Design and Fluid Dynamic Analysis of a Three-Fluidized-Bed Reactor

Article pubs.acs.org/IECR

Design and Fluid Dynamic Analysis of a Three-Fluidized-Bed Reactor System for Chemical-Looping Hydrogen Generation Zhipeng Xue, Shiyi Chen, Dong Wang, and Wenguo Xiang* Key Laboratory of Energy Thermal Conversion and Control, Ministry of Education, School of Energy and Environment, Southeast University, Nanjing 210096, China ABSTRACT: Chemical-looping hydrogen generation (CLHG) can produce hydrogen from fossils fuels with inherent separation of CO2. Iron oxide is a suitable oxygen carrier for this process. The CLHG process basically involves three reactors, a fuel reactor (FR), a steam reactor (SR), and an air reactor (AR). In the FR, the carbon-containing fuel gases react with hematite (Fe2O3). The product solids are wüstite (FeO), and the product stream is a mixture of carbon dioxide and water vapor. After water condensation, pure carbon dioxide can be obtained. FeO then enters the SR and react with steam, giving the gas product hydrogen and the solid product magnetite (Fe3O4). In the AR, Fe3O4 is reoxidized to Fe2O3. Through this cycle, hydrogen is generated with inherent separation of CO2. In this article, a novel compact fluidized-bed fuel reactor is proposed. It integrates a bubbling fluidized bed and a riser to obtain full conversion of unreacted fuel gases through the thermodynamic equilibrium limit. Based on this fuel reactor, a cold-flow model of the three-fluidized-bed reactor system with a 50-kW CLHG design scheme was built to test the feasibility of this CLHG process. A series of tests with respect to solids circulation rate, gas leakage, and stability of long-term operation were performed by varying the inlet gas flow and total solids inventory. The results showed that the threefluidized-bed reactor system can run steadily. The solids circulation rate could be changed in a wide range by adjusting the inlet gas flows. The gas leakage was associated with both the solids circulation rate and the pressure difference balanced by the downcomer. The system showed a stable pressure difference and solids circulation rate during a test of long-term operation.

1. INTRODUCTION Hydrogen is expected to be a clean energy carrier in the foreseeable future. It can be produced from many primary energy fuels, such as coal, oil, natural gas, and renewable sources. To use hydrogen in a clean manner, the hydrogen production process should also be environmentally benign.1 Considering production costs, fossil fuels will still be the dominant sources for hydrogen production for a long time. However, current technologies of hydrogen production from fossil fuels with CO2 capture have significant energy penalties that increase the production costs. Therefore, the major, perhaps dominating, criterion is to develop innovative technologies that can substantially reduce the cost of CO2 capture while simultaneously producing H2 from fossil fuels.2 Recently, technology based on chemical-looping combustion (CLC) and the steam−iron process to generate hydrogen with inherent separation of CO2 has received increasing attention. CLC is a novel combustion technology with inherent separation of CO2. The standard CLC system consists of two reactors: an air reactor and a fuel reactor. The essential oxygen for fuel oxidation is supplied by an oxygen carrier that transports the oxygen between the two reactors. In this way, the fuel and the air are never mixed, and the flue gas from fuel reactor contains only CO2 and H2O if full oxidation of the fuel gas is achieved, which means that pure CO2 can be obtained by cooling the gas and removing the water.3 The chemical-looping concept can also be used to produce hydrogen. The steam−iron process was one of the earliest methods for producing hydrogen, using iron oxides as the oxygen carrier. It was first proposed in 1910 by Messerschmitt.4 The advantage of the steam−iron process is that it can produce almost pure hydrogen with simple purification. However, the low conversion rate of the reducing gas and poor recyclability of the © 2012 American Chemical Society

iron-based carrier limited its development and application with the emergence of steam−methane reforming technology.5−9 Chemical-looping hydrogen generation (CLHG) is a novel technology that integrates the steam−iron process with CLC to produce H2 and simultaneously capture CO2. The CLHG system consists of three reactors: a fuel reactor (FR), a steam reactor (SR), and an air reactor (AR). The iron-based oxygen carrier circulates through these three reactors, as shown in Figure 1. In the FR, fuel gases (taking syngas as an example) react with hematite (Fe2O3) according to the following reactions Fe2O3 + CO → 2FeO + CO2

ΔH298 = 5.50kJ/mol (1)

Fe2O3 + H2 → 2FeO + H2O ΔH298 = 46.63kJ/mol

(2)

Ideally, if the fuel gas can be fully converted, the exhaust gas from the FR contains only CO2 and H2O. After water condensation, almost-pure CO2 can be obtained. The reduced FeO is transferred to the SR, where it reacts according to the equation 2 2 2 2FeO + H2O → Fe3O4 + H2 3 3 3 ΔH298 = ‐49.83kJ/mol Received: Revised: Accepted: Published: 4267

(3)

May 17, 2011 December 22, 2011 February 28, 2012 February 28, 2012 dx.doi.org/10.1021/ie201052r | Ind. Eng. Chem. Res. 2012, 51, 4267−4278

Industrial & Engineering Chemistry Research

Article

successful coupling of a chemical-looping system with a gasifier. Gnanapragasam et al.13 proposed coal direct chemical looping (CDCL) substituting the gasification process in syngas chemical looping (SCL), thus eliminating the need for higher oxygen consumption. The process simulation results showed that CDCL has a higher hydrogen-to-CO2 ratio than SCL, along with overall advantages based on resource requirements and specific hydrogen produced. Many studies have focused on process simulation to confirm the feasibility of CLHG; however, limited experimental research has been carried out. Chen et al.14 studied CLHG using iron oxides as the oxygen carrier and CO as the fuel in a batch fluidized bed. The effects of reduction time, operating temperature, particle size, and addition of CO2 in the reduction gases on hydrogen production were investigated. Yang et al.15 verified the feasibility of chemical-looping hydrogen generation with direct reduction of iron oxides by coal char in a fluidized bed. Three processes, including Fe2O3 reduction by char in a direct path (without a gasifying agent) based on CLC, H2 production by the steam−iron process, and oxidation of Fe3O4 to the initial Fe2O3 by air, can be performed. However, because of thermodynamic limitations, CO cannot be fully oxidized to CO2 in the reduction process, which might lose the advantage of CLHG in terms of CO2 capture. This problem can be resolved by countercurrent operation in a moving bed, as stated by Gupta et al.16 and a bench-scale moving-bed reducer based on this concept was designed by Li et al.17 A syngas conversion in excess of 99.5% and an oxygen carrier conversion of nearly 50% were obtained in this reducer. However, a slight increase in the flow rate of syngas led to incomplete conversion of the syngas. Therefore, finding a suitable fuel reactor to solve this problem is an important research topic for the development of CLHG. In this article, we present a novel fuel reactor concept for CLHG that can obtain sufficient conversion of the fuel gas. Based on this fuel reactor concept, a three-fluidized-bed reactor system representing a 50-kW prototype for CLHG was designed, and a cold-flow model of it was manufactured. We report herein the results of an extensive testing program to assess the feasibility and understand the operating rules and hydrodynamic characteristics of the three-fluidized-bed reactor system based on the cold-flow model.

