Three Dimensional Modeling of a Coal-Fired Chemical Looping

Article pubs.acs.org/EF

Three Dimensional Modeling of a Coal-Fired Chemical Looping Combustion Process in the Circulating Fluidized Bed Fuel Reactor Xiaojia Wang, Baosheng Jin,* Yong Zhang, Yi Zhang, and Xianli Liu Key Laboratory of Energy Thermal Conversion and Control, Ministry of Education, School of Energy & Environment, Southeast University, Nanjing 210096, People’s Republic of China ABSTRACT: In this work, a comprehensive three-dimensional numerical model was developed to simulate the coal-fired chemical looping combustion (CLC) process using ilmenite as oxygen carrier in a pressurized circulating fluidized bed (CFB) fuel reactor. Both gas−solid flow and chemical reactions were considered. The chemical reactions contain steam gasification of coal and subsequent reduction reactions of intermediate gasification products with the oxygen carrier. This model predicted the main features of the complex gas−solid flow behaviors from the velocity and voidage profiles. The concentrations of gas−solid components, the conversions of char and oxygen carrier, and the distributions of reaction rates were also obtained. Meanwhile, further simulations were performed to evaluate the effects of operating variables on the fuel conversion in the fuel reactor. The conversion of carbon in char is directly related to the extent of gasification, which is promoted by increasing the temperature or the fraction of steam in gasification agent. This work indicates a promising way to research the CLC process of solid fuels in a CFB fuel reactor.

1. INTRODUCTION It is generally accepted that the increasing emission of carbon dioxide (CO2) may greatly affect the global climate. Hence, interest has arisen in capturing and sequestering CO2 generated in fossil fuel combustion. However, considerable energy is required to separate and collect CO2 because CO2 is diluted by N2 from air in the conventional combustion system. It is thus important to search for a new combustion method where CO2 can be separated with low energy consumption. Chemical looping combustion (CLC), which provides an inherent feature of isolating CO2, has become an attractive combustion technology with low cost of CO2 separation.1,2 As shown in Figure 1, a conventional CLC system involves two interconnected reactors: a fuel reactor and an air reactor.3,4 The

fuel introduced into the fuel reactor is oxidized to CO2 and H2O by a metal oxide (oxygen carrier). The reduced oxygen carrier particles are transferred to the air reactor, where they are regenerated by the contact with air. In the whole combustion process, the direct contact between air and fuel is circumvented; thus, the flue gas leaving the fuel reactor only includes CO2 and H2O. After the condensation of H2O, almost pure CO2 is obtained with a small energy loss. A key factor in the CLC performance is the selection of oxygen carrier with adequate behavior and properties. In the CLC process with solid fuel, the fuel is physically mixed with the oxygen carrier. Thus, it is inevitable that a part of the oxygen carrier particles would be lost, together with the draining stream of coal ashes. In this context, the use of low-cost materials such as natural minerals or industrial waste products as oxygen carriers becomes very attractive. Ilmenite has been proposed as one of the most promising candidates in the CLC process of solid fuels due to its low price and adequate reactivity and oxygen transport capacity.5−7 In addition, ilmenite has good mechanical stability and fluidizing properties.8 Since the introduction of CLC in 1983, the gaseous fuel CLC process has been successfully demonstrated in different prototype reactors.9−11 However, as coal is a much more abundant fossil fuel in comparison to gas fuel, the adaptation of CLC with solid fuels such as coal and biomass is necessary. At present, the application of solid fuels to CLC mainly involves two possible ways. The first possibility can be easily accomplished by a process integrating coal gasification and chemical looping combustion (IGCC−CLC).11,12 In this process, the syngas produced from coal gasification is used as fuel in a CLC system for power generation with CO2 capture. Nevertheless, the defect of this approach is that the addition of a gasifier would increase the complexity of the system. The second Received: May 8, 2012 Revised: February 22, 2013 Published: February 28, 2013

Figure 1. Schematic diagram of the coal-fired chemical looping combustion. © 2013 American Chemical Society

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dx.doi.org/10.1021/ef302075n | Energy Fuels 2013, 27, 2173−2184

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fluidized bed was simplified as 2D, which would bring large deviations, including the influences of bed walls, the geometry on the flow patterns and other aspects. Hence, it is necessary to develop 3D numerical models to study the CLC performance of solid fuel. In this study, a comprehensive three-dimensional numerical model was developed to simulate the behavior of coal-fired chemical-looping combustion. The gas−solid flow patterns, species distributions, fuel conversion, reaction rates and other important characteristics were predicted. The main contributions of this simulation work compared to the previous research are listed as follows: (1) successful run of the 3D model for the CLC of solid fuels, (2) successful application of CFB technologies into the fuel reactor construction, (3) realization of high-flux (Gs ≥ 200 kg/m2 s) in the CFB fuel reactor, (4) realization of pressurized operation, and (5) flexible optimization of temperature and the fraction of steam in gasification agent to help achieve the high fuel conversion and CO2 concentration together with energy saving.

