Article pubs.acs.org/IECR
Catalytic NOx Reduction in a Novel i‑CFB Reactor: I. Kinetics Development and Modeling of Reduction Zone Xingxing Cheng†,‡ and Xiaotao T. Bi*,‡ †
National Engineering Laboratory of Coal-Fired Pollutants Emission Reduction, School of Energy & Power Engineering, Shandong University, Jinan, China 250061 ‡ Fluidization Research Centre, Department of Chemical and Biological Engineering, The University of British Columbia, Vancouver, Canada V6T 2K9 ABSTRACT: A model was developed for the hydrocarbon selective catalytic reduction (HC-SCR) of NOx in a fluidized bed reactor. For the kinetics, the reaction was divided into two submodels: (i) NO oxidation and adsorption; (ii) hydrocarbon oxidation and NOx reduction. The kinetic parameters were obtained from fitting a fixed bed model to fixed bed experimental data. Mass transfer between solids and gas phases was also considered in this fixed bed model with the reaction term being embedded into the solids phase mass balance equation. Comparing with fixed bed experimental data, the model showed an average error of about 7% for NOx conversion and about 5% for HC conversion. The fitted kinetics was then incorporated into a fluidized bed reactor model, which included gas flow in the bubble phase, gas flow in the dense phase and solids flow in the dense phase. The fluidized bed reactor model was compared to the fluidized bed experimental data. The model showed acceptable agreement with measured NOx conversion, but poor agreement with observed HC conversion data. Suggestions were proposed to further improve the model and experimental data. The fluidized bed model was then applied to simulate NOx reduction in the reduction zone of the internal circulating fluidized bed (i-CFB) reactor. The simulation results showed that NOx conversion could be improved if NOx is fed into the reactor via the solids phase, the same as what happens in the reduction zone of an i-CFB reactor. Also, for the reduction zone of the i-CFB reactor, a higher solids circulation rate is preferred if NOx feed rate is kept at a constant. These simulations could provide good suggestions for the i-CFB design and operation.
1. INTRODUCTION Nitrogen oxides (NOx) are emitted primarily from transportation and other industrial sources, which contribute to many problems that threaten the quality of the environment and human health: the formation of acid rain and ground level ozone (smog), and general atmospheric visibility degradation.1 Due to the increasingly stringent emission regulations on NOx all over the world, the method for the abatement of NOx emissions has gained extensive attention from academic as well as industrial research groups. As a technology under development, selective catalyst reduction by hydrocarbon (HC-SCR), which uses hydrocarbons as the reducing agent, shows satisfactory performance in deNOx reactions,2 although the conversion efficiency is significantly affected by the oxygen concentration present in the flue gases.3−6 To avoid the negative impact of excess O2 and other poisoning components in the flue gas, Yang and Bi7 proposed a new integrated NOx adsorption-reduction process, where the NOx adsorption and reduction are carried out in two separate zones of the reactor. The flue gas is passed into the adsorption zone where NOx is absorbed by the catalyst. The NOxadsorbed catalyst particles then move into the reaction zone where NOx is reduced by injected hydrocarbons at controlled oxygen concentrations, and at the same time, other adsorbed flue gas contaminants are also stripped from the catalyst. The regenerated catalyst particles are then recirculated back to the adsorption zone to establish a continuous operation. The concept has then been validated in a novel internal circulating fluidized bed reactor (i-CFB). © 2014 American Chemical Society
In order to evaluate and further improve the performance of the i-CFB reactor, a reactor model needs to be developed. Different from the conventional fluidized bed catalytic reactor where adsorption and reaction take place at the same time, adsorption and reduction take place in two different zones of the i-CFB reactor. Adsorption and reaction therefore should be decoupled in the kinetics and modeled separately for the adsorption and reduction zones separately and the performance of the adsorption and reduction zones is then linked by the solid circulation between the two zones. In catalytic reaction kinetic models, surface reaction and surface adsorption are usually embedded into the gas phase mass balance equation, even with NOx storage being considered.8,9 A few other models consider adsorption and reduction separately. In the NSR (NOx storage reduction) reaction, NOx is first adsorbed onto the catalyst surface in the lean cycle and then is reduced in the rich cycle. Lindholm10 modeled NOx storage and reduction with hydrogen as the reducing agent over Pt/Ba/Al catalyst in a monolith reactor. The mass transport to the catalyst is modeled for adsorption, desorption, and reaction on the catalyst surface in stages, separately, over a lean and rich cycle. Reaction takes place on the catalyst surface between adsorbed NOx and reduction agent during the rich cycle. The same model was also used for the modeling of NOx adsorption Received: Revised: Accepted: Published: 9365
March 17, 2014 May 7, 2014 May 9, 2014 May 9, 2014 dx.doi.org/10.1021/ie5011228 | Ind. Eng. Chem. Res. 2014, 53, 9365−9376
Industrial & Engineering Chemistry Research
Article
later be used for the evaluation of fluidized bed reactor model. The experimental conditions of both the fixed and fluidized beds were summarized in Table 1.
and NO oxidation over Cu/ZSM-5 monolith catalyst.11 Metkar et al.12 recently developed a kinetic model for NH3−SCR over Fe- and Cu-zeolite monolith catalyst in which gas diffusion in the washcoat was considered. It was also reported12 that diffusion is very important in the SCR reactor and should be included in the reactor model. These models could work well with the monolith reactors where there is a distinct boundary between the gas phase and washcoat and reactants diffusing through the layer of the washcoat. However, gas flow and mass transfer in the packed bed could be different, where the gas species diffuses into the pores of each particle. Also, model developed for monolith reactors could not be extended to fluidized bed reactors, where both diffusion inside the particles and the mixing of the solids need to be considered. In our earlier work, a reaction kinetic model was proposed for the HC-SCR reaction on Fe/ZSM-5.13 Adsorption and reduction were decoupled into two submodels, with the parameters being fitted by adsorption and reduction experimental data separately, so that the model can be applied for modeling the adsorption zone and reduction zone of the i-CFB reactor. However, because both interphase and intraphase mass transfer between gas and particles were neglected in the model, it cannot be directly applied to model fluidized bed reactors. Moreover, although adsorption and reduction are decoupled in the kinetics, only gas phase mass balance equations were considered. To model the adsorption-reduction process in both the fixed bed and fluidized bed reactors, solid phase mass balance should be considered in the adsorption model in which the reaction term should be incorporated into the solid phase to account for catalytic reaction of adsorbed NOx on the solids surface. This work extended our previous kinetics model13 by incorporating mass transfer between solid and gas phase, with the mass transfer coefficients being obtained from fitting fixed bed adsorption breakthrough curves.14 Both solid and gas phase mass balance equations were considered with the reaction term embedded into the solid phase mass balance equation. The model is then applied to simulate NOx reduction in the reduction zone of the i-CFB reactor before is extended to simulate the performance of the whole i-CFB reactor.