Figure 1. Conceptual scheme of chemical-looping hydrogen generation.

In the SR, hydrogen is generated by oxidizing FeO. The gaseous outlet stream is a mixture of hydrogen and H2O, and almost-pure hydrogen can be received after the water vapor is cooled. However, the iron oxides can be oxidized only to magnetite (Fe3O4); thus, further oxidation to more stable iron oxides occurs in the AR. The reaction in the AR is 2 1 Fe3O4 + O2 → Fe2O3 3 6

ΔH298 = −44.22kJ/mol (4)

Fe3O4 is fully oxidized back to Fe2O3 by air, which is then returned to the FR for a new cycle, and the AR exit stream contains N2 and some excess O2. This reaction is strongly exothermic to sustain the thermal balance of the whole system. The CLHG process is intrinsically very attractive, because the exhausts from the SR consist mainly of H2 and water vapor. Moreover, the gas from the reduction reactor consists of CO2 and water vapor, so a condenser is the only equipment needed to obtain almost-pure CO2 or H2. Compared to conventional technologies, such as steam reforming of natural gas, thermal cracking of natural gas, partial oxidation of heavy fractions, or coal gasification, additional CO2 separation costs, H2 purification, and the water−gas shift are avoided. Thus, no energy is lost in the separation, and no costs associated with gas separation equipment and operation are incurred.10 To develop this promising technology for hydrogen generation, many studies have been carried out. Chiesa et al.11 analyzed the three-reactor chemical-looping process for hydrogen production from natural gas using iron oxides as the oxygen carrier. In their study, two plant configurations based on this process were presented in comparison with a “steam reforming” plant. The results showed that the global efficiency can reach 76−78% and that an impressive potential exists for hydrogen production, so that substantial research and development activities are warranted in the near future. Cleeton et al.12 simulated a CLHG system in conjunction with a steam−coal gasification process using Fe2O3 as the oxygen carrier. Their results showed that the peak exergetic efficiencies could reach 53.7% and 59.7% at operating pressures of 1 and 10 atm, respectively. These values compare favorably with those achieved by hydrogen production through the steam reformation of methane. The system show considerable promise for the

2. THERMODYNAMIC ANALYSIS AND REACTOR DESIGN 2.1. Thermodynamic Analysis. It is important that the iron oxides used as oxygen carriers can fully convert the fuel gas to CO2 and H2O. Fuel gases, such as CH4, CO, and H2, leaving the fuel reactor, would mean a loss in combustion efficiency.18 Mattisson et al.19 simulated chemical reactions and performed equilibrium calculations using the Gibbs energy minimization method. Svoboda et al.6,20 studied the reduction of Fe3O4 to Fe by H2, CO, CH4, and syngas from the thermodynamic and chemical equilibrium points of view. Their results showed that all iron oxides except Fe2O3/Fe3O4 are thermodynamically unsuitable as oxygen carriers because of thermodynamic equilibrium limitations.18 The Bauer−Glaessner diagram in Figure 2 shows the equilibrium gas compositions of the Fe−CO−CO2 and Fe−H2−H2O systems at different temperatures. It can be seen that the conversion of H2 and CO is complete when Fe2O3 is reduced to Fe3O4 (reactions 5 and 6). For example, the equilibrium concentration of CO/(CO2 + CO) is about 0.005% and that of 4268

dx.doi.org/10.1021/ie201052r | Ind. Eng. Chem. Res. 2012, 51, 4267−4278


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Figure 3. Fuel reactor concept for CLHG. Figure 2. Bauer−Glaessner diagram.

H2/(H2 + H2O) is 0.003%, when Fe2O3 is converted to Fe3O4 at 900 °C. 3Fe2O3 + CO → 2Fe3O4 + CO2 ΔH298 = −50.73kJ/mol

(5)

3Fe2O3 + H2 → 2Fe3O4 + H2O ΔH298 = −9.59kJ/mol

(6)

However, the reduction of Fe3O4 to FeO is a reversible reaction: the greater the reduction of iron oxides, the higher the equilibrium conversions of CO/(CO2 + CO) and H2/ (H2 + H2O) (reactions 7 and 8). For example, when reactions 7 and 8 reach thermodynamic equilibrium at 900 °C, the equilibrium concentration of CO/(CO2 + CO) is about 45% and that of H2/(H2 + H2O) is 37%. This means that it is difficult to fully convert the fuel gas to CO2 and H2O when Fe3O4 is deeply reduced to FeO.