approach is to integrate gasification and oxygen carrier reduction in the fuel reactor. The reaction steps include gasification of solid fuels and subsequent reactions of intermediate gasification products with the oxygen carrier.13,14 This technical approach has been demonstrated in a variety of laboratory experiments,8,13,15−21 and all these tests show that this technology of solid-fuel CLC is promising. The fuel reactor is usually designed as a bubbling/spout fluidized bed in the CLC system. However, this kind of design has an inherent disadvantage of producing unconverted fuel gas. In the fluidized region, a large proportion of fuel gas passes through the bed in the bubble phase. In the freeboard, the gas− solid reactions are greatly weakened because of the absence of solid particles. Moreover, compared with the CLC process of gaseous fuel, the reactions are much more complex for solid fuels. There are more reaction steps, and volatile matter in the coal could lead to unburnt products.22 Thus, the inadequate gas−solid contact will further lower the conversion of fuel gas. Admittedly, this shortage can be improved by increasing the bed height or decreasing the superficial gas velocity. However, these measures will result in a larger system size and higher solid inventories.23 At the same time, for direct coal-fired CLC, another challenge is to achieve the complete combustion of char. However, for typical bubbling/spout fluidized bed fuel reactors, the carbon loss is inevitable due to an easy elutriation of fine char particles from the freeboard. In order to get higher carbon conversion efficiency, it was proposed to separate the char that has not been gasified and recirculate it back to the fuel reactor,17,24 but this method will obviously increase the complexity of the system. In a word, due to the gas bypassing in the fluidized region, inadequate gas−solid contact in the freeboard, and elutriation of fine char particles, typical bubbling/spout fluidized bed fuel reactors have inherent disadvantages in CO2 exit concentration and fuel conversion. In these respects, yet, a circulating fluidized bed (CFB) fuel reactor is competitive and promising. It is capable of working in the fast fluidization regime, which brings a favorable gas−solid contact over the whole reactor height and allows the operation with lower solid inventories.23,25,26 Meanwhile, as proposed by Shen et al.,17 the conversion of carbon in char could be greatly increased with the recirculation. With the development of hardware and numerical methods, numerical simulations have become more and more important to deal with complex gas−solid hydrodynamics and chemical reactions in the processes such as combustion and gasification. Chen et al.27 has proposed an Eulerian−Eulerian model incorporating the kinetic theory of granular flow to describe the gas−solid two-phase flow in fluidized bed polymerization reactors. Zhou et al.28 established a 2D combustion model to study the air− coal two-phase flow and combustion characteristics in a 50 kW circulating fluidized bed (CFB) combustor. Wang et al.29 developed a 3D numerical model to simulate the coal gasification process in a bubbling fluidized bed gasifier. However, very little attention has been paid to CLC studies via the CFD approach. Jung and Gamwo30 simulated the chemical looping combustion process in a bubbling fluidized bed fuel reactor with Ni-based oxygen carrier. A sensitivity analysis of the operation parameters on the CLC performance with CaSO4 as the oxygen carrier was carried out by Deng et al.31 In our previous work, we successfully simulated the process of coal gas fueled chemical-looping combustion and further studied the effects of several important operating conditions on the CLC performance.32,33 However, the fuels used were all gaseous. Mahalatkar et al.34 took the lead in studying the CLC performance of solid fuels. However, it was not feasible for continuous charging in their work. Moreover, the 3D

2. MODELS DESCRIPTION AND SIMULATION METHOD As shown in Figure 2, the fuel reactor used in this study was selected as a circulating fluidized bed riser with a height of 10 m and an internal

Figure 2. Sketch of the CFB fuel reactor. diameter of 76 mm. A detailed description of the experimental system can be found elsewhere.35,36 In order to eliminate the impact of the intense nonlinear character of the model and make a rigorous analysis, the comprehensive models are simplified as follows: (1) Gas flows into the inlets with a uniform velocity distribution. 2174


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constants C1 = 1.44 and C2 = 1.92. The turbulent Prandtl numbers for κ and ε are σκ = 1.0 and σε = 1.3, respectively. The momentum equation of particle phase is

(2) The particles in the simulations are spherical, inelastic, and uniform in size. (3) Small interaction forces such as Brownian force, lift force, thermophoretic force, and virtual mass force are neglected. 2.1. Gas−particle Hydrodynamics. 2.1.1. Continuity Equations. The continuity equations for gas and particle phases are given by37

∂ (αpρp vp⃗ ) + ∇·(αpρp vp⃗ vp⃗ ) ∂t = − αp∇pp + ∇·τp + αpρp g ⃗ + ṁ pg vp⃗ + β(vg⃗ − vp⃗ )

∂ (αgρg ) + ∇· (αgρg vg⃗ ) = ṁ gp ∂t

(1)

∂ (αpρp ) + ∇·(αpρp vp⃗ ) = ṁ pg ∂t

(2)

where α, ρ, and v⃗ are the volume fraction, the density, and instantaneous velocity, respectively. ṁ on the right-hand side is source term and it characterizes the mass transfer between the phase and particle phase. 2.1.2. Momentum Equations. The momentum equation of phase is

(12)

where, τ p is the particle phase stress−strain tensor and pp is the particle phase pressure due to particle collisions. They are expressed as37

the the gas gas

⎛ 2 ⎞ τp = αpμp (∇vp⃗ + ∇vp⃗ T ) + αp⎜ξp − μp ⎟∇· vp⃗ I ⎝ 3 ⎠

(13)

pp = αpρp Θ + 2g0αp2ρp Θ(1 + e)

(14)

In eq 13, ξp is the particle phase bulk viscosity and μp is the particle phase shear viscosity. In eq 14, g0 is the radial distribution function. They have the following form:

∂ (αgρg vg⃗ ) + ∇·(αgρg vg⃗ vg⃗ ) ∂t (3)

ξp =

4 Θ αpρ d pg (1 + e) 3 p 0 π

where g⃗ is gravity, τ g is the gas phase stress−strain tensor, and β is the interphase momentum exchange coefficient between the gas phase and particle phase, which can be written in the following form:38

μp =

10ρp d p π Θ ⎡ ⎤2 4 4 Θ αpρp d pg0(1 + e) + ⎢⎣1 + αpg0(1 + e)⎥⎦ 5 5 96αpg0(1 + e) π