Table 1. Experimental Conditions expt.
NO inlet, ppm
inlet HC, ppm
inlet O2, %
gas velocity, m·s−1
fixed bed O2 adsorption fixed bed NO adsorption
N/A
N/A
0.4, 1
N/A
4
fixed bed reduction
200, 400, 600, 800, 1000 600, 800, 1000
300, 600, 1200, 2400 300, 600, 900 N/A
0.5, 1, 0.07 (GHSV = 2, 4, 8 5000 h−1) 4
0.04, 0.06, 0.08
600
1
0.2−0.6
fluidized bed NO adsorption fluidized bed reduction
600, 1200
0.07 (GHSV = 5000 h−1) 0.07 (GHSV = 5000 h−1)
3. REACTOR MODELS 3.1. Fixed Bed Model. 3.1.1. Kinetics. The kinetic model was divided into two submodels as we proposed in our previous work:13 (i) NO oxidation and adsorption, eqs R1 and R2 in Table 2, (ii) C3H6 oxidation and NOx reduction, eqs R3 and R4. [*] is denoted as the number of vacant sites on the catalyst surface, and [NO2*] is the number of sites occupied by NO2. Here, it is assumed that oxygen and NO2 are absorbed on the surface of catalyst and then react with hydrocarbon reductant that is diffused onto the catalyst surface. No hydrocarbon adsorption is considered in the kinetics. The Freundlich adsorption equilibrium submodel was applied here for adsorption of O2 and NO2, the same as the one used in our previous work.13 The equations are shown as eqs 1 and 2 in Table 2. HC oxidation and NOx reduction are expressed by power law equations, as eqs 3 and 4. The net reaction rates of hydrocarbon, adsorbed NO2 and O are then given by eqs 5 to 7. 3.1.2. Governing Equations. An axial dispersion model is developed for the fixed bed reactor, in which both solid phase and gas phase mass balances are considered. Reactants are exchanged between the gas and particles, which accounts for the adsorption and desorption process. The reaction term is embedded in the solid phase mass balance equation, representing the catalytic surface reaction. The adsorption in the fixed bed takes place in the following steps: (1) external diffusion; (2) diffusion inside the catalyst pores; (3) adsorption onto the catalyst surface. The effect of these individual steps on the breakthrough curves is difficult to be determined independently from the macroscopic experiments.15 Thus, simplified models are often used. In the film diffusion model, introduced by Whitman,16 NOx is adsorbed onto the catalyst surface and the pores of the catalyst after it diffuses through a thin film, with the porous solid being treated as a pseudohomogeneous medium where reaction takes place. All mass transfer resistances are located across the film adjacent to the particle surface as shown in Figure 1. The effective mass transfer constant of the film, kf, lumps both interphase and intraphase mass transfer resistances of the particle. NOx concentrations in the film are in equilibrium with the concentrations in the gas phase. The same method is applied for O2 adsorption and hydrocarbon transferring into the solid phase. For hydrocarbon, the concentration in the film is assumed to be the same as
2. EXPERIMENTS The catalyst used in this study is Fe/ZSM-5, which showed good performance in our previous i-CFB study. The catalyst support is H/ZSM-5 (average particle size of 155 μm, apparent bulk density of 968 kg·m−3, and BET surface area of 171 m2·g−1), which was obtained as a free sample from Albemarle Corporation. The impregnation method was used to prepare the fine Fe/ZSM-5 catalyst, which was described in detail in our earlier paper.6 The prepared fine Fe/ZSM-5 catalyst had an iron content of 5%, and a BET surface area of 163 m2·g−1, which was slightly lower than the original ZSM-5 catalyst support. For the reducing agent, propylene, C3H6, was selected due to its reported good performance over the prepared Fe/ZSM-5 catalyst. In order to obtain the adsorption and reaction characteristics of the catalyst, the catalyst was first tested in a fixed bed for its performances on O2 adsorption, NOx adsorption and NOx reduction. Details of the fixed bed experiments were shown in our previous work.13 Since the catalyst showed the best deNOx performance around 350 °C, here the characteristic of the catalyst was studied only at 350 °C. The performance of the catalyst at other temperatures will be further investigated in the future. The deNOx performance was also tested in a fluidized bed. The configuration of the fluidized bed and details of the experiment were shown in a previous work.7 The data will 9366
dx.doi.org/10.1021/ie5011228 | Ind. Eng. Chem. Res. 2014, 53, 9365−9376
Industrial & Engineering Chemistry Research
Article
Table 2. Equations for the Developed Kinetic Model kinetic reactions
k1, k−1
O2 + 2* ←⎯⎯→ 2O*
(R1)
k 2 , k−2
NO + O* ←⎯⎯⎯→ NO2 *
(R2)
k3
C3H6 + 9O* → 3CO2 + 3H 2O
(R3)
k4
2NO2 * + C3H6 + 5O* → N2 + 3CO2 + 3H 2O adsorption equilibrium
0.5[O*] = qeO2 = [NO2 *] = qeNOx =
power law kinetics
m1 r3 = k 3C HC [O*]m2
n kFO2COO2 2
3
mol/m cat
n nNO2 kFO2COO2 2kFNO2C NO
(1) n
n
NO2 ′ 2COO2 2C NO = kNO mol/m 3cat
(2)
(3)
m4 r4 = k4[NO2 *]m3 C HC [O*]m5
net reaction rates
(R4)
(4)
rNO2 * = − 2r4
(5) rHC = − r3 − r4 (6) rO * = − 9r3 − 5r4
(7)
(4) No catalyst deactivation. (5) Only O2 and NO2 are assumed to be adsorbed onto the catalyst, while the adsorption of NO and N2O are not considered. The governing equations are derived and listed as eqs 8 and 9 in Table 3. Since particles are stationary in the fixed bed, the particle velocity and solids dispersion coefficient equal zero. The solid phase equation is simplified to eq 10. In the equation, Cs* is the equilibrium concentration of adsorbed reactant in the catalyst, determined by the solid phase concentration and the adsorption equilibrium. Initial and boundary conditions are presented as eqs 11 to 13. For the reduction in the fixed bed, only steady state is assumed, which is reached when (∂C/∂t) < 1 × 10−6 mol·s−1. 3.1.3. Model Parameters. Gas mixing coefficient Dg and mass transfer coefficient kf were obtained by fitting the model to fixed bed adsorption breakthrough curves.14 The fitted values for Dg and kf are 1.2 × 10−4 m2·s−1 and 1.776 × 10−3 s−1, respectively. The fitted values were in the same order as those values calculated from empirical correlations in references.17,18 It should be noted that kf was fitted from fixed bed NOx adsorption breakthrough curves at inlet O2 = 4%, NO = 600 ppm. However, values of kf depend on the adsorption capacity of each reactant. Here, the fitted kf value is denoted as kf0. kf will be adjusted in the model for different reactants and different reaction
Figure 1. Schematic of mass transfer around particles.
