen carriers, the pyrolysis liquid products of biomass (e.g., tar) can be catalytic into small molecular gas (e.g., H2, CO) according to reaction 16. The catalyst activity of oxygen carrier depended on the availability of reducible iron on the surface particles. The reducible iron could lead to the formation of the active metallic iron to mediate C− C and C−H bond cleavage.34 As a result, it might have a positive effect on the increasing of CO and H2. The effect of the reaction temperature on the exit gas composition of the air reactor was also investigated, as indicated in Figure 5. In the air reactor, oxygen carriers recovered the lattice oxygen to its original state in combustion air and the major reactions can be presented as below: Fe3O4 + O2 → Fe2O3

(17)

FeO + O2 → Fe2O3

(18)

C + O2 → CO2

(19)

Table 3. Effect of Temperature on the Synthesis Gas from CLG of Sawdust Temp (°C)

Gas yield (Nm3/kg)

LHV (MJ/Nm3)

Carbon conversion (%)

Cold gas efficiency (%)

670 700 750 800 850 900

0.95 1.07 1.09 1.11 1.13 1.17

9.26 9.91 10.60 10.97 12.13 12.19

83.94 90.72 92.92 93.16 93.33 93.76

41.56 48.05 52.67 54.61 60.53 61.11

elevation of temperature. As stated above, high temperature benefited the thermal cracking of biomass tar and steam reforming,35 which were endothermic reactions. Therefore, more carbon and hydrogen can be converted to syngas though reactions 5−16 at elevated temperature. Accordingly, it caused the gas yield, LHV, carbon conversion, and cold gas efficiency to increase respectively from 0.95 to 1.17 N m3/kg, 9.26 to 12.19 MJ/Nm3, 83.94% to 93.76%, and 41.56% to 61.11% with the rising temperature, as shown in Table 3. 3.3. Effect of the Biomass Feeding Rate. 3.3.1. Gas Composition of Both Fuel Reactor and Air Reactor. The influence of the ratio of oxygen carrier to biomass introduced into the fuel reactor on the exit gas composition from both the air reactor and fuel reactor was also studied. Changing the circulation of the solid oxygen carrier particles may result in abnormal fluidization of the whole interconnected fluidizing beds system. Therefore, we investigated the effect of the ratio of oxygen carrier to biomass in the course of the CLG by varying the biomass feeding rate from 0.72 kg/h to 3.66 kg/h and keeping the circulation rate of the oxygen carriers as a constant (0.017 kg/s). The value of the biomass feeding rate was obtained from the designed parameters of the prototype and the variation amount of the screw feeder. The influences of the variation of biomass feeding rate on the gas composition of the fuel reactor and air reactor at the temperature 850 °C are shown in Figures 6 and 7, respectively. As seen from Figure 6, the CO, H2, and CH4 concentrations increased smoothly with the increasing of biomass feeding rate,

The oxygen fraction in the flue gas of the air reactor was maintained in the range of 9%−11% and increased at elevated reaction temperatures, whereas the CO2 concentration showed an opposite trend, decreasing with increasing temperatures. The low fraction of O2 can be attributed to consuming reactions 17 −19, and the CO2 in the air reactor was mainly produced from the residual char burning with air. The residual char was entrained to the air reactor by the external circulation of oxygen carrier from the fuel reactor. As mentioned above, the biomass CLG reactions 5−16 were principally endothermic processes, and most reactions were remarkably strengthened with the rising fuel reactor temperature; as a result, more biomass was converted into gaseous product, which lead to a decrease of residual char coming into the air reactor. Consequently, the content of O2 ascended slightly while CO2 descended with the increasing of reaction temperature. 3.2.2. Carbon Conversion and Cold Gas Efficiency. Two parameters of carbon conversion, ηc and cold gas efficiency η in the system, can be used to investigate the effect of the reaction temperatures on the process of CLG of biomass in the

Figure 6. Effect of feeding rates on gas composition of fuel reactor. 237


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Figure 7. Effect of feeding rates on gas composition of air reactor.

Figure 8. Effect of feeding rate on the carbon conversion and cold gas efficiency.