= − αg∇pg + ∇·τg + αgρg g ⃗ + ṁ gpvp⃗ + β(vp⃗ − vg⃗ )

⎧ αpαgρg |vp⃗ − vg⃗ | ⎪ 0.75C D αg −2.65 (αg > 0.8) ⎪ dp ⎪ β=⎨ ⎪ αp2μg ρg αp|vp⃗ − vg⃗ | ⎪150 (αg ≤ 0.8) + 1.75 2 ⎪ dp αgd p ⎩

(16) −1 ⎡ ⎛ αp ⎞1/3⎤ ⎥ ⎢ ⎟⎟ g0 = ⎢1 − ⎜⎜ ⎥ ⎝ αp,max ⎠ ⎦ ⎣

(4)

Rep =

24 [1 + 0.15(αgRep)0.687 ] αgRep

∂ (αpρp Θ) + ∇·(αpρp vp⃗ Θ) ∂t 2 = [(− pp I + τp): ∇vp⃗ + ∇·(ΓΘ∇Θ) − γp + ϕp] 3

(5)

ρg αg|vp⃗ − vg⃗ |d p μg

(6) 2 αgμ ∇· vg⃗ I 3 g

(7)

γp =

where

μg = μgl + μgt

κ2 ε

dp π

ρp αp2 Θ3/2

ϕp = − 3β Θ ΓΘ =

(19) (20)

150ρp d p π Θ ⎡ ⎤2 6 Θ 2 ⎢⎣1 + (1 + e)g0αp⎥⎦ + 2ρp αp d pg0(1 + e) 5 384(1 + e)g0 π

(21)

(9)

2.1.3. Energy Conservation Equations. To describe the conservation of energy in Eulerian multiphase, a separate enthalpy equation is written for each phase. Heat transfers in each phase and the heat exchange between two phases are taken into account, but the pressure work, kinetic terms, and viscous heating are negligible:

where Cμ is a constant, which is set as 0.09. The governing transport equations for κ and ε are ⎛ μgt ⎞ ∂ (αgρg κ ) + ∇·(αgρg vg⃗ κ ) = ∇·⎜αg ∇κ ⎟ + αgGκ − αgρg ε + αgρg Πκ ∂t ⎝ σκ ⎠

∂ (αgρg Hg) + ∇·(αgρg vg⃗ Hg) ∂t

(10) ∂ (αgρg ε) + ∇·(αgρg vg⃗ ε) ∂t ⎛ μgt ⎞ ε = ∇·⎜αg ∇ε⎟ + αg (C1Gκ − C 2ρg ε) + αgρg Πε κ ⎝ σε ⎠

12(1 − e2)g0

(8)

In eq 8, μg is the gas phase shear viscosity, μgl is the gas-phase laminar viscosity, and μgt is the gas-phase turbulent viscosity, which is computed as a function of κ and ε:

μgt = ρg Cμ

(18)

where γp, ΓΘ, and ϕp represent the collisional dissipation of energy, the diffusion coefficient, and the energy exchange between the gas phase and particle phase, respectively. They are given by

and τg = αgμg [∇vg⃗ + (∇vg⃗ )T ] −

(17)

where, Θ is the granular temperature, which is proportional to the kinetic energy of the random motion of the particles. The transport equation derived from kinetic theory takes the form

where

CD =

(15)

= ∇·(λg ∇Tg) + hgp(Tg − Tp) + ṁ gpHp

(22)

∂ (αpρp Hp) + ∇·(αpρp vp⃗ Hp) ∂t

(11)

where Πκ and Πε represent the influence of the particle phase on the gas phase, and Gκ is the production of turbulent kinetic energy. The

= ∇·(λp∇Tp) + hpg (Tp − Tg) + ṁ pg Hp 2175

(23)


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where H, λ, and h (hgp = hpg) are the specific enthalpy, the mixture thermal conductivity, and the transfer coefficient between phases, respectively. The third term on the right-hand is the heat transfer from the solid phase to gas phase. The heat transfer coefficient, which is related to the solid phase Nusselt number, Nup, is given by

hpg = and

⎛ −E ⎞ a(b) ⎟⎟ ka(b) = A a(b) exp⎜⎜ RT ⎝ p ⎠

The parameters in the model are selected as follows: Aa = 2 × 105 s−1, Ab = 1.3 × 107 s−1, Ea = 104.6 kJ/kmol, Eb = 167.4 kJ/kmol, Ya = 0.3, and Yb = 1.0. 2.3.2. Char Gasification. In the gasification process, the most common gasification agents are steam and CO2. Through the gasification, H2 and CO are generated as for the following reactions R2 and R3.

6λg αpαgNu p d p2

(24)

39

Nu p = (7 − 10αg + 5αg 2)(1 + 0.7Rep0.2 Pr1/3) + (1.33 − 2.4αg + 1.2αg 2)Rep0.7 Pr1/3

(25)

(27)

where Sct is the turbulent Schmidt number, which is set as 0.7. Di,m is the diffusion coefficient for species i in the mixtures. 2.3. Chemical Reactions. In the fuel reactor, the chemical process mainly contains the following reactions: (1) coal pyrolysis and devolatilization, (2) char gasification, (3) water gas shift (WGS) reaction, and (4) ilmenite reduction reaction. 2.3.1. Coal Pyrolysis Reaction Kinetics. Coal pyrolysis and composition balance of coal are considered as follows:

CH4

H2O

H2

100% H2O 50% H2O + 50% H2O 100% CO2

5.7 23.1 46.5

42.8 12.7 −29.1

7.1 8.8 9.9

−27.8 −14.7 4.8

5.5 3.4 1.2

(R4)