in the bulk, because hydrocarbon is only transferred to the solids and is in amount more than required stoichiometrically. General assumptions for the system include (1) Since NOx concentrations were at ppm levels, the exothermic heat of adsorption was negligible and isothermal condition was assumed. (2) Pressure drop through the bed is neglected, and thus, the pressure and gas velocity are uniform throughout the reactor. Radial concentration gradient in the reactor is also neglected. (3) All mass transfer resistances are lumped onto a mass transfer film around the particle. Concentration gradients are negligible within both gas and solid phases, but not across the film. Table 3. Equations for the Fixed Bed Model governing equations
∂Cs ∂ 2C ∂C − Ds 2s − Us s ∂t ∂z ∂z = − k f (Cs − Cs*) + R i
Solid phase:
Gas phase: =
2
∂Cg ∂t
εs 1 − εs
− Dg
∂ Cg
∂z 2 k f (Cs − Cs*)
+ Ug
(8)
∂Cg ∂z
(9)
∂C Solid phase (simplified): s = − k f (Cs − Cs*) + R i ∂t
initial and boundary conditions
C = C0
z = 0,
Cg = Cg,in ,
z = H,
9367
(11)
t = 0,
∂Cg ∂z
(10)
= 0,
∂Cs =0 ∂z
(12)
∂Cs =0 ∂z
(13)
dx.doi.org/10.1021/ie5011228 | Ind. Eng. Chem. Res. 2014, 53, 9365−9376
Industrial & Engineering Chemistry Research
Article
Figure 2. Schematics of mass transfer in the fluidized bed.
Table 4. Equations for the Fixed Bed Model governing equations
Particles in the dense phase:Dsεs = εs
∂ 2Cs ∂z 2
± Usεs
εs ∂Cs − k f (Cs − Cs*) + R i 1 − εs ∂z
∂Cs ∂t
(16)
∂ 2C ∂C Gas in the dense phase:DH ΦH(1 − εs) 2H − UH ΦH(1 − εs) H ∂z ∂z − k f (Cs − Cs*) + kHLab(C L − C H) = ΦH(1 − εs)
∂C H ∂t
(17) 2
Gas in the bubble phase:DL ΦL = ΦL initial and boundary conditions
∂C L ∂t
z = 0,
C H = Cg,in ,
z = H,
∂C H = 0, ∂z
∂Cs = 0, ∂z
C H = 0,
condition by eq 14 using adsorption equilibrium constant K, in which i denotes different reactants and 0 represents the reaction condition at the fitting condition. Detailed discussion of kf is shown in the Appendix.
Cs = 0,
C L = Cg,in ∂C L =0 ∂z
CL = 0
(19) (20)
(21)
tions in the dense phase, the bubble phase, and the equilibrium concentration in the dense phase, respectively, with the same unit, mol·m−3 gas. Cs is the average concentration in the particles per unit volume with the unit, mol·m−3 cat. Cs* could be calculated from the adsorption isotherm fitted from fix bed experiments. ΦH and ΦL are the fraction of dense phase and bubble phase, respectively, with ΦL equal to bubble fraction εb and ΦH = 1 − ΦL. Parameter ab is the specific surface area between the bubble and dense phases, defined as
(14)
It is assumed that hydrocarbon is transferred from the gas phase to solids and then reacts with adsorbed NO on the catalyst surface. Therefore, different kf value is used for HC mass balance equation, as shown in eq 15.
Rp R p2 1 = + kf 3kc 15De
(18) ∂Cs = 0, ∂z
t = 0,
k f = k f0K i0/K i
∂ CL ∂C − UL ΦL L − kHLab(C L − C H) ∂z ∂z 2
ab =
6εb db
(22)
For the reduction in fluidized bed, only steady state is considered, which is assumed to be reached when (∂C/∂t) < 1 × 10−6 mol·s−1, the same as fixed bed model calculation. 3.2.2. Hydrodynamics. The total voidage consists of two parts, voidage in the dense phase and the fraction of bubbles. The hydrodynamic parameters are calculated in the same way as the model developed previously for NOx adsorption in fluidized bed using Geldart Group B powders.19 Voidage in the dense phase is assumed to be the voidage at minimum fluidization, εmf, and the gas velocity in the dense phase is assumed to be at the minimum fluidization velocity Umf. The equations for the estimation of other parameters are given in Table 5. 3.2.4. Mass Transfer. The same intraphase mass transfer coefficient kf as in the fixed bed is used for the fluidized bed
(15)
Value of kc is calculated from the correlation given in the Appendix. 3.2. Fluidized Bed Model. 3.2.1. Governing Equations for Fluidized Bed. A fluidized bed consists of two phases, bubble phase and dense phase. Different from fixed beds, mass transfer of NOx in the fluidized bed not only exists between the bubble phase and dense phase but also between the gas phase and particle phase, as shown in Figure 2. It is assumed that there are no solid particles in the bubble phase. Adsorption and reaction take places only on the particles of the dense phase. Governing equations for different phases are eqs 16 to 18 in Table 4. Initial and boundary conditions are eqs 19 to 21. In these equations, CH, CL, and Cs* are the NOx gas concentra9368
dx.doi.org/10.1021/ie5011228 | Ind. Eng. Chem. Res. 2014, 53, 9365−9376
Industrial & Engineering Chemistry Research
Article
Table 5. Equations for the Estimation of Hydrodynamic Parameters param. mean bubble size
correlations
Db =
0.21Hf0.8P 0.06(U
0.42
− Umf )
ref −4 2
2
exp[− 1.4 × 10 P − 0.25(U − Umf )
20
− 0.1P(U − Umf )]
a
bubble rise velocity
Ub = (U − Umf ) + 0.711 gD b
21
bubble phase flow rate
Q b = Y (U − Umf )A
22
coefficient Y
Y = 0.116Hs0.26(U − Umf )0.09 d p−0.48ρp−0.19
20a
Hs
Hs = Hmf /(1 − εmf )
20
bubble volume fraction
εb = Q b/(UbA)
expanded bed height
H = Hmf /(1 − εb)
20
Note: Cai’s equation was slightly modified by fitting to our bed expansion experimental data.