Figure 4. Schematic diagram of the compact fluidized bed.

fuel reactors are designed as a bubbling bed and a fast bed, respectively, through the selection of a suitable cross-section ratio γ. The solids inlet is located near the bottom of the upper fuel reactor. Fe2O3 enters from the solids inlet and is then thrown upward by the high-speed gas. Part of Fe2O3 is reduced to Fe3O4 in the upper fuel reactor by using the unconverted fuel gases from the lower part. After that, Fe3O4 and unreacted Fe2O3 pass through the cyclone, downcomer, and loop seal continuously and finally enter the lower fuel reactor, where the Fe3O4 and the unreacted Fe2O3 are further reduced to FeO by fuel gas. The major advantage of the compact fluidized bed in this process is that the fuel gas can be completely converted to CO2 and H2O when Fe2O3 is reduced to FeO. Therefore, CO2 can be readily captured by condensing water vapor, eliminating the need for an additional energy-intensive CO2 separation. The process simulation of a compact fluidized bed using Aspen Plus software was performed by Xiang et al.21 Their results showed that, when syngas was assumed as the fuel gas, the unconverted CO from lower fuel reactor was 21.92% and the unconverted H2 was 7.81% whereas the CO and H2 from the lower fuel reactor were almost completely converted into CO2 and H2O after the upper fuel reactor. Although process simulation gives only ideal results compared with the real process, it still indicates that this reactor concept is of value to research. Simultaneously, the structural features of the compact fluidized bed also offer some other advantages for CLHG. The reduction of Fe3O4 to FeO requires more reaction time than that of Fe2O3 to Fe3O4, so the lower fuel reactor is operating in the bubbling fluidized-bed regime, which supplies a sufficient mean particle residence time for the deep reduction of Fe3O4 to FeO. Compared with a traditional bubbling bed, the slip of fuel gas passing through the bubble phase is not tolerable in this fuel

2Fe3O4 + 2CO → 6FeO + 2CO2 ΔH298 = 67.22kJ/mol

(7)

2Fe3O4 + 2H2 → 6FeO + 2H2O ΔH298 = 149.49kJ/mol

(8)

From this analysis, it can be concluded that the conversions of CO and H2 are low when Fe2O3 is reduced to FeO in a single step.17,21 Consequently, it is necessary to find a rational way to solve this problem. 2.2. Compact Fluidized-Bed Concept. As the key part of the CLHG system, the fuel reactor must satisfy at least two requirements: (1) In the fuel reactor, the Fe2O3 should be deeply reduced to FeO that can be used to generate hydrogen. (2) The conversion of the fuel gas should be high, which means that high-purity CO2 can be obtained after water condensation. In this section, a novel fuel reactor concept for CLHG that can meet these two requirements is proposed (as shown in Figure 3). This concept is best described as follows: The fuel gas, for example, syngas, first reduces Fe3O4 to FeO, and the unreacted syngas in those reactions is fully converted to CO2 and H2O by participating in the irreversible Fe2O3 reduction to Fe3O4. All of these reactions take place in a novel fuel reactor, a compact fluidized bed, presented schematically in Figure 4. The compact fluidized bed is considered as consisting of two parts, where the lower fuel reactor has a wider cross section than the upper fuel reactor. Different cross sections lead to different superficial gas velocities in these two parts. The lower and upper 4269



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was chosen. The cross-section ratio γ was determined by the superficial gas velocities of the lower and upper parts of the fuel reactor, ulow and uup u up S γ = low = S up ulow (11)

reactor, as the unconverted fuel gas could be oxidized in the upper fuel reactor, and therefore, the measures used to minimize the gas slip are not necessary, such as reducing fluidization numbers or providing sufficient bed height, which will lead to relatively large fuel-reactor-bed cross-sectional areas and high solids inventories.22 Another beneficial peculiarity of this fuel reactor is that it can decrease carbon deposition due to the reverse Boudouard reaction or methane decomposition 2CO → CO2 + C

ΔH298 = − 172.42kJ/mol

CH 4 → C + 2H2

ΔH298 = 74.60kJ/mol

where the cross section of the lower fuel reactor, Slow, was determined by the thermal power and the particle properties. To adapt to different types of fuel gas, the fuel gas was assumed to undergo no volume expansion, although methane was selected as the fuel gas. Therefore, the cross section of the upper fuel reactor, Sup, could be easily calculated as

(9) (10)

S u S up = low = Slow low u up γ

The deposited carbon mainly reduces the purity of hydrogen when the carbon is carried over to the steam reactor.16 Both reactions can be readily catalyzed by many transitional metals such as nickel and iron. For example, the carbon deposition will be enhanced by the direct contact of fuel gas with Fe particles. Gupta et al.16 reported that the carbon deposition will be significantly diminished if the fuel gas makes contact with iron oxide particles that are not fully reduced. Hence, the compact fluidized bed can help reduce carbon deposition caused by Fe catalysis, as all types of iron oxides are distributed evenly by the strong disturbance in the lower fuel reactor. 2.3. Three-Fluidized-Bed Reactor System. Based on the novel concept of compact fluidized-bed fuel reactor, a threefluidized-bed reactor system representing a 50-kW CLHG prototype was designed and a cold-flow model was erected at Southeast University (Nanjing, China) in 2010. This hot rig is operated at atmospheric pressure. The type of fuel, characteristics of the oxygen carrier, operating temperature, and bed operating regime are as follows: Many types of fuel, even coal, could be chosen for this system. However, methane is considered most suitable for the process. Because one molecule of methane can generate three molecules of reactants, gas volume expansion further increases the gas velocity in the upper fuel reactor, which supports solids entrainment. For example, the conversion of the methane after the lower fuel reactor was assumed to be 80%, as calculated according to the Aspen Plus at 900 °C, so the gas volume in the upper fuel reactor would be 1.15 times that in the lower fuel reactor. In other words, methane can help to extend the operating range of fuel gas flow. Thus, the fuel reactor can still operate at a relatively low fuel gas flow. In addition, some other fuels, such as syngas from coal gasification or even coal, could be considered in our future work. The type of oxygen carrier is considered to be one of the most crucial inputs in the design process.23 In this work, iron oxides were used as the oxygen carrier for CLHG, because they provide the best conversion rate of fuel gas along with a high conversion of steam to hydrogen.16 The mean diameter of the oxygen-carrier particles was chosen to be between 100 and 350 μm. The range of particle density was between 2500 and 5600 kg/m3, as some inert carrier support and binder might be mixed to improve the reactivity, durability, and longevity of the oxygen carrier. The design temperature was assumed to be about 900 °C in the three reactors, ignoring the effects of the energy balance. The lower fuel reactor was designed as a low-velocity bubbling bed for a higher particle residence time, with the superficial gas velocity assumed to be 0.7 times the terminal velocity. However, for the entrainment requirement of the upper fuel reactor, a superficial velocity close to 4−10 times the terminal velocity

(12)

The steam reactor (SR) has the main requirement of obtaining hydrogen and transporting the oxygen carrier to the air reactor (AR) for deep oxidization. Therefore, the SR was designed as a fast fluidized bed. Similarly, to provide sufficient driving force for the solids circulation, the AR was designed as a fast fluidized bed. Similarly to the transport reactor developed by the Chalmers University of Technology,23 both the SR and AR had wider bottom sections in which a higher mean particle residence time could be obtained. A sketch of the prototype of the three-fluidized-bed reactor system for CLHG is shown in Figure 5, and the most important geometric data and other design parameters are summarized in Table 1.