Fe2TiO5(s) + TiO2 (s) + H 2(g) → 2FeTiO3(s) + H 2O(g) 3Fe2O3(s) + H 2(g) → 2Fe3O4 (s) + H 2O(g)

(R5) (R6)

for reduction with CO

Fe2TiO5(s) + TiO2 (s) + CO(g) → 2FeTiO3(s) + CO2 (g) 3Fe2O3(s) + CO(g) → 2Fe3O4 (s) + CO2 (g)

(R7) (R8)

for reduction with CH4

4Fe2TiO5(s) + 4TiO2 (s) + CH4(g) → 8FeTiO3(s) + CO2 (g) + 2H 2O(g)

released devolatilized species for different fluidizing gas composition is because that there is some CH4 re-forming taking place with H2O and CO2. Also note that H2O or CO2 have negative values in some cases because some H2O and CO2 were taken from the gasification agent flow for the partial CH4 re-forming.24 The well-known two-equation method for coal devolatilization is adopted in this work.40 The reaction rate is given by the expression

rpyroly = (Yaka + Ybk b)Ccoal

(30)

for reduction with H2

Table 1. Mass (g) of Different Gaseous Species Generated from the Release of Volatile Matter of 100 g of Colombian Coal after CH4 Re-Forming, for Different H2O−CO2 Mixtures as Gasification Agent CO2

k1preact 1 + k 2preact + k 3pprod

2.3.4. Ilmenite Reduction Reaction. During the ilmenite reduction, the oxygen transferred to the intermediate gasification products is the sum of the contribution of the reduction of Fe2TiO5 and Fe2O3 (reactions R5−R10):

As shown in eq R1, the volatile matter in coal was released mainly in the form of H2, CO, CH4, and CO2 as gaseous species. Other small amounts of light hydrocarbons and tars were neglected in this paper. Coefficients a−e depend on the distribution of gases generated during the devolatilization in steam−CO2 mixtures, whose correlation used in this model are described in Table 1. The difference between the

CO

(R3)

CO(g) + H 2O(g) → CO2 (g) + H 2(g)

(R1)

gasification agent

C(s) + CO2 (g) → 2CO(g)

where preact is the partial pressure of the gaseous reactants, i.e. H2O or CO2, pprod is the partial pressure of the gasification products, i.e. H2 or CO, and k1, k2, and k3 are the kinetic constants. Table 2 lists the gasification kinetic constants in case of using H2O or CO2 as gasification agent.24 2.3.3. Water−Gas Shift (WGS) Reaction. Reaction R4 corresponds to the water−gas shift (WGS) reaction. The fulfillment of it could modify the final concentrations of CO, H2, H2O, and CO2. However, previous CLC experiments performed with solid fuels show that the WGS reaction has no important influence on the gas product composition.8,24 Hence the WGS reaction was neglected in this paper.

where Jg,i is the diffusion flux of species i in the gas phase, and Si is the net rate of production of species i. Diffusion flux Jg,i is calculated by

coal → aCO + bH 2 + cCH4 + dCO2 + e H 2O + f char

(R2)

rgasif =

(26)

μgt ⎞ ⎛ Jg, i = − ⎜ρg Di ,m + ⎟∇Yg, i Sct ⎠ ⎝

C(s) + H 2O(g) → CO(g) + H 2(g)

It was assumed that char gasification is controlled by a chemical reaction.24 The gasification rate rgasif is defined by eq 30.

2.2. Species Transport Equations. Gas phase is assumed to be a mixture of five species, represented as follows: CO2, H2, CO, CH4, and H2O. The transport equations for species take the general form of eq 26 for these chemical species but the H2O, which are computed from the fact that the sum of all mass fractions is equal to one in the gas phase ∂ (αgρg Yg, i) + ∇·(αgρg Yg, ivg⃗ ) = −∇·αgJg, i + Si ∂t

(29) 40

(R9)

12Fe2O3(s) + CH4(g) → 8Fe3O4 (s) + CO2 (g) + 2H 2O(g) (R10) In the air reactor, the oxygen carrier is regenerated by taking up oxygen from the air:

(28)

where Ccoal is the concentration of unreacted coal in the solid particles. ka and kb are rate constants with Arrhenius form

4FeTiO3(s) + O2 (g) → 2Fe2TiO5(s) + 2TiO(s)

(R11)

4Fe3O4 (s) + O2 (g) → 6Fe2O3(s)

(R12)

In this study, the IRoR (integrated rate of reduction) model was used, by which the total rate of reduction of a whole particle is adopted 2176


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Table 2. Gasification Kinetic Constants for Char from Pretreated Colombian Coal k1,H2O (s−1·bar−1) k0 Ea (kJ/mol)

K2,H2O (s−1·bar−1)

K3,H2O (s−1·bar−1)

−6

−9

2.81 × 10 −135.1

52.6 95.1

8.1 × 10 −218.5

(R13)

Fe2O3(s) + CO(g) → 2FeO(s) + CO2 (g)

(R14)

(R16)

(31)

where τ is the time for complete conversion of oxygen carrier for the reduction reaction and is calculated from eq 32: ρm rg τ= b ̅ kCg n (32) C, n, b̅, rg and ρm in eq 32 represent the concentration of gas reactant, reaction order, stoichiometric factor, mean radius of the grains, and molar density, respectively. The rate constant k follows Arrhenius-type dependence with temperature (33)

where k0 is the preexponential factor of the rate constant, E is the activation energy, and R is the constant of the ideal gases, R = 8.314 J·mol−1·K−1. The reaction rates of R13−R15 are expressed as follows:

⎛ ρm αpYFe2O3 dX ⎞ (− r )j = ⎜ ⎟ (mol · m−3· s−1) b̅ dt ⎠ j ⎝

K3,CO2 (s−1·bar−1) 1.84 × 10−6 −157.6

(36)

where m is the instantaneous mass of oxygen carrier, mox is its initial mass when fully oxidized, and mred is its mass when reduced completely. 2.5. Numerical Considerations. In this study, the commercial CFD software code Fluent 6.3 was employed. Simulations were performed in three-dimensional domains. To obtain more structured grid cells, geometrical systems were segmented into several relatively regular parts. The hexahedral grids were applied in most zones and tetrahedral grids were applied near the inlets. After mesh-independent analysis, the meshes composed of 143 067 control volumes were adopted. The governing equations mentioned above were solved using a finite volume method. A first-order upwind discretization was chosen for all solutions. The phase-coupled SIMPLE (PCSIMPLE) algorithm, which is an extension of the SIMPLE algorithm to multiphase flows, was used to solve the pressure−velocity coupling and correction. The time step was set as 1 × 10−3 s. At the inlets, the velocity inlet condition was selected. Both the velocities and the concentrations of the gas and solid phases were specified, according to the superficial gas velocities and solids mass fluxes. At the outlet, the outflow condition was adopted. The no-slip wall condition was applied for the gas phase and solid phase. Initially, the solids concentration in the reactor was zero. The velocities of gas phase and solid phase were both set to zero. Table 4 lists the main parameters applied to the simulation.

In this work, the particle was assumed to be composed of a number of nonporous spherical grains with uniform initial radius, rg. The grain model with uniform reaction in the particle with changing grain size model in the grains, controlled by chemical reaction, was used to determine the kinetic parameters. The equations that describe this model are as follows:6,24

k = k 0 exp(− E /RT )

3.28 × 10 −158.5

−7

where YC is the mass fraction of char in the original coal, QCoal,in is the mass flux of coal at the solids inlet, and QC,out is the mass flux of unreacted char at the outlet. 2.4.3. Conversion of Oxygen Carrier. The conversion of oxygen carrier was defined as mox − m X= mox − mred (37)

(R15)

t = 1 − (1 − Xs)1/3 τ

K2,CO2 (s−1·bar−1)

⎛ Q C,out ⎞ ⎟ × 100% XC = ⎜⎜1 − Q Coal,inYC ⎟⎠ ⎝

4Fe2O3(s) + CH4(g) → 8FeO(s) + CO2 (g) + 2H 2O(g)

4FeO(s) + O2 (g) → 2Fe2O3(s)

4.53 × 10 160.1

3

2.4.2. Conversion of Carbon in Char. The conversion of carbon in char (XC) is calculated as

instead of the local rate of reduction (LRoR). This model has been widely used for the analysis of the reduction of iron ores.5,41 In the IRoR model, the Fe2TiO5 was considered to be a mixture of Fe2O3 and TiO2, FeTiO3 was a mixture of FeO and TiO2, and Fe3O4 was a mixture of Fe2O3 and FeO. TiO2 was considered to be an inert material. Thus, the reduction and oxidation sequence of ilmenite (reactions R5−R12) can be summarized and simply expressed as the following reactions (R13−R16):

Fe2O3(s) + H 2(g) → 2FeO(s) + H 2O(g)

k1,CO2 (s−1·bar−1)

Table 4. Initial Conditions and Model Parameters (34)

where j represents the jth reaction. Table 3 includes the detailed kinetic parameters.24

Table 3. Kinetic Parameters for Ilmenite Reduction with H2, CO, and CH4 gas reactants

b̅

k0 (mol1‑n·m3n‑2·s−1)

E (kJ/mol)

n

H2 CO CH4

1.45 1.45 5.78

6.2 × 10−2 1.0 × 10−1 9.8

65 80.7 135

1 0.8 1

description

value

reactor diameter (mm) reactor height (m) initial/maximum solids packing inlet boundary conditions outlet boundary conditions time step size (s) convergence criteria

76 10 0/0.64 velocity inlet outflow 10−3 10−4

The bed material in this study was Norwegian ilmenite.5,24 Because the active Fe2TiO5 was considered to be a mixture of Fe2O3 and TiO2, the total mass fraction of Fe2O3 in the ilmenite was considered to be about 47.7%. The fuel used was a bituminous Colombian coal.24 The mean particle diameter was 150 μm. Tables 5 and 6 list the main parameters of the used ilmenite and coal.

2.4. Data Evaluation. 2.4.1. CO2 Concentration (dry basis). The CO2 concentration (dry basis) ( f CO2) in the flue gas could be calculated as xCO2 fCO = 2 xCO + xCO2 + xCH4 + x H2 (35)

3. RESULTS AND DISCUSSION This work mainly devotes to the research on the gas−solid flow behaviors and reaction performance in the CFB fuel reactor for

where xi is the molar fraction of species i in the gas phase. 2177


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Table 5. Composition (wt %) and Physical Properties of Activated Ilmenite Norwegian ilmenite main species true density (kg/m3) grain radius, rg (μm) porosity (%)

Fe2TiO5 (38.5%), Fe2O3 (22.0%), TiO2 (34.0%), inerts (5.5%) 4250 1.25 12.7

Table 6. Proximate and Ultimate Analyses and Lower Heating Value of Pre-Treated Colombian Coal proximate analysis (ad,a wt %) ultimate analysis (ad,a wt %) lower heating value (kJ/kg) a

moisture 2.3 carbon 65.8 21899

volatile 33.0 hydrogen 3.3

fixed carbon 55.9 nitrogen 1.6

ash 8.8 oxygen 17.6

sulfur 0.6

Figure 3. Comparisons of Solids Holdup Distributions between Experimental Data and Predictions.