modeling. Interphase mass transfer coefficient is estimated by the Sit and Grace24 equation for three-dimensional freely bubbling beds: kHL
⎡ Dmεmf Ub ⎤0.5 Umf = + 2⎢ ⎥ 3 ⎣ πD b ⎦
Table 6. Fitted Kinetic Model Parameters submodel adsorption
(23) reduction
3.2.5. Gas and Solid Mixing. Dispersion in the bubble phase was assumed to be 0, DL = 0, according to the two-phase model of Grace.25 For axial solids dispersion in bubbling beds, there are several correlations in the literature.26,27 The correlation of Morooka28 was selected in this study due to its similar experimental settings, with U ≥ 0.02 m·s−1, D = 3.2−19.5 cm, H = 1−3 m for FCC particles of dp = 80 μm. ⎡ ⎛ U ⎞0.8⎤ U ⎢ UD Pes = = 19(D + 0.23) / 1 + 6.5⎜ ⎟ ⎥ gD ⎢⎣ Ds ⎝ gD ⎠ ⎥⎦
23
model and param.
fitted experiment (fixed bed)
0.5[O*] = qeO2 = 6.88CO0.736 2
O2 adsorption
0.505 [NO2 *] = qeNOx = 350CO0.736 C NO 2
NOx adsorption
0.8 r3 = 2.5C HC [O*]0.7
HC oxidation
0.1 r4 = 0.0155[NO2 *]0.012 C HC [O*]0
NOx reduction
the reaction of adsorbed NO2 on the catalyst surface for the currently tested reaction conditions. Parts a and b of Figure 3 present the NOx and HC conversion at different O2 and hydrocarbon concentrations, with the symbols representing the experimental data and lines for the modeled values. Experimental data at O2 = 0.5%, 4%, and 8% were used for the fitting of kinetic parameters and data at O2 = 1% and 2% were used for verification. The model shows a good fitting to the experiment for most of the data points with an average error of about 7% for NOx conversion and about 5% for HC conversion. However, poor agreement was observed for HC conversion at inlet HC/NO = 4. It is suspected to be caused by the catalyst deactivation at higher HC concentrations since unreacted HC tends to deposit on the catalyst surface. At a given O2 concentration, NOx conversion is seen to increase with increasing the inlet hydrocarbon concentration or inlet HC to NOx molar ratio. For experimental data, NOx reduction increased as O2 increased from 0.5 to 1%, but then decreased with further increase in O2 concentration. This is because that small amount of O2 is required to oxidize NO to NO2, but higher O2 concentration in the flue gas leads to the hydrocarbon oxidation, leaving less HC reductant for NOx reduction and thus decreasing the NOx conversion. The model could accurately predict the negative effect of oxygen at high oxygen concentrations but could not well capture the promotional effect of oxygen at low oxygen concentrations, in the range 0.5%−1%. Increasing the number of data points at oxygen concentration lower than 1% may be able to improve the accuracy of current kinetics. Although the accuracy of the present model is not much improved over our previous work,13 the model itself shows a much better capability over the previous one. With the solid and gas phase equations considered and reaction term embedded into the solid phase, this new model could decouple the adsorption from reaction, enabling it to be applied for the modeling of i-CFB reactors, in which the catalyst
(24)
Gas mixing coefficient in the dense phase is assumed to be the same as dispersion coefficient of solids phase Ds. This assumption was confirmed by the study of Abba,29 who measured both the solid dispersion and dense phase gas dispersion and found the values were similar. There are several correlations in the literature for gas dispersion, either vertical or radial, in fluidized beds.30 However, these dispersion coefficients correlated in the literature are for the overall dispersion, different from the DL and DH in eq 17 and 18 for dense and dilute phases, separately. 3.2.6. Freeboard. Chen and Wen31 developed a comprehensive model incorporating elutriation, entrainment, and reaction to predict the chemical reaction and hydrodynamics occurring in the freeboard region of a fluidized bed. It showed a good agreement with the experimental data obtained from fluidized bed coal combustion. In this study, the Chen and Wen31 model is used to predict the NOx conversion in the freeboard region.
4. MODELING RESULT AND DISCUSSION 4.1. Fixed Bed Kinetic Modeling Result. O2 and NO adsorption were correlated in a previous work.14 Reduction model was coded into Matlab and fitted with experimental data. The best-fitting parameters were found by using the function of “fminsearch” (unconstrained nonlinear optimization) in Matlab and shown in Table 6. In the expression of r4, the reaction order for [O*] is found to be 0 based on regression of experimental data, indicating that oxygen is not important for 9369
dx.doi.org/10.1021/ie5011228 | Ind. Eng. Chem. Res. 2014, 53, 9365−9376
Industrial & Engineering Chemistry Research
Article
Figure 3. Simulated and experimental conversions at different HC/NO ratios in fixed bed, (a) NOx conversion; (b) HC conversion. (Symbols, experimental data; lines, model fitting).
Figure 4. Modeled NOx and HC conversion as a function of superficial gas velocity in the fluidized bed. (Symbols, experimental data; lines, model fitting).
from 1 to 2. As a comparison, NOx conversion at O2 = 1% in the fixed bed, as shown in Figure 3a, increases from 50% to 60% when HC/NO ratio increases from 1 to 2. Therefore, it could be concluded that HC shows similar effect on NOx conversion in both the fluidized bed and fixed bed. For O2 effect, the model could well predict outlet NOx concentrations at lower inlet O2 concentrations with an average error of about 4% at O2 = 1%. However, when inlet O2 concentration is higher, the model tends to overestimate the NOx conversion, especially at lower Ug. The reason may be that catalysts used in the fixed bed and fluidized bed reaction were prepared separately and a slight difference in the preparation could lead to some variation in the reaction kinetics. In modeling the fluidized bed, kinetics used is developed from the fixed bed catalyst performance data, which may lead to the under-prediction of negative O2 effect. To further verify this assumption, measured NOx conversions between the fluidized bed and fixed bed were compared at HC/NO = 2. In the fluidized bed experiment, NOx conversion decreases from about 62% to 20% when O2 concentration is increased from 1% to 8%. However, in the fixed bed experiment, as shown in Figure 3a, NOx conversion decreases from about 60% to 40%
carries NOx from adsorption zone to the reduction zone and the solid phase needs to be modeled separately. 4.2. Fluidized Bed Modeling Result. The model is coded in MATLAB and the differential equations are solved with Crank−Nicolson method. The outlet concentration of the fluidized bed is calculated by eq 25. The reaction in the freeboard region contributes 4−8% to the total conversion in the fluidized bed reactor. Cout =
C LUL ΦL + C HUH ΦH UL ΦL + UH ΦH
(25)
Figure 4a presents the modeled and experimental NOx conversions as a function of superficial gas velocity at different inlet O2 concentrations and HC/NO ratios. It could be seen that the model could well capture the effect of inlet O2 and HC concentration. Experimental data show that increasing HC/NO ratio has a positive effect on NOx conversion, while increasing O2 concentration has a negative impact on the NOx conversion. This observation is consistent with the modeling results. At O2 = 1%, NOx conversion in the fluidized bed is predicted to increase from about 52% to 62% when HC/NO ratio increases 9370