3. EXPERIMENTS 3.1. Experimental Setup. A cold-flow model for CLHG of gaseous fuels was designed and built in 2010 by applying the scaling rules of Glicksman and co-workers,24,25 and the relationships between the hot rig and the cold-flow model are listed in the Table 2. Glass beads were used as the particles with a particle density of 2500 kg/m3 and mean particle diameter of 200 μm. To maintain the density ratio between the solids and the gas in the hot rig, a gas mixture, for example, air/helium, could be chosen. A schematic of the cold-flow model setup is shown in Figure 6. The model is composed of three interconnected fluidized-bed reactors, the FR, SR and AR, arranged in triangle and connected by three cyclones, four downcomers, a ball valve, an L-type loop seal, and three V-type loop seals. To balance the pressure difference between the FR and the SR, a ball valve was placed near the L-type loop seal. The handle turning angle of the valve was held constant at 45° in all of the experiments. All of the reactors and parts were built from acrylic glass, and the cyclones were designed according to the design formulas of Hoffmann and Stein.26 The distributor plates of both the reactors and the loop seals were of nozzle type. Pressure fluctuations in this system were obtained by a multichannel differential pressure signal sampling system. A total of 30 pressure taps and 20 pressure differential transmitters (type Xi’an XinMin) with a scale of 0−10 kPa were used. The voltage signals were sent to a computer through an analogto-digital converter. A Roots-type blower supplied the fluidizing gas. The gas flow rates were measured with 11 flow meters (type ChangZhou ShuangHuan flow). A gas holder was installed to avoid pressure oscillations and to achieve a steady gas flow. The gas leaving the reactors was analyzed with an Emerson gas analyzer (including O2, CO, CO2, CH4, and H2). 4270



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Figure 5. Design layout of a three-fluidized-bed reactor system for CLHG: (1) lower fuel reactor, (2) upper fuel reactor, (3) low-velocity zone of the steam reactor, (4) riser of the steam reactor, (5) low-velocity zone of the air reactor, (6) riser of the air reactor, (7) loop seal, (8) L-type loop seal, (9) ball valve, (10) screwed feeder.

Table 1. Design Values of the CLHG Hot Rig parameter

units

thermal power fuel type operating temperature pressure fluidization gas in the FR fluidization gas in the SR fluidization gas in the AR particles mean particle diameter particle size distribution

kW

particle density gas fluidization velocity in the lower fuel reactor (u/ut) gas fluidization velocity in the upper fuel reactor (u/ut gas fluidization velocity in the low-velocity zone of the SR and AR (u/ut) gas fluidization velocity in the riser of the SR and AR (u/ut) lower-fuel-reactor diameter lower-fuel-reactor height upper-fuel-reactor diameter upper-fuel-reactor height diameter of the low-velocity zone of the SR height of the low-velocity zone of the SR SR riser diameter SR riser height diameter of the low-velocity zone of the AR height of the low-velocity zone of the AR AR riser diameter AR riser height

kg m−3

°C Pa

m m

Table 2. Relationships between the Hot Rig and the ColdFlow Model

value 50 methane 900 1 × 105 methane steam air iron oxides 260 × 10−6 (100−350) × 10−6 3850 0.7 4−10 1.5−3

parameter temperature pressure fluidization gas in the FR fluidization gas in the SR fluidization gas in the AR particles mean particle diameter particle size distribution particle density length superficial velocity solids flux volume flow

°C Pa

m m kg m−3

hot rig

cold-flow model

900 1 × 105 CH4 steam air oxygen carrier 260 × 10−6 (100−350) ×10−6 3850 1 1 1 1

25 1 × 105 He/air He/air He/air glass beads 200 × 10−6 200 × 10−6 2500 0.77 0.87 0.57 0.52

were carried out in the cold-flow model system. The main purpose of the tests was not to simulate hydrodynamics of hot rig but rather to assess the feasibility and understand the operating rules of this system. Therefore, air was selected as the fluidization gas in these tests for qualitative study. More accurate hydrodynamics simulations using helium as the fluidization gas will be performed in future work. For the solids circulation rate test, a tracer particle method was used. Black tracer particles were added to the reactor system from the screwed feeder in small quantities, and both the black tracer particles and white glass beads had similar particle properties to improve the accuracy of the tests. As can be seen from Figure 7, the solids circulation rate was determined by repeatedly measuring the time required for a tracer particle to pass a fixed length in the downcomer. All values for the solids flow in this section represent averages from a minimum of 10 measurements each. Compared with methods that abruptly stop the fluidization

4−10 m m m m m m m m m m m m

units

0.17 0.78 0.05 1.90 0.10 0.71 0.05 1.78 0.10 1.04 0.05 2.92

3.2. Experimental Procedures. The solids circulation rate measurements, gas leakage tests, and long-term operation test 4271



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Figure 6. Schematic of the cold-flow model setup: (1) CO2, (2) Roots-type blower, (3) gas holder, (4) flow meters, (5) ball valve, (6) fuel reactor, (7) steam reactor, (8) air reactor.