indicates that the model with Gidaspow drag correction is suitable for the simulation of gas−solid flow in this CFB riser. In order to verify the reaction model, a comparison of reaction performance was carried out beforehand between the calculations and hot experimental data on a BFB fuel reactor.7 Figure 4 displays the comparison of gas concentration

ad: air-dried basis.

coal-fired CLC. Table 7 lists the main operating parameters of the reference condition. Table 7. Main Operating Parameters of the Reference Condition parameters

value

coal feed rate (kg/h) steam supply (kg/h) solid flux (kg/m2·s) reactor temperature (K) pressure (MPa)

3 20 300 1243 0.5

In order to further optimize the operating conditions for direct coal-fired CLC, different cases have been applied to study the effects of operating variables on the one-pass conversion of carbon in char through the fuel reactor. The higher the onepass conversion of carbon in the fuel reactor, the higher the thermal power that could be achieved. Then for investigating the effect of one parameter more rationally, other parameters should be kept constant. Table 8 lists all the operating

Figure 4. Comparisons of gas concentrations at the outlet between experimental data and predictions.

distributions at the fuel reactor outlet. The simulation conditions were set consistent with the experimental operation conditions. Under one of their operating conditions, i.e., 83 g/h for the coal feeding flow, 190 LN/h for the steam flow, and 890 °C for the temperature, the dry basis concentrations of gas species at the outlet were about 52.8% for CO2, 26.5% For H2, 19.1% for CO, and 1.6% for CH4. As shown in Figure 4, the corresponding gas concentrations obtained in our simulation were 55.2% for CO2, 21.2% For H2, 21.9% for CO, and 1.7% for CH4. It could be observed that the predictions were very close to the experiment data, and the maximum relative error between the simulation and experiment is less than 25%, implying that the reaction simulations are reasonable and the validity of reaction models is verified. 3.2. Flow Characteristics. Initially, there were not particles in the reactor. At the beginning of the simulation, the gasification agent (steam) was fed into the bed through a distributor at the bottom of the reactor, and the particles (mixture of coal and oxygen carrier) were transported by steam into the solids inlet.

Table 8. Range of operating conditions for the simulations case

temp (K)

H2O:mixture (H2O + CO2)

1−3 4−6

1193−1243−1273 1243

1 0−0.5−1

conditions used in the simulations, where cases 2 and 6 are the same for the reference condition. 3.1. Validity of Flow Characteristics and Reaction Models. Before this comprehensive modeling research, a validation of flow characteristic was performed through the comparison with the experimental data.35 Figure 3 shows the comparison of time-averaged solids holdup distributions in different radial regions along the riser between experimental data and predictions. The experimental operation conditions were described in detail35 and the simulation conditions were set accordingly. The calculation results of solids holdups show a good agreement with the experimental data, also with the simulation results obtained from other researches.36 This 2178


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Figure 5 shows the axial and radial distributions of solids holdup at 70 s in the fuel reactor. It could be found from Figure 5a

Figure 6. Radial profiles of predicted particle velocities at z = 1 and 7 m.

could also observe that the solids velocity becomes larger in the higher section of the reactor due to the drive of gas. These observations agree well with the cold experimental data from Jin et al.43 3.3. Spatial and Temporal Distributions of Gas−Solid Compositions. The patterns of char mass fraction in the solid phase and concentrations of CO2 and CO in the gas phase at 70 s are presented in Figures 7−9. There are significant differences in the spatial distributions among different species. Char is generated from coal pyrolysis in abundance at the bottom; thus, the mass fraction of char reaches the maximum in this section. Afterward, it is gradually consumed through gasification reactions during the ascent. CO2 is produced from coal pyrolysis and especially the combustion reactions of intermediate gasification products (H2, CO, CH4) and oxygen carrier (Fe2O3). It could be seen that the concentration of CO2 keeps increasing with the bed height, which reaches the maximum at the outlet. Moreover, we can observe that the CO2 concentration is larger near the wall than that in the center region. It is probably because that the gas−solid reactions are promoted near the wall due to the higher solids holdups. As shown in Figure 9, the reverse is realized for the distributions of intermediate gasification products (H2, CO, CH4). Figure 10 shows the changes over time of gas components at the outlet during the quasi-equilibrium state (t > 50 s). As expected, only slight fluctuations of gas concentrations are observed at this stage. The mean dry basis concentrations of gas components are about 99.0% for CO2, 0.42% for H2, and 0.58% for CO, while CH4 is almost totally consumed. Figure 11 presents the time-averaged conversions of carbon in char and oxygen carrier. The one-pass conversion of carbon in char is about 82.5%, and the conversion of oxygen carrier is about 2.6% at the outlet. It implies that the high CO2 concentration and one-pass conversion of carbon in char are achievable adopting this kind of pressurized CFB fuel reactor. 3.4. Distributions of Heterogeneous Reaction Rates. Figures 12 and 13 show the rates of representative heterogeneous reactions R2 and R14 referred to in this CLC model. Through the comparison of the reaction rates of combustion (R14) and char gasification (R2), we could find that the combustion rates are generally limited by the char gasification rates, because the gasification is the rate-controlling step.20 But for the reason that the combustion of volatiles is also included in the combustion rate, we can observe that the combustion rates exceed the char gasification rates a little in

Figure 5. Axial and radial distributions of solids holdup.