dx.doi.org/10.1021/ie5011228 | Ind. Eng. Chem. Res. 2014, 53, 9365−9376
Industrial & Engineering Chemistry Research
Article
when O2 concentration is increased from 1% to 8%. Although the data could not be directly compared due to different catalyst loading and gas velocity, the trend clearly shows that O2 concentration imposed a more serious negative effect in the fluidized bed than in the fixed bed, which may be attributed to the slight difference in kinetics for the catalyst separately prepared for the fluidized bed and/or the more complex gas and solids mixing patterns in the fluidized bed. In terms of gas velocity, measured NOx conversion is not affected noticeably by superficial gas velocity for most data, which is likely associated with a better gas−solid mass transfer at higher Ug. However, the modeling results could not adequately capture this trend. NOx conversion of the modeling results clearly decreases with Ug, although it is not very sensitive to gas superficial velocity. The differences between the experiment and modeling could be caused by the variation of reactor temperature in the fluidized bed, which fluctuated between 340 and 360 °C during the experiment. Comparing to NOx conversion, the model shows a very poor agreement on HC conversion, as shown in Figure 4b. Most of the measured HC conversion data ranges from 30% to 60%, while corresponding modeling data range from 65% to 100%. The difference between the experimental data and modeling results could be caused partly by experimental errors. The HC conversion was estimated based on measured CO and CO2 concentrations by assuming that all converted HC goes to CO/ CO2, ignoring other possible reaction intermediates. However, in our experiments some unidentified byproducts were detected by gas chromatography, which was believed to be HCN or cyanogens.32 This will likely cause the underestimation of HC conversion. On the other hand, the reaction kinetics fitted from fixed bed data, as shown in Figure 3b, tends to overpredict fixed bed HC conversion data. This overprediction is then likely carried over to the fluidized bed, contributing to the overprediction of HC conversion in the fluidized bed. Although the model overpredicts the experimental data significantly, it captures the right trends of the effect of O2, HC/NO ratio on HC conversion. HC conversion is predicted to decrease with increasing HC/NO ratio and decreasing O2 concentration. However, the experimental HC conversion decreases with Ug, while modeled conversion increases with Ug. Besides experimental errors, it is also suspected that the model may have overpredicted the mass transfer of hydrocarbons at higher superficial gas velocities. But overall speaking, both experimental and modeling curves are not very sensitive to Ug. To further improve the model, it is recommended that the catalyst particles be tested in a fixed bed reactor after running in the fluidized bed reactor. By doing so, changes of the kinetics over time could be detected and updated kinetic model equations could be used in the reactor simulation. In order to well capture the mass transfer of hydrocarbons in fluidized beds, the parameters used in the mass transfer model should be correlated directly from fluidized bed experimental data. 4.3. Simulation of Reduction Zone of an i-CFB. As shown in Figure 5, different from the batch operated fluidized bed in which solids stay in the bed all the time, solids are fed to the reduction zone of the reactor from the bottom and discharged from the top into the adsorption zone. The solids fed to the reduction zone are saturated with NOx, which is adsorbed from the flue gas in the adsorption zone. NOx on the catalyst surface is then reduced by reducing agent in the reduction zone. Different from conventional fluidized bed, gas fed into the reduction zone of the i-CFB contains only hydrocarbon and
Figure 5. Flow diagram of fluidized bed and decoupled i-CFB.
a small amount of O2, while NOx is carried into the bottom by solids from the adsorption zone. Two more parameters should be considered in modeling the reduction zone of the i-CFB, solids circulation rate, Gs, and NOx concentration on the fed particle surfaces, Cs,feeding. The solids circulation rate, Gs, determines how much solids are fed into the reduction zone and how long the solids could stay to have the NOx on the surface reduced by hydrocarbon. In most cases, the catalyst is not completely saturated by NOx in the flue gas and the NOx concentration on the solids surface leaving the adsorption zone, defined as Cs,feeding, depends on the design and operation of the adsorption zone of the reactor. In a conventional fluidized bed, NOx in the flue gas is adsorbed onto the surface of the catalyst and then reduced by hydrocarbon reductant. But in the reduction zone of the i-CFB, NOx on the surface of catalysts is directly reduced by hydrocarbon reductant. At the same time, NOx could also desorb into the gas phase and slips out of the reactor. Therefore, different from the conventional fluidized bed, the catalytic NOx reduction efficiency in the reduction zone of the i-CFB is given by ⎛ Q NO ,slip + Q NO ,d − a ⎞ x x ⎟100 Xr = ⎜⎜1 − ⎟ Q ⎝ ⎠ NOx ,feeding
(26)
where QNOx,feeding, QNOx,slip, and QNOx,d‑a are the amounts of NOx fed into the reduction zone by the catalysts, escaped from the reactor with the exhaust gases because of desorption, and returned to the adsorption zone with the circulating catalysts, respectively, in unit of mol·s−1. Another parameter used to evaluate the performance of decoupled i-CFB is the slip ratio, defined as Sl =
Q NO ,slip x
Q NO ,feeding x
100 (27)
where QNOx,feeding is the total amount of NOx fed to the i-CFB and the unit is also mol·s−1. The slip ratio, Sl, which represents the percentage of NOx escaped from the reduction zone, is an important parameter for the calculation of overall deNOx efficiency in the i-CFB. When simulating the reduction zone in the i-CFB, the governing equations are still eqs 16 to 18. It should be noted that the adsorption term in the governing equations accounts for NOx desorption from the catalyst when NOx is fed through the solids. The initial and boundary conditions now become z = 0, C H = Cg,in , Cs = Cs,feeding , C L = Cg,in (28) 9371
dx.doi.org/10.1021/ie5011228 | Ind. Eng. Chem. Res. 2014, 53, 9365−9376
Industrial & Engineering Chemistry Research ∂C H = 0, ∂z
z = H,
t = 0,
C H = 0,
∂Cs = 0, ∂z
Cs = 0,
∂C L =0 ∂z
CL = 0
Article
(29) (30)
In order to evaluate the performance of the i-CFB, NOx reduction in a conventional fluidized bed (A) and reduction zone of an i-CFB reactor fed with solids saturated with NOx (B) is directly compared. Bed catalyst loading is set to be 3 kg and the bed temperature is set at 350 °C. Inlet conditions for both gas and solids are shown in Table 7. The total amounts of Table 7. Inlet Conditions for Different Fluidized Bed reactor A: conventional fluidized bed feeding gas concn.
superficial gas velocity feeding solids concn. solid circulation rate
reactor B: decoupled i-CFB fed with NOx saturated solids
Cg0,NOx = 600 ppm
Cg0,NOx = 0 ppm
Cg0,HC = 600 ppm Cg0,O2 = 1%
Cg0,HC = 600 ppm Cg0,O2 = 1%
Ug = 0.25 m·s−1 ∼ 0.45 m·s−1
same as reactor A
N/A
Cs0,NOx = Cs,NOx*
0
determined by NOx feeding rate
Figure 6. Dimensionless NOx concentration as a function of Ug for different reactors.