A series of experiments about solids circulation was performed. The gas flows to the three reactors and L-type loop seal were changed. Also, the influence of the total solids inventory was studied. The main purpose of these tests was to understand the variation of the solids circulation rate when operating conditions were changed and to gain initial operating experience. For the gas leakage tests a tracer gas method was used.27−29 CO2 was added to the inlet air in one of the reactors (FR, SR, or AR). The leakage is defined as the fraction of gas added to one reactor that escapes out of the other reactor or cyclone. In this test, the leakages include gas leakage from the AR to the hydrogen cyclone, from the FR to the air cyclone, and from the FR to the SR. Concentrations at all measuring points were measured with a gas analyzer, and the quality of the results was verified by performing concentration measurements of the air/ CO2 gas before leaking to another reactor.28 3.3. Data Evaluation. For better comparison of the leakage, for example, from reactor A to reactor B, the leakage was magnified by assuming 100% CO2 was used as the fluidization gas in reactor A, even though the CO2 gas stream was only part of the fluidization gas in reactor A. The leakage was thus calculated as leakage(%) =

xCO2B xCO2A

× 100% (13)

where xCO2A denotes the concentration of CO2 in reactor A measured before leaking to reactor B and xCO2B is the concentration of CO2 leaked to reactor B. The variation in xCO2i caused by the flow rate to reactors A and B was corrected before calculatino of the leakage. Because most of the bed materials accumulated in the lower fuel reactor, another important parameter, the bed mass

Figure 7. Schematic diagram of the tracer particle method for the solids circulation rate test.

of loop seal, this method is simple to apply and has a high accuracy because it involves almost no disturbance of the system. 4272



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distribution, λFR, describes the distribution state of the particles in the three reactors and is calculated from the static volumes of bed materials in the FR, SR, and AR λFR =

VFR VFR + VSR + VAR

structure. It can be seen that the lower fuel reactor shows a distinct decay of pressure. However, there is almost no pressure drop in the transition section of the FR. This means that few particles pass through the transition section. From the perspective of the actual solids distribution in the system, fewer solids are present in the SR and AR than in the lower fuel reactor, and the upper fuel reactor is relatively lean in solids. 4.2. Solids Circulation Rate. In the CLHG system, the solids circulation rate has to provide sufficient oxygen transport capacity and energy transfer among the reactors.28 Therefore, it is necessary to investigate the solids circulation rate of the three-fluidized-reactor system with respect to different flow rates to reactors, total solids inventories (TSIs), and bed mass distributions. The downcomer below the FR was the measuring point determining the specific solids circulation rate, Gs. Gs can also be considered as the downcomer flux, and all values of Gs reported in this article are referred to the cross section of the downcomer. Figure 9 shows the specific solids circulation rate, Gs, as a function of the flow rate to the AR and the total solids inventory.

(14)

where Vi denotes the static volume of bed materials in reactor i. The volumes of bed materials in downcomers and loop seals were almost constant for different total solids inventories and are not included in eq 14. A more detailed solids distribution for a typical operating case (Table 3) that contains the bed materials in the downcomers and loop seals is reported in Table 4. Table 3. Typical Operating Case air flow

units

fuel reactor steam reactor air reactor L-type loop seal total solids inventory

3

Nm N m3 N m3 N m3 kg

value −1

h h−1 h−1 h−1

24 31 30 0.3 5

Table 4. Solids Distribution for the Typical Operating Case position

units

value

fuel reactor steam reactor air reactor downcomers and loop seals

% % % %

76.1 ∼0.4 ∼0.2 23.2

4. RESULTS AND DISCUSSION 4.1. Pressure Loop. Figure 8 shows the pressure loop profile of the cold-flow model for a typical operating case according to

Figure 9. Specific solids circulation rate as a function of the flow rate to the AR and the total solids inventory.

The volume flows in the FR and SR were kept constant at 24 N m3 h−1 (1 atm, 293 K) and 31 N m3 h−1, respectively. It was found that the specific solids circulation rate, Gs, increased as the flow rate to the AR increased. This is most likely because a higher flow rate to the AR provides a higher driving force for the overall system. In addition, with increasing total solids inventory, the specific solids circulation rate, Gs, increased. As can be seen from Figure 10, the bed mass distribution, λFR, was strongly influenced by the flow rate to the AR. Also, it can

Figure 8. Pressure loop profile of the cold-flow model (total solids inventory = 5 kg, VFR = 24 m3/h, VSR = 31 m3/h, VAR = 30 m3/h).

Table 3, where the pressure values are relative to the ambient atmosphere. As can be seen, the loop seals and L-type loop seal can balance the pressure differences between the reactors, and the system can maintain a stable and continuous circulation of solids. The pressure profile also gives a good indication regarding the distribution of solids. Both the AR and SR show typical distributions for a fast fluidized riser. The decay is almost linear, indicating an almost-constant void fraction over the whole height. However, the FR shows a particular pressure curve for its special

Figure 10. Bed mass distribution as a function of the flow rate to the AR and the total solids inventory.

be observed from the cold-flow model that the bed height of the FR declined gradually and more materials accumulated at 4273



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the bottom of the AR as the velocity in the AR decreased. Relatively low fluidization rates in the AR led to the bottom of the AR operating in the bubbling fluidized-bed regime. However, owing to the smaller cross section of the riser, the AR could still supply sufficient solids entrainment for the system mass balance. Figure 10 also shows the bed mass distribution as a function of the total solids inventory. It can be seen that the total solids inventory had no significant effect on the bed mass distribution. The influence of the specific solids circulation rate, Gs, on the flow rate to the SR and the total solids inventory are presented in Figure 11. Generally, the solids circulation rate decreased

Figure 12. Bed mass distribution as a function of the flow rate to the SR and the total solids inventory.

and 4 kg, respectively. This is because the relatively low bed mass distributions of points B and C lead to an increase of the specific solids circulation rate, especially for point C (shown in Figure 12), with the lowest bed mass distribution giving an obvious increase of the specific solids circulation rate. Figure 13 presents the specific solids circulation rate, Gs, as a function of the flow rate to the FR and the total solids

Figure 11. Specific solids circulation rate as a function of the flow rate to the SR and the total solids inventory.