that the solids holdup generally decreases along with the bed height. Meanwhile, the change of solids holdup is much more intense at the bottom of the riser than that in the upper region, indicating a more uniform distribution of particles in the upper part. Thus, the riser can be divided into three longitudinal sections: bottom high-density section (εs ≥ 0.1), middle transition section, and upper dilute section.35,42 But, apparently, the adoption of high-flux (Gs ≥ 200 kg/m2·s) condition greatly promotes the solids holdups in the whole reactor. From Figure 5b−d we could further observe that the solids holdup in the central region is obviously lower than that near the wall at a given bed height, meaning a gathering of particles due to the wall effect.43 The calculated results predicted a highly efficient gas−solid mixing as well as high solids holdups over the whole reactor height, which would greatly promote the complex gas−solid reactions and further the CO2 concentration and fuel conversion. Figure 6 presents the radial velocity distributions of particles in different bed heights. The particle velocity is flat in the center region and then turns smoothly downward toward the wall. This trend is opposite to the distribution of solids holdup. We 2179


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Figure 7. Spatial distributions (axial dimension in horizontal and radial position in vertical) of char mass fraction at three different sections: (a) 1.4− 1.5 m, (b) 4.5−4.6 m, and (c) 8.9−9.0 m.

Figure 8. Spatial distributions (axial dimension in horizontal and radial position in vertical) of CO2 concentration at three different sections: (a) 1.4−1.5 m, (b) 4.5−4.6 m, and (c) 8.9−9.0 m.

Figure 9. Spatial distributions (axial dimension in horizontal and radial position in vertical) of CO concentration at three different sections: (a) 1.4− 1.5 m, (b) 4.5−4.6 m, and (c) 8.9−9.0 m.

By comparing the distributions of reaction rates (Figures 12 and 13) and species (Figures 7 and 8), it can be found that there is a close relationship between the reaction rates and

some places. Thus, the realization of pressurized operation accelerates not only the gasification reaction but also the reactions of the whole system. 2180


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3.5. Influence of the Fuel Reactor Temperature. The temperature in the fuel reactor was evaluated as one of the main operating variables affecting the performance of the CLC system.26,44,45 Cases 1−3 are designed to study the influence of fuel reactor temperature on the one-pass conversion of carbon in char through the fuel reactor. The temperatures tested are in the range of 1193−1273 K. Figure 14 presents the comparison of one-pass conversion of carbon in char at the quasi-equilibrium state with different fuel reactor temperatures. As expected, the one-pass conversion of carbon is about 70.8% for the temperature of 1193 K and increases to 87.8% for the temperature of 1273 K. This is because both gasification and oxidation reactions are accelerated at higher temperature. Meanwhile, the accelerated consumption of reducing gases (i.e., CO, H2, and CH4) by oxidation reactions will further promote the gasification reactions of char. 3.6. Influence of the Gasification Agent Type. The motivation of using CO2 as a gasification agent is that the recirculation of part of the CO2 obtained at the fuel reactor outlet will lead to savings of energy needed for steam generation and therefore increase the efficiency of the whole process.24 Cases 4−6 are chosen to study the influence of gasification agent type on the one-pass conversion of carbon in char through the fuel reactor. The steam fraction in the gasification agent ranges from 0 to 1. Figure 15 exhibits the effect of steam fraction in the gasification agent on one-pass conversion of carbon in char. It could be seen that the conversion of carbon in char is enhanced when there is higher steam fraction in the gasification. The conversion of carbon in char increases from 57.4% to 82.5% when the steam fraction increases from 0 to 1. This is mainly because the gasification rate by steam is faster than gasification by CO2 for this fuel, which has been assessed during kinetics determination. Thus, when CO2 is recirculated to the fuel reactor as a gasification agent, the thermal power will be decreased due to the reduction of conversion of carbon in char while the energy efficiency of the whole system will be increased. Hence it is necessary to find a balance between the energy efficiency of the whole system and the gasification rates in the actual design and operation process. 3.7. Performance Comparison with Typical SpoutFluid Bed Fuel Reactor. In order to exhibit the potential of CFB fuel reactor, we carried out a comparison of reaction

Figure 10. Variation of gas components (dry basis) at the outlet during the quasi-equilibrium state.

Figure 11. Conversions of carbon in char and Fe2O3.

distributions of species. For the char gasification R2, the reaction rate gradually decreases along with the bed height due to the consumption of char. For the combustion of gasification products R14, the reactions are more intense at the bottom of the reactor because of the higher concentrations of intermediate gasification products (H2, CO, CH4) in this zone. Meanwhile, at the bottom of the reactor, the reaction rates are increased from the center toward the wall due to the higher solids holdup near the wall.

Figure 12. Spatial distributions (axial dimension in horizontal and radial position in vertical) of reaction rate for R2 at three different sections: (a) 1.4−1.5 m, (b) 4.5−4.6 m, and (c) 8.9−9.0 m. 2181


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Figure 13. Spatial distributions (axial dimension in horizontal and radial position in vertical) of reaction rate for R14 at three different sections: (a) 1.4−1.5 m, (b) 4.5−4.6 m, and (c) 8.9−9.0 m.

Table 9. Comparison of Reaction Performance between the CFB Riser and Typical Spout-Fluid Bed

Figure 14. Effect of temperature on one-pass conversion of carbon in char through the fuel reactor.

description

Shen et al.’s experiment

this work

fuel reactor type material of oxygen carrier gasification agent type coal flow (kg/h) solids mass in fuel reactor (kg) solids inventory (kg/MWth) temperature (K) dry-basis concn of CO2 at the outlet (%) carbon conversion efficiency (%)

spout-fluid bed NiO steam/CO2 1.2 11 ∼1300 1243 95.2 92.8

CFB riser ilmenite steam 3 20 ∼1100 1243 99.0 ∼82.5 (one-pass)

Figure 15. Effect of gasification agent type on one-pass carbon conversion in char.