Therefore, a shorter residence time of the solids at a higher Ug leads to a lower NOx reduction and NOx desorption, which can further lead to a higher solids phase NOx concentration at the reactor outlet. Figure 7 shows the overall NOx reduction efficiency, Xr, in the two reactors. NOx conversion in reactor A is determined by
NOx in the feed are the same for the two reactors, although the gas carriers are different. In reactor A, NOx is carried into the reactor by the gas feed (flue gas), while in reactor B, NOx is supplied by the solids feed. The NOx concentration in the fed solids of reactor B is set to be in equilibrium with the flue gas, calculated by the equilibrium of NOx: nNO2 n * = kFNO Cg0,NO Cs0,NOx = Cs,NO C O2 x g0,O2 x 2
(31)
The total amount of NOx fed into the two reactors is the same. Thus, solid circulation rate in reactor B could be calculated as Gs =
UCg0,NOxρp Cs0,NOx
kg/m 2·s (32)
NOx reduction in the two reactors is then simulated at different superficial gas velocities. Outlet concentrations of NOx in both the solids and gas phase are plotted in Figure 6. The concentrations investigated are dimensionless values, divided by Cg0 or Cs0. Reactor A shows the highest gas and solids phase NOx concentrations. In reactor A, NOx is transferred from gas phase to the solids phase. So, the dimensionless concentration in the gas phase is higher than the concentration in the solids phase. The superficial gas velocity, Ug, shows little influence on the outlet NOx concentration. When NOx is fed through the solids phase in reactor B, outlet NOx concentration of the gas phase is very low, because NOx in the gas phase comes from NOx desorption from the solids surface. A low gas phase concentration in reactor B could guarantee a low NOx slip, improving the overall performance of i-CFB. A high solids phase NOx concentration at the outlet of reduction zone, however, implies a high NOx concentration on the catalyst returning to the adsorption zone, imposing a negative impact on NOx capture in the adsorption zone of i-CFB. Different from reactor A, the outlet NOx concentration in reactor B increases substantially with increasing Ug, especially in the solids phase. Based on the equation for estimating solids circulation rate, eq 32, Gs is expected to be higher at higher Ug.
Figure 7. Xr as a function of Ug in different reactors.
the gas phase outlet concentration only since there are no solids leaving the reactor. The values are around 50%, slightly lower than the conversion simulated in the previous section due to a smaller bed catalyst loading. NOx conversion in reactor B is determined by both gas and solids phase NOx concentrations. The overall conversion in reactor B is much higher than the conversions in reactor A. This is anticipated since one important step in the deNOx reaction, adsorption, is already completed before the solids enter the reactor. NOx reduction decreases very quickly with increasing Ug, due to the shorter solids residence time at a higher solids circulation rate. Solid circulation rate, Gs, in i-CFB is a very important operating parameter for determining both the NOx capture efficiency in the adsorption zone and reduction efficiency in the reduction zone. It is anticipated that at a higher Gs, more solids will be recirculated to the adsorption zone to capture NOx, leading to lower NOx concentration on catalyst particles leaving 9372
dx.doi.org/10.1021/ie5011228 | Ind. Eng. Chem. Res. 2014, 53, 9365−9376
Industrial & Engineering Chemistry Research
Article
the adsorption zone. It is thus worth to investigate how Gs could affect the overall reduction efficiency and NOx slip ratio in the reduction zone. The performance of i-CFB is simulated at Ug = 0.35 m·s−1 with Gs changing. Other conditions are the same with the settings of reactor B in Table 7. To investigate the effect of Gs on the performance of the reduction zone of i-CFB, NOx molar flow rate is set to a constant while decreasing the solids phase NOx concentration at higher solids circulation rate, as shown in eq 33. Gs C = s0 Gs0 Cs
To investigate the influence of HC concentration on the performance of the reduction zone of i-CFB, NOx reduction at Ug = 0.35 m·s−1 is simulated with other conditions kept the same as those for reactor B given in Table 7. Solids circulation rate Gs is increased from an initial value Gs0, to twice of Gs0, 2*Gs0, while solids phase NOx concentration is kept at a constant, Cs0. Here, both Gs0 and Cs0 are the same as the values used in Figure 8. To keep feed HC/NOx = 1, either superficial gas velocity Ug or HC concentration CHC is increased proportionally. The two cases are simulated with the results of Xr and Sl plotted as a function of dimensionless Gs, Gs/Gs0, in Figure 9. Xr for either case illustrates that the overall NOx
(33)
This is based on the assumption that the total NOx feeding rate and capture efficiency in the adsorption zone of the i-CFB reactor remain the same. The simulated Xr and Sl values are plotted in Figure 8. It can be seen that the overall NOx
Figure 9. Xr and Sl values at different solid circulation rate.
reduction efficiency decreases when Gs increases, due to a shorter residence time of solids in the reactor. NOx slip ratio, Sl, increases at the same time. Both Xr and Sl seem to be more sensitive to the change in gas velocity, Ug, than to the change in hydrocarbon concentration, CHC. It is seen that increasing Ug could lead to a lower overall NOx reduction and more NOx slippage, comparing to increasing the HC concentration. This is likely caused by the shorter contacting time between catalyst and gases at the higher Ug. Therefore, a reductant gas stream with higher HC concentrations is preferred in an i-CFB to improve the overall deNOx performance. It should be also noted that this conclusion is valid only when other operating parameters of the i-CFB are kept at constants. In a real i-CFB, when HC concentration is higher, gas velocity in the reduction zone becomes lower, which could further change the solids circulation rate in the i-CFB and affect NOx capture in the adsorption zone. An integral model is thus needed to explore the performance of the i-CFB.
Figure 8. Overall efficiency, Xr, and slip ratio, Sl, as a function of Gs/Gs0.
reduction efficiency is slightly lower at higher Gs, indicating that Xr is hardly influenced by Gs if inlet NOx molar flow rate is maintained at a constant value. On the other hand, the change of NOx slip ratio is much bigger than overall reduction efficiency. The slip ratio increases by about 10% when Gs is decreased from 2*Gs0 to Gs0/2. This could be explained by the difference in solid phase NOx concentration Cs at different solids circulation rate Gs. Higher Cs at a lower Gs increases NOx mass transfer between solid and gas phase, which leads to a higher NOx desorption and slip ratio. Therefore, higher solid circulation rate is preferred for the i-CFB if the total feed NOx molar flow rate is kept constant. It should be noted that the solid phase NOx concentration at the outlet of the reduction zone is expected to be higher at a lower slip ratio, at a constant overall reduction efficiency. Another important issue for i-CFB operation is how to set the hydrocarbon concentration in the reduction zone. In conventional fluidized bed or fixed bed reactor, HC concentration is determined based on the given NOx/HC ratio, since the superficial gas velocity is determined by the flow rate of flue gas. However, gas streams of reducing agent and flue gases are fed into different zones of i-CFB, making HC concentration and gas flow rate adjustable independently in the reduction zone.