as the flow rate to the SR increased. This is a consequence of the increased pressure of the SR, and thus the increased pressure difference between the FR and SR, which enhanced the resistance of solids entering the SR. However, the curve for 6 kg increased first at low fluidization rates and then decreased after passing through a maximum. This is because of the decrease in the entrainment capability of the SR at low velocity, with more materials accumulating at the bottom of the SR. The pressure drop in the SR was mainly a result of overcoming the gravity of the solids, which weakened the influence of pressure variations on the specific solids circulation rate. Therefore, the specific solids circulation rate, Gs, increased with improving entrainment capability until the SR had little accumulation. Then, the pressure became the main factor influencing the specific solids circulation rate. The curves for 4 and 5 kg at low fluidization rates are discussed below. The specific solids circulation rate is also influenced by the total solids inventory (Figure 11). It can be observed that the specific solids circulation rate increased when the total solids inventory change from 5 to 6 kg. However, the specific solids circulation rate exhibited little change when the total solids inventory increased from 4 to 5 kg. This phenomenon can be explained by the bed mass distribution shown in Figure 12. It can be seen that total solids inventories of 5 and 6 kg have similar bed mass distributions, so it can be considered that the specific solids circulation rate increased with the increase in the total solids inventory under the same bed mass distribution conditions. However, the bed mass distribution for 4 kg was substantially lower than that for 5 and 6 kg, so the phenomenon observed in the experiments showed that the decrease in the bed mass distribution can lead to an increase in the specific solids circulation rate, which is similar to the results reported by Kronberger et al.23 In addition, as can be seen from Figure 11, point A represents a relatively low value on the 6-kg curve, whereas points B and C represent the maximum values for 5

Figure 13. Specific solids circulation rate as a function of the flow rate to the FR and the total solids inventory.

inventory. The flow rate to the FR strongly influences the specific solids circulation rate. On one hand, it was found that, for gas flow in the FR lower than 15 N m3 h−1, the entrainment capability of the upper fuel reactor decreased and part of the solids directly fell into the lower fuel reactor. This problem can be solved by introducing secondary air below the inlet of the FR. On the other hand, for the gas flow in the FR higher than about 30 N m3 h−1, the bed height of the lower fuel reactor increased, and part of the solids in the lower fuel reactor rose to the upper fuel reactor. Although the entrainment in the lower fuel reactor does not affect the operation of the overall system, this is considered not to be a desired operating situation for the system. The influence of the total solids inventory on the specific solids circulation rate is similar to that discussed before. As expected, the specific solids circulation rate increased when more bed material was present in the system. This is because the hydrostatic pressure pushing particles from the FR to the SR increased and, thereby, increased the solids circulation rate of the system.30 With the aid of Figure 14, it can be seen that the specific solids circulation rate was again influenced by the variation of bed mass distribution. The relatively small bed mass distribution for 4 kg led to an increase of the specific solids circulation rate. It can be seen from Figure 15 that the flow rate to the L-type loop seal had less of an influence on the specific solids circulation rate. However, in general, the specific solids circulation rate 4274



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These velocities are related by the equation Δu = ug + us

(15)

When |ug| > |us|, gas is able to move upward against the downward-flowing solids, and leakage occurs. To investigate the leakage from the AR to the hydrogen cyclone, the flow rates to both the AR and SR were varied, and CO2 was added to the fluidizing gas of the AR. The solid line in Figure 16 shows the leakage as a function of the flow rate to the

Figure 14. Bed mass distribution as a function of the flow rate to the FR and the total solids inventory.

Figure 15. Specific solids circulation rate as a function of the flow rate to the L-type loop seal and the total solids inventory. Figure 16. Leakage from the air reactor to the hydrogen cyclone as a function of the flow rates to the AR (solid line) and SR (dashed line).

increased first at low flow rates and then basically stabilized after passing through a maximum. This might be because the solids were fluidized by the gas at 0.6−0.7 N m3 h−1, and the three curves in Figure 15 have similar characteristics to the pressuredrop-versus-velocity diagram used to ensure umf. In other words, the L-type loop seal can influence the solids circulation rate only by changing the fluidized state of the downcomer. In addition, because the flow rate of the L-type loop seal had almost no influence on the bed mass distribution, the specific solids circulation rate increased with increasing total solids inventory, as expected. In general, a wide operating range of solids circulation rate can be achieved by varying the fluidization conditions in three reactors. Moreover, the solids circulation rate is hardly influenced by the flow rates in the L-type loop seal. On the other hand, the experimental results discussed in this section can also explained from the viewpoint of mass balance. The FR can be considered as the center reactor of the whole system because most of the material is distributed in it. Then, the AR and the SR are located in the inlet and outlet of the fuel reactor, respectively. Therefore, increasing the drive force of the inlet can help increase the solids circulation rate, whereas increasing the fluidization rate of the outlet might prevent solids circulation. 4.3. Leakage. It is important to minimize the leakage among the reactors in a CLHG system. The basic reason for leakage is gas passing through the downcomer from the higherpressure part to the lower-pressure part. In this system, all of the downcomers were operated as moving beds. Consider a fixed bed of solids moving downward at a velocity us (downward positive) with a gas flowing upward at a velocity of ug, where Δu is the relative velocity of the gas with respect to the solids.

AR, where the flow rates to the FR and SR were held constant at 24 and 31 N m3 h−1, respectively. As can be seen, the leakage increased as the flow rate to the AR increased. This is because the pressure difference balanced by the downcomer increased as the flow rate to the AR increased, which led to an increase of the gas velocity in the downcomer, ug. In addition, as the velocity in the AR increased, the inventory in the AR decreased, and the pressure increased more obviously. This explanation agrees well with the gradient of the solid line. The dashed line in Figure 16 shows the leakage when the flow rate to the SR was varied. The flow rates to the AR and FR were held constant at 26 and 24 N m3 h−1 respectively. It can be seen that the leakage passed through a maximum and decreased at high velocities. This might be caused by the combined effects of the solids circulation rate and the pressure difference in the downcomer. On one hand, the solids circulation rate decreased as the flow rate to the SR increased, thus decreasing the velocity of the solids in the downcomer, us. On the other hand, the pressure difference in the downcomer increased as the flow rate to the SR increased, thus increasing the gas velocity in the downcomer, ug. When the velocity in the SR was low, the relatively large inventory of solids in the SR reduced the effect of the pressure difference, and the leakage increased because the effect of the solids circulation rate dominated. When the velocity in the SR was high, the SR operated in the fast fluidization regime. The variation in the pressure became the primary factor affecting the leakage, and the leakage decreased as the pressure difference decreased. The leakage from the FR to the air cyclone was tested by changing the flow rates to the FR and AR. The influence of 4275