while CO2 declined sharply. This can be explained by reactions 5−16. When the feeding rate of sawdust was too small, while the oxygen carrier circulation was kept a constant, reactions 10, 11, and 14 of the total oxidation of biomass occurred to generate CO2 and H2O due to the lattice oxygen provision excessively. With the increasing of sawdust feeding rate, the sawdust pyrolysis reactions 5−7 and 9 were reinforced gradually, whereas the reactions between biomass pyrolysis intermediate and iron oxides (eqs 10−15) cannot be reinforced greatly due to the constant available providing of lattice oxygen, leading to the rise of CO, H2, and CH4.36 Simultaneously, the Boudouard reaction 6 was promoted by the excessive residual carbon to generate CO as well as consumed CO2, resulting in the decreasing of CO2. Finally, the pyrolysis reaction of biomass (eq 5) became overwhelming as the biomass feeding rate increased. The effect of the biomass feeding rate on the gas composition of the air reactor at the temperature 850 °C was indicated in Figure 7. The oxygen composition in the flue gas of the air reactor was varied from 13%−7.6%, and it decreased with the increasing of the biomass feeding rate, while the CO2 fraction increased rapidly. As mentioned above, the CO2 was mainly generated though reaction 19 from residual chars burnt with air in the air reactor. The residual char was entrained to the air reactor by the external circulation of oxygen carrier from the fuel reactor. With the increase of biomass feeding rate, more residual char was entrained into the air reactor from the fuel reactor. Consequently, the concentration of O2 decreased while CO2 increased with the elevated biomass feeding rate. 3.3.2. Carbon Conversion and Cold Gas Efficiency. Carbon conversion ηc and cold gas efficiency η in the system were investigated to evaluate the effect of the variations of biomass feeding rate on the CLG of biomass, as indicated in Figure 8. It was observed that the carbon conversion maintained a steady descent while the cold gas efficiency increased slightly and then decreased with the increasing of feeding rate. There was an optimal value of biomass feeding rate at 2.24 kg/h corresponding to the maximal cold gas efficiency. Lower biomass feeding rate gave a higher carbon conversion because the lattice oxygen was sufficient to mainly oxidize the biomass into CO2 and H2O. However, with the further increasing of biomass feeding rate, pyrolysis of biomass occurred to mainly generate intermediates. Due to the deficient lattice oxygen provision, part of carbon and hydrogen cannot be

converted into syngas. Consequently, there was more and more residual char in the fuel reactor and the gas yield as well as carbon conversion continuously decreased from 1.12 to 0.86 N m3/kg and 97.73% to 69.45%, respectively, as shown in Table 4 Table 4. Effect of Feeding Rate on the Synthesis Gas from CLG of Biomass Feeding rate (kg/h)

Gas yield (Nm3kg−1)

LHV (MJ/Nm3)

Carbon conversion (%)

Cold gas efficiency (%)

0.72 1.52 2.24 2.93 3.66

1.12 1.11 1.02 0.98 0.86

7.14 10.74 12.84 12.99 13.46

97.73 92.85 84.42 80.19 69.45

42.74 63.74 70.01 68.06 61.86

and Figure 8. However, the gas LHV and cold gas efficiency presented a different profile, shown in Table 4. The gas LHV increased from 7.14 to 13.46 MJ/Nm3 as a consequence of gas concentration variations, especially the increase of CO, H2, and CH4 and the decrease of CO2. As for the cold gas efficiency, it depends on the CO, H2, and CH4 composition of the syngas and the gas yield. With more biomass being pyrolyzed, the syngas LHV increased while the syngas yield decreased. These two aspects together caused the cold gas efficiency to slightly ascend and then showed a downtrend. The cold gas efficiency achieved maximum of 70.01% at the feeding rate 2.24 kg/h. The profile of the cold gas efficiency indicated that the ascending of syngas LHV cannot compensate the descending of the gas yield, leading to the rapid decrease of cold gas efficiency with the further increasing feeding rate. 3.4. Effect of Reaction Time. To evaluate the reactivity of the oxygen carrier and the stability of the system, the typical carbon conversion and cold gas efficiency were measured at the temperature of 850 °C and biomass feeding rate of 1.2 kg/h for a 60 h test term, shown in Figure 9. It is observed that the carbon conversion and cold gas efficiency of the interconnected fluidized respectively decreased from 94.38% to 89.76% and 60.82% to 58.25% after a 60 h reaction, which shows that there was a slight loss in reactivity of oxygen carriers, but the particles still can be used to transmit lattice oxygen. However, the oxygen carrier can still be used in the process of CLG of 238


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Figure 10. XRD patterns of fresh and used oxygen carrier. Figure 9. Carbon conversion and cold gas efficiency as a function of reaction time and number of cycles.

Meanwhile, it is observed that the XRD patterns of the oxidized oxygen carriers are in good agreement with the patterns of fresh samples. The result indicates that the physical crystal form of the oxygen carrier did not change during the reaction process. The oxygen carriers were well regenerated and maintained a good crystalline state after 60 h of reaction. 3.6. SEM Micrographs. The shape and morphological features of fresh and reacted oxygen carriers were characterized by scanning electron microscopy (SEM), as shown in Figure 11. The morphology of the fresh Fe2O3/Al2O3 particles exhibits an irregularly blocky structure with an average size of around 3−6 μm. Particles of Al2O3 were dispersed on that of Fe2O3. Small interconnected pores exit the surface of particles, which are beneficial to diffusion and penetration of reactant gas into solid particles. After 20 h reactions in the interconnected fluidized beds, significant physical change occurred on the oxygen carries, and the particles had a relatively regular blocky structure. However, the particle converted to small grains with the reaction proceeding at the experiment time of 60 h. The more porous surface of a reduced oxygen carrier was formed as a consequence of accumulative effects, which included long time altering reduction and oxidation as well as attrition in the fuel reactor and air reactor. Meanwhile, it was observed that there was agglomeration on the surface of oxidized oxygen carrier particles after 60 h reactions, which can be ascribed to the slight sintering of the samples. Due to a relatively high oxygen concentration in the bottom of the air reactor, the reduced oxygen carriers were oxidized by air very fiercely with an intensively exothermic process. The surface of oxygen carrier grains became rougher and agglomerative, which produced a limitation on reactant gas diffusion into the core of oxygen carrier particles.38 3.7. Specific Surface Area Analysis. The surface area, total pore volume as well as average pore size of the fresh and used oxygen carriers are illustrated in Table 6. The porous properties had a significant change after 60 h tests. The surface area increased from 1.727 m2/g to 2.703 m2/g as the operation time increased from 20 h to 40 h. Additionally, the total pore volume and average pore size of the reacted oxygen carrier all increased; however, with the reaction time increasing, the surface area of the oxygen carrier decreased from 2.703 to 2.582 m2/g after 60 h reactions; similarly, the total pore volume and