5. CONCLUSIONS A model, which includes a bubble phase, gas in the dense phase and solids in the dense phase, was developed for the selective catalytic reduction of NOx in a fluidized bed reactor and was validated by the NOx conversion data measured in a fluidized bed reactor. The fluidized bed deNOx model was then applied to simulate NOx reduction in the reduction zone of the i-CFB reactor. Comparing with conventional fluidized beds, NOx conversion could be improved if NOx is fed through the solids phase, such as what happens in the reduction zone of an i-CFB. Also, for the reduction zone of the i-CFB, a higher solids circulation rate is preferred if NOx feed rate is kept at a 9373
dx.doi.org/10.1021/ie5011228 | Ind. Eng. Chem. Res. 2014, 53, 9365−9376
Industrial & Engineering Chemistry Research
Article
constant. This is because NOx slip ratio is lower at higher Gs, which could improve the overall NOx reduction efficiency of the i-CFB. If NOx feed rate is increased, HC feed rate should be increased proportionally. In this case, changing HC concentration and keeping gas velocity at a constant are preferred than changing gas velocity and keeping HC concentration at a constant. The overall NOx conversion becomes higher and NOx slip ratio is lower when the HC concentration is increased. These simulations could provide valuable criteria for the i-CFB design and operation. Further modeling of the i-CFB reactor and comparison with experimental data from an i-CFB reactor will be presented in part II of this study entitled ‘Catalytic NOx reduction in a novel i-CFB reactor: II. Modeling and simulation of i-CFB reactors’.
From eq A2, kK, as a lumped factor, is determined by the external and internal mass transfer characteristics and is independent of adsorption equilibrium K. Therefore, eq A2 could be simplified to eq A5, in which α is a constant calculated from kc, Rp, and De. Then it becomes clear that k is linear to 1/K, as shown in eq A6.
k = α /K
k f = α /K
APPENDIX: DETAILED DISCUSSION OF KF Mass transfer term kf (Cs−S*s ) used in eqs 8 and 9 is slightly different with the classic term k(q*−q)̅ for mass transfer18 given in eq A1.
If there is a kf value (kf0) fitted at one condition, other kf values could be calculated from eq A8. Here, kc, Rp, and De remain unchanged. It should be noted that Ki is the adsorption equilibrium of linear isotherm.
(A1)
Rp R p2 1 = + kK 3kc 15De
■
AUTHOR INFORMATION
Notes
The authors declare no competing financial interest.
■
ACKNOWLEDGMENTS The authors are grateful to the Natural Science and Engineering Research Council (NSERC) of Canada for the financial support in the form of a discovery grant. X. Cheng is also grateful to a scholarship from the China Scholarship Council (CSC).
■
(A3)
For the estimation of the external mass transfer coefficient kc in fixed beds, the correlation of Wakao34 could be used: 3 < Re < 104
(A9)
*Tel: (604) 822-4408. Fax: (604) 822-6003. Email: xbi@ chbe.ubc.ca.
where the second term on the right side represents the internal resistance and the first term represents the external mass transfer resistance, kcav, since for a sphere the surface area/unit volume, av, is given by
= 2.0 + 1.1Sc1/3Re1/2
K i = qi /Cg , i = kFiCg,nii /Cg, i
Corresponding Author
(A2)
4πR p2/[(4/3)πR p3] = 3/R p
(A8)
The value of Ki should be obtained according to eq A9 at each condition since the adsorption of NOx on Fe/ZSM-5 follows Freundlich isotherm.
where q* is the adsorbate loading in equilibrium with the solute concentration, Cg, in the bulk fluid; Cg* is the concentration in equilibrium with the average loading q;̅ k is the overall mass transfer coefficient, which includes both external and internal transport resistances; and K is the adsorption equilibrium constant for a linear adsorption isotherm of the form q = KCg. For the term kK(Cg−Cg*) in eq A1, it is assumed that the adsorption follows a linear isotherm. However, the adsorption of NOx on Fe/ZSM-5 follows the Freundlich type isotherm.14 The model with mass transfer treated as kK(Cg−Cg*) could not adequately capture the adsorption characteristics of Fe/ZSM-5. Thus, the mass transfer term was modified to capture the Freundlich type adsorption performance. The modified model was investigated in fixed bed adsorption14 and then confirmed to provide better fit with the experiment than the classic model. Similar modification of the model was also adopted by Dasgupta33 to enable the Langmuir form of isotherm to be used. However, in Dasgupta’s model, only intraphase mass transfer resistance was considered. In order to investigate the characteristics of kf, there is a need to quote eq A1 again. Since the modified term kf(Cs−Cs*) is comparable to the classic term k(q*−q), ̅ it is expected that characteristics of parameter kf is similar to k. However, there is a correlation in the literature to calculate k directly. Equation A2 is the most often used equation to calculate kK as a lumped factor.
Dm
(A7)
k f = k f 0K i0/K i
∂q ̅ = k(q* − q ̅ ) = kK (Cg − Cg*) ∂t
kcd p
(A6)
It could be further expected that there exists a similar relationship between kf and K.