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pressure variation. The leakage was mainly influenced by the solids circulation rate. This indicates a close relationship between gas leakage and the solids circulation rate. 4.4. Long-Term Operation. It is important to check the feasibility of stable and continuous operation in the rather complex arrangement of a three-circulation-loop system. A longterm operation test with little adjustment of operating variables was performed. The test was performed for a typical operating case according to Table 3. Figure 19 shows part of the pressure

leakage on the flow rate to the FR is shown by the solid line in Figure 17. It indicates that the leakage can be reduced by

Figure 17. Leakage from the fuel reactor to the air cyclone as a function of the flow rates to the FR (solid line) and AR (dashed line).

increasing the gas flow of the FR. Although the pressure difference balanced by the downcomer increased as the gas flow rate of the FR increased, which led to an increase in the gas velocity in the downcomer, ug, the influence of pressure in the upper fuel reactor on the fluidization rates in the FR is not obvious, as most of the pressure drop is due to the solids gravity in the bubbling bed in the lower fuel reactor. However, when the gas flow rate of the FR was high, it was difficult for the gas to move up against the downward-flowing solids.29 Therefore, the leakage was influenced mainly by the solids circulation rate, and the leakage decreased as the solids circulation rate increased. Based on the analysis in this section, it can be known that, as the flow rate to the AR increased, both an increase of the solids circulation rate and a decrease of the pressure difference in the downcomer can reduce the gas leakage from the FR to the air cyclone. This analysis is consistent with the test results, and it can be seen from Figure 17 that the leakage was reduced by increasing the gas flow rate of the AR. Figure 18 shows the leakage from the FR to the SR as a function of the flow rates to the FR and SR. It can be seen that

Figure 19. Pressure drop profiles during long-term operation.

drop profiles of the three reactors during stable operation. As can be seen, both the upper and lower fuel reactors had larger pressure drop fluctuations than the SR and AR. This was caused by the bubbles passing through the fuel reactor. Moreover, through continuous running, the results show that the system is stable and able to operate for long times with minimum adjustments of the flow rate to the loop seals.

5. CONCLUSIONS A novel compact fluidized-bed fuel reactor concept for CLHG has been proposed in this article. It can help reduce the iron oxides deeply to FeO or Fe while simultaneously fully converting the fuel gas to CO2 and H2O; after the water has been condensed, high-purity CO2 can be obtained. Based on the compact fluidized-bed concept, a 50-kW hot rig scheme of a three-fluidized-bed system designed for CLHG was presented; a cold-flow model was designed and built; and a series of test with respect to solids circulation rate, gas leakage, and long-term operation were performed. In more detail, the following conclusions can be drawn: The experiments on solids circulation rate showed that the solids circulation rate can be varied over a wide range and can be effectively influenced by the fluidization conditions in the three reactors. Moreover, the solids circulation rate also depends on the total solids inventory and the bed mass distribution. Those operating experiences on solids circulation rate can help to control the heat and mass balance of the three-fluidized-bed reactor system, such as adjusting the oxygen transport capacity and the temperature difference among reactors. The gas leakage among the reactors was measured by CO2 tracer gas methods. In general, it was found that gas leakage is directly related to the solids circulation rate and the pressure difference balanced by the downcomer. The leakage was normally lower than 2% when the system was operated near the standard

Figure 18. Leakage from the fuel reactor to the steam reactor as a function of the flow rates to the FR (solid line) and SR (dashed line).

increasing the flow rate to the FR reduced the leakage. As in the discussion of Figure 16, compared with the pressure difference between the downcomer, which can also reduce the leakage in this case, the solids circulation rate is a major contributing factor to the gas leakage. The influence of the leakage on the flow rate to the SR is also shown in Figure 18: When the SR was operated in the bubble bed regime with a low gas flow, there was little 4276