First, compared with expensive NiO-based oxygen carrier used in Shen et al.’s experiment, ilmenite adopted in this work is much cheaper and is more suitable for the future application in CLC power plants. Then, as expected, the dry-basis CO2 concentration at the fuel reactor outlet in Shen et al.’s experiment can only achieve 95.2% due to the gas bypassing, although highly reactive oxygen carrier is used. However, due to a favorable gas−solid contact over the whole reactor height, the CO2 concentration could reach 99% in this CFB reactor. Finally, the maximum carbon conversion efficiency in Shen et al.’s experiment is only 92.8%, indicating that the carbon loss is inevitable due to an easy elutriation of fine char particles from the freeboard of the spout-fluid bed. But for the CFB fuel reactor, high conversion of carbon in char could be achieved with the recirculation. At the same time, as shown in Table 9, the one-pass carbon conversion efficiency through the fuel reactor has achieved about 82.5%, meaning the thermal power of system could also be assured. Overall, this comparison shows the potential and competitiveness of CFB fuel reactor in the future large-scare coal-fired CLC power plants.

performance between the CFB riser and typical spout-fluid bed.17 As shown in Table 9, the main operating conditions (fuel and gasification agent types, solids inventory, temperature) we used to simulate the process in CFB fuel reactor are very similar with Shen et al.’s spout-fluid bed experiment, except that the NiO-based oxygen carrier they used is much more reactive than the ilmenite adopted in this work.

4. CONCLUSIONS A comprehensive three-dimensional numerical model has been developed to simulate the coal-fired chemical looping combustion process in a pressurized CFB fuel reactor. The model consists of complex gas−solid flow and chemical reactions. The results show reasonable solids holdups, velocities of solid phase, species distributions, fuel conversion, and 2182


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Nu p Pr Q r rg R Re ṁ Sc Si v⃗ XC XOC

reaction rates in the fuel reactor. The main innovations and conclusion points from these studies are listed as follows: (1) In order to overcome the chronic shortcomings of typical BFB fuel reactor, i.e., gas bypassing in the fluidized region and inadequate gas−solid contact in the freeboard, we developed a CFB fuel reactor model for coal-fired CLC. Therefore, the inherent feature of CFB of sufficient gas− solid contact over the whole reactor height ensured that the gas−solid reactions were performed in a homogeneous and favorable system. (2) We innovatively grafted the technologies of high-flux CFB into the CLC application. Thus, the solids holdups in the whole reactor were greatly increased, and further, the gas−solid reaction performance was enhanced. (3) The pressurized condition was successfully adopted to increase the gasification reaction. This is because the coal gasification is the rate-controlling step in the coal-fired CLC process. Moreover, in a pressurized CLC system, only a little additional power is required for the further compression of CO2 as a high-pressure gas. (4) By giving consideration to the effects of various operation conditions comprehensively, one operating condition with 1243 K for temperature, 300 kg/m2·s for solids flux, 0.5 MPa for pressure, and 0.5 for the steam ratio in the gasification agent was recommended as the optimum. Under this condition, the CO2 concentration at the outlet could reach over 99%, and the one-pass carbon conversion in char could achieve more than 75% while energy saving and avoiding of sintering could also be guaranteed due to the recirculation and process compression of CO2 and the relatively muted operating temperature.

■

Greek Letters

α β γ ε Θ κ λ μ ξp ρ σε σκ ∏ε Πκ τ ΓΘ ϕp

AUTHOR INFORMATION

Corresponding Author

*Tel.: +86-25-83794744. Fax: +86-25-83795508. E-mail: bsjin@ seu.edu.cn.

g i l p t

The authors declare no competing financial interest.

■

ACKNOWLEDGMENTS Financial support from the National Natural Science Foundation of China (51076029), the Collaboration Project of China and British (2010DFA61960), the Scientific Research Foundation of Graduate School of Southeast University (YBJJ1119), the National Natural Science Foundation of China (50706007), and the Foundation of Excellent Young Scholar of Southeast University (4003001039). C CD dp D E g0 g⃗ Gk h H Ji k k0

volume fraction drag (kg/m3·s) collisional dissipation of energy (W/m3) dissipation rate of turbulent kinetic energy (m−2·s−3) particle phase pseudotemperature (m2/s2) turbulent kinetic energy (m2/s2) thermal conductivity (W/m2·K) viscosity (kg/m·s) particle phase bulk viscosity (kg/m·s) density (kg/m3) turbulent Prandtl number for ε turbulent Prandtl number for κ influence of the particle phase on the gas phase (m−2·s−4) influence of the particle phase on the gas phase (m2/s3) stress−strain tensor (Pa) the diffusion coefficient between the gas phase and particle phase energy exchange between the gas phase and particle phase (W/m3)

Subscripts

Notes

■

Nusselt number pressure (Pa) Prandtl number mass flux (kg/h) reaction rate (mol/m3·s) radius of a grain (m) constant of the ideal gases (J/mol·K) Reynolds number mass source term (kg/m3·s) Schmidt number net rate of production of species i (kg/m3·s) instantaneous velocity (m/s) conversion of carbon in char conversion of oxygen carrier

■

gas phase the ith specie laminar flow particle phase turbulent flow

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NOMENCLATURE gas concentration (mol/m3) drag coefficient granular diameter (m) diffusion coefficient (m2/s) activation energy (J/mol) radial distribution function gravity (m/s2) production of turbulent kinetic energy (kg/m·s3) heat transfer coefficient (W/m2·K) specific enthalpy (J/kg) diffusion flux of species i (kg/m2·s) chemical reaction rate constant (mol1‑n·m3n‑2·s−1) preexponential factor of the rate constant 2183


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