■
Sh =
(A5)
kK = α
(A4) 9374
NOMENCLATURE A = reactor cross area, m2 ab = contacting area of bubbles in unit volume, m−1 C = concentration, mol·m−3 Cg = concentration in gas phase, mol·m−3 Cg,in = inlet concentration in gas phase, mol·m−3 Cg,i = gas phase concentration of species i, mol·m−3 Cg0,NOx = initial NOx concentration in gas phase, mol·m−3 Cg0,O2 = initial O2 concentration in gas phase, mol·m−3 CH = concentration in high (dense) phase, mol·m−3 cat CHC = hydrocarbon concentration, mol·m−3 CL = concentration in low (bubble) phase, mol·m−3 cat CNO = NO concentration, mol·m−3 CO2 = O2 concentration, mol·m−3 Cout = outlet concentration, mol·m−3 Cs = concentration in solids phase, mol·m−3 cat Cs0 = initial concentration in solids phase, mol·m−3 cat Cs* = solids phase NOx concentration in equilibrium with gas phase inlet NOx concentration, mol·m−3 cat Cs,feeding = concentration in feeding solids, mol·m−3 cat Cs,in = inlet concentration in solids phase, mol·m−3 dx.doi.org/10.1021/ie5011228 | Ind. Eng. Chem. Res. 2014, 53, 9365−9376
Industrial & Engineering Chemistry Research
Article
Cs0,NOx = initial NOx concentration in solids phase, mol·m−3 Cs,NOx* = NOx concentration in solids phase, mol·m−3 Cg* = gas phase concentration in equilibrium with solids phase, mol·m−3 D = diameter, m Db = bubble diameter, m De = effective diffusivity, m2·s−1 Dg = gas axial dispersion coefficient, m2·s−1 DH = axial dispersion coefficient of high (dense) phase, m2· s−1 DL = axial dispersion coefficient of low (bubble) phase, m2·s−1 Dm = molecular diffusivity, m2·s−1 Ds = axial dispersion coefficient of solids phase, m2·s−1 dp = particle diameter, m g = acceleration due to gravity, 9.8 m·s−2 Gs = solids circulation rate, kg·m−2·s−1 Gs0 = initial solids circulation rate, kg·m−2·s−1 H = bed height, m Hf = expanded bed height, m Hmf = bed height at minimum fluidization, m Hs = bed height, m k = overall mass transfer coefficient, s−1 K = adsorption equilibrium in linear adsorption isotherm k1 = reaction rate coefficient of reaction 1 (forward) k1′ = intraphase mass transfer coefficient, s−1 k−1 = reaction rate coefficient of reaction 1 (reversed) k2 = reaction rate coefficient of reaction 2 (forward) k2′ = intraphase mass transfer coefficient, s−1 k−2 = reaction rate coefficient of reaction 2 (reversed) k3 = reaction rate coefficient of reaction 3 k4 = reaction rate coefficient of reaction 4 kc = external mass transfer coefficient, m·s−1 kf = intraphase mass transfer coefficient, s−1 kf0 = intraphase mass transfer coefficient, s−1 kFi = Freundlich adsorption constant related to sorption capacity of species i kFNO2 = Freundlich adsorption constant related to sorption capacity of NO2 kFO2 = Freundlich adsorption constant related to sorption capacity of O2 kHL = interphase mass transfer coefficient, m·s−1 Ki = adsorption equilibrium of species i Ki0 = adsorption equilibrium of species i kNO2′ = lumped Freundlich adsorption constant related to sorption capacity of NO2 m1 = reaction order m2 = reaction order m3 = reaction order m4 = reaction order m5 = reaction order ni = constant in Freundlich adsorption equations of species i nNO2 = constant in Freundlich adsorption equations of NO2 nO2 = constant in Freundlich adsorption equations of O2 P = operation pressure, Pa Pes = Péclet number of solids phase q = adsorbate loading, mol·m−3 cat q* = adsorbate loading in equilibrium, mol·m−3 cat qi = adsorbate loading of species i, mol·m−3 cat Qb = volumetric gas flow rate in bubble phase, m3·s−1 qeNOx = adsorption capacity of NOx, mol·m−3 cat qeO2 = adsorption capacity of O2, mol·m−3 cat QNOx,d‑a = molar flow rate of NOx returned from the annulus to the draft tube, mol·s−1
QNOx,feeding = molar flow rate of NOx in the feeding flue gas, mol·s−1 QNOx,slip = molar flow rate of NOx slipping out of the reduction zone, mol·s−1 r3 = reaction rate of reaction 3, mol·s−1·m−3 r4 = reaction rate of reaction 4, mol·s−1·m−3 rO* = reaction rate of O*, mol·s−1·m−3 rNO2* = reaction rate of NO2*, mol·s−1·m−3 rHC = reaction rate of HC, mol·s−1·m−3 Re = Reynolds number, Re = ρUL/μ Ri = reaction rate of species i, mol·s−1·m−3 Rp = radius of particles, m Sc = Schmidt number, Sc = μ/ρDm Sh = Sherwood number, Sh = kfl/Dm Sl = NOx slip ratio from reduction zone, % t = time, s U = superficial gas velocity, m·s−1 Ub = bubble rise velocity, m·s−1 Ug = gas velocity, m·s−1 UH = gas velocity in high (dense) phase, m·s−1 UL = gas velocity in low (bubble) phase, m·s−1 Umf = minimum fluidization velocity, m·s−1 Us = solids velocity, m·s−1 Xr = NOx conversion in reduction zone, % Y = dimensionless coefficient z = axial position in the reactor, m Greek Symbols
■
α = constant εb = bubble fraction εmf = voidage at minimum fluidization εs = solids fraction ρp = density of catalyst particles, kg·m−3 ΦH = fraction of high (dense) phase ΦL = fraction of low (bubble) phase
REFERENCES
(1) Wark, K.; Warner, C. F.; Davis, W. T. Air Pollution: Its Origin and Control, 3rd ed.; Prentice Hall: Englewood Cliffs, NJ, 1998. (2) Liu, Z.; Woo, S. I. Recent advances in catalytic deNOx science and technology. Catal. Rev. 2006, 48, 43−89. (3) Li, Y. J.; Armor, J. N. Selective reduction of NOx by methane on Co-ferrierites 0.1. Reaction and kinetic-studies. J. Catal. 1994, 150 (2), 376−387. (4) Corma, A.; Fornes, V.; Palomares, E. Selective catalytic reduction of NOx on Cu-β zeolites. Appl. Catal., B 1997, 11 (2), 233−242. (5) Chen, H.-Y.; Sachtler, W. M. H. Activity and durability of Fe/ ZSM-5 catalysts for lean burn NOx reduction in the presence of water vapor. Catal. Today 1998, 42 (1−2), 73−83. (6) Yang, T. T.; Bi, H. T.; Cheng, X. X. Effects of O2, CO2, and H2O on NOx adsorption and selective catalytic reduction over Fe/ZSM-5. Appl. Catal., B 2011, 102 (1−2), 163−171. (7) Yang, T. T.; Bi, H. T. A novel fluidized bed reactor for integrated NOx adsorption-reduction with hydrocarbons. Environ. Sci. Technol. 2009, 43 (13), 5049−5053. (8) Leistner, K.; Nicolle, A.; Da Costa, P. Modelling the kinetics of NO oxidation and NOx storage over platinum, ceria, and ceria zirconia. Appl. Catal., B 2012, 111, 415−423. (9) Rankovic, N.; Nicolle, A.; Da Costa, P. Detailed kinetic modeling study of NOx oxidation and storage and their interactions over Pt/Ba/ Al2O3 monolith catalysts. J. of Phys. Chem. C 2010, 114 (15), 7102− 7111. (10) Lindholm, A.; Currier, N. W.; Li, J. H.; Yezerets, A.; Olsson, L