biomass on account of a relatively higher carbon conversion. Additionally, the carbon balance analysis of two reactors was performed, as shown in Table 5. The input carbon was Table 5. Carbon Balance Analysis of Air Reactor and Fuel Reactor Output carbon mass (mol/h) AR

FR

Temp (°C)

Input carbon mass (mol/h)

Gas

Gas

Fine carbon

Recovery (%)

850

44.6

2.66

41.18

0.19

98.72

calculated according to the feeding rate of biomass. The output carbon mass was estimated by the carbonaceous gas concentration from the air reactor and fuel reactor. Also, the fine carbon was collected from the cyclone of the fuel reactor where the residue char was separated from the oxygen carrier. The mass of liquid products was relatively low enough to be neglected. The recovery was defined as the ratio of output carbon mass and input carbon mass in the process of the CLG reaction. It is observed that the recovery of both the carbonaceous gas and fine carbon is 98.72%, which suggests that the test process is reliable. Also, it is found that 2.66 mol/h carbon which contains 5.96% of total input carbon is consumed in the air reactor and that the fine char escaping from the cyclones is about 0.43%. 3.5. XRD Patterns of Oxygen Carriers. Figure 10 shows the XRD spectra of the oxygen carrier in its fresh, reduced, and reoxidized states, respectively. The reduced samples and oxidized samples were collected from the fuel reactor and air reactor separately after 60 h of operation. There were three crystalline phases of Fe2O3, Fe3O4, and Al2O3 in the three samples (JCPS card 01-089-8104, 01-088-0315, 010-070-3322), as indicated in Figure 10. FeO or Fe phase was not found in the reduced samples of the fuel reactor, which manifested that Fe2O3 was mostly reduced to Fe3O4. This result was in agreement with the literature.31,37 The main reduction reactions of iron oxide particles with biomass pyrolysis products in the system are listed in reactions 10−16. 239


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Figure 11. SEM analysis of fresh and used oxygen carriers.

rate and cold gas efficiency all increased, while the CO2 concentration decreased due to the Boudouard reaction. The feeding rate of biomass had a significant effect on the syngas composition and cold gas efficiency. There was an optimal value of biomass feeding rate at 2.24 kg/h corresponding to the maximum cold gas efficiency. The results of XRD, SEM, and BET showed that oxygen carrier was reduced to Fe3O4 from Fe2O3 in the process of iron oxide reaction with biomass. Though thermal sintering occurred in partial oxygen carrier particles, the oxygen carrier particles were well regenerated and maintained a good crystalline state after 60 h reactions, which indicated that the synthesized oxygen carrier had a stable reactivity and relative resistance to agglomeration.

Table 6. Specific Surface Area of the Synthesized Oxygen Carriers Fe-based oxygen carrier (OC)

BET-surface area (m2/g)

Total pore volume (cc/g)

Avg pore size (nm)

Fesh Used 20 h, oxidized Used 40 h, oxidized Used 60 h, oxidized

1.727 2.022 2.703 2.582

0.0038 0.0064 0.0170 0.0165

8.73 12.02 25.22 25.13

average pore size showed the same changed trend decreasing from 0.017 to 0.0165 cc/g and 25.22 to 25.13 nm, respectively. The increase of the surface area, pore volume, and average pore size may be due to the crack and interstice of particles, which can be obviously confirmed by the SEM images of Figure 11. The decrease of surface area might be ascribed to the agglomerated grains after sintering in the cyclic reactions. This result is in agreement with the literature.39−41 Besides, though thermal sintering occurred in the oxygen carrier particles after 60 h cyclic reactions, the reactivity of the oxygen carrier did not significantly decline, which can be verified by the XRD patterns in Figure 10 and the carbon conversion as well as cold gas efficiency in Figure 9.

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The financial support of the National Natural Science Foundation of China (51076154) is gratefully acknowledged. This work was also supported by the National Key Technology R&D Program of the 12th Five-Year Plan of China (2011BAD15B05).

4. CONCLUSIONS The CLG of biomass was performed in a 10 KW th interconnected fluidized bed reactor by using an iron-based oxygen carrier. The effects of temperature and biomass feeding rate on the syngas composition, carbon conversion rate, and cold gas efficiency were discussed. The results indicated that appropriate composition syngas could be obtained from the reactors, and with the increasing of temperature, the fractions of CO, H2, and CH4 in the syngas as well as carbon conversion

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