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(10) Abad, A.; Mattisson, T.; Lyngfelt, A.; Johansson, M. The use of iron oxide as oxygen carrier in a chemical-looping reactor. Fuel 2007, 86 (7−8), 1021−1035. (11) Chiesa, P.; Lozza, G.; Malandrino, A.; Romano, M.; Piccolo, V. Three-reactors chemical looping process for hydrogen production. Int. J. Hydrogen Energy 2008, 33 (9), 2233−2245. (12) Cleeton, J. P. E; Bohn, C. D.; Muller, C. R.; Dennis, J. S.; Scott, S. A. Clean hydrogen production and electricity from coal via chemical looping: Identifying a suitable operating regime. Int. J. Hydrogen Energy 2009, 34 (1), 1−12. (13) Gnanapragasam, N. V.; Reddy, B. V.; Rosen, M. A. Hydrogen production from coal using coal direct chemical looping and syngas chemical looping combustion systems: Assessment of system operation and resource requirements. Int. J. Hydrogen Energy 2009, 34 (6), 2606−2615. (14) Chen, S. Y.; Xiang, W. G.; Xue, Z. P.; Sun, X. Y. Experimental investigation of chemical looping hydrogen generation using iron oxides in a batch fluidized bed. Proc. Combust. Inst. 2011, 33, 2691− 2699. (15) Yang, J. B.; Cai, N. S.; Li, Z. S. Hydrogen production from the steam−iron process with direct reduction of iron oxide by chemical looping combustion of coal char. Energy Fuels 2008, 22 (4), 2570− 2579. (16) Gupta, P.; Velazquez-Vargas, L. G.; Fan, L. S. Syngas redox (SGR) process to produce hydrogen from coal derived syngas. Energy Fuels 2007, 21 (5), 2900−2908. (17) Li, F. X.; Zeng, L. A.; Velazquez-Vargas, L. G.; Yoscovits, Z.; Fan, L. S. Syngas chemical looping gasification process: Bench-scale studies and reactor simulations. AIChE J. 2010, 56 (8), 2186−2199. (18) Jerndal, E.; Mattisson, T.; Lyngfelt, A. Thermal analysis of chemical-looping combustion. Chem. Eng. Res. Des. 2006, 84 (9), 795− 806. (19) Lyngfelt, T. M. a. A. Capture of CO2 using chemical-looping combustion. Presented at the First Biennial Meeting of the ScandinavianNordic Section of the Combustion Institute, Göteborg, Sweden, April 18− 20, 2001. (20) Svoboda, K.; Siewiorek, A.; Baxter, D.; Rogut, J.; Pohorely, M. Thermodynamic possibilities and constraints for pure hydrogen production by a nickel and cobalt-based chemical looping process at lower temperatures. Energy Convers. Manage. 2008, 49 (2), 221−231. (21) Xiang, W.; Chen, S.; Xue, Z.; Sun, X. Investigation of coal gasification hydrogen and electricity co-production plant with threereactors chemical looping process. Int. J. Hydrogen Energy 2010, 35 (16), 8580−8591. (22) Kolbitsch, P.; Proll, T.; Bolhar-Nordenkampf, J.; Hofbauer, H. Design of a chemical looping combustor using a dual circulating fluidized bed reactor system. Chem. Eng. Technol. 2009, 32 (3), 398− 403. (23) Kronberger, B.; Lyngfelt, A.; Loffler, G.; Hofbauer, H. Design and fluid dynamic analysis of a bench-scale combustion system with CO2 separation-chemical-looping combustion. Ind. Eng. Chem. Res. 2005, 44 (3), 546−556. (24) Glicksman, L. R.; Hyre, M. R.; Farrell, P. A. Dynamic similarity in fluidization. Int. J. Multiphase Flow 1994, 20, 331−386. (25) Glicksman, L. R.; Hyre, M.; Woloshun, K. Simplified scaling relationships for fluidized beds. Powder Technol. 1993, 77 (2), 177− 199. (26) Hoffmann, A. C.; Stein, L. E Gas Cyclones and Swirl Tubes: Principles, Design and Operation; Springer: New York, 2004 (27) Ryu, H. J.; Lee, S. Y.; Park, Y. C.; Park, M. H.; Ardil, C. Solid circulation rate and gas leakage measurements in an interconnected bubbling fluidized beds. World Acad. Sci., Eng. Technol. 2007, 22, 169− 174. (28) Kronberger, B.; Johansson, E.; Loffler, G.; Mattisson, T.; Lyngfelt, A.; Hofbauer, H. A two-compartment fluidized bed reactor for CO2 capture by chemical-looping combustion. Chem. Eng. Technol. 2004, 27 (12), 1318−1326. (29) Johansson, E.; Lyngfelt, A.; Mattisson, T.; Johnsson, F. Gas leakage measurements in a cold model of an interconnected fluidized

operating conditions. However, the leakage from the FR to the air cyclone could reach 5% under extreme conditions, which means that about 5% of the CO2 would be lost to the atmosphere. To avoid this leakage, it is possible to increase the length of the downcomer below the air cyclone. Moreover, the three-fluidized-bed reactor system showed a stable pressure difference and solids circulation rate during the long-term operation test, and pressure self-balance was achieved in the whole system with little adjustment. Further work will focus on a more accurate hydrodynamic simulation of the three-fluidized-bed reactor system using the existing cold-flow model. Of special interest are the pressure characteristics of different operating conditions. Such data will be central to operating and controlling the hot rig in the future. Another interesting topic is to assess the possibilities for the compact fluidized bed to apply to other areas. For example, it can help to enhance the oxygen-carrying capacity of iron oxides in a chemical-looping combustion system. In addition Ca-looping CO2 capture is also consider as a suitable application fields for this reactor.

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AUTHOR INFORMATION

Corresponding Author

*Tel.: +86 25 8379 5545. Fax: +86 25 8771 4489. E-mail: wgxiang@ seu.edu.cn.

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ACKNOWLEDGMENTS The authors express thanks to the National Natural Science Foundation of China (51176033) and the Special Fund of the National Priority Basic Research of China (2007CB210101) for financial support of this project.

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REFERENCES

(1) Winter, C. J. Hydrogen energyAbundant, efficient, clean: A debate over the energy-system-of-change. Int. J. Hydrogen Energy 2009, 34 (14), S1−S52. (2) Proll, T.; Bolhar-Nordenkampf, J.; Kolbitsch, P.; Hofbauer, H. Syngas and a separate nitrogen/argon stream via chemical looping reformingA 140 kW pilot plant study. Fuel 2010, 89 (6), 1249− 1256. (3) Mattisson, T.; Lyngfelt, A.; Cho, P. The use of iron oxide as an oxygen carrier in chemical-looping combustion of methane with inherent separation of CO2. Fuel 2001, 80 (13), 1953−1962. (4) Messerschmidt, A. Process of Producing Hydrogen. U.S. Patent 971,206, 1910. (5) Ryu, J. C.; Lee, D. H.; Kang, K. S.; Park, C. S.; Kim, J. W.; Kim, Y. H. Effect of additives on redox behavior of iron oxide for chemical hydrogen storage. J. Ind. Eng. Chem. 2008, 14 (2), 252−260. (6) Svoboda, K.; Slowinski, G.; Rogut, J.; Baxter, D. Thermodynamic possibilities and constraints for pure hydrogen production by iron based chemical looping process at lower temperatures. Energy Convers. Manage. 2007, 48 (12), 3063−3073. (7) Lorente, E.; Pena, J. A.; Herguido, J. Separation and storage of hydrogen by steam−iron process: Effect of added metals upon hydrogen release and solid stability. J. Power Sources 2009, 192 (1), 224−229. (8) Urasaki, K.; Tanimoto, N.; Hayashi, T.; Sekine, Y.; Kikuchi, E.; Matsukata, M. Hydrogen production via steam−iron reaction using iron oxide modified with very small amounts of palladium and zirconia. Appl. Catal. A 2005, 288 (1−2), 143−148. (9) Li, F. X.; Kim, H. R.; Sridhar, D.; Wang, F.; Zeng, L.; Chen, J.; Fan, L. S. Syngas chemical looping gasification process: Oxygen carrier particle selection and performance. Energy Fuels 2009, 23 (8), 4182− 4189. 4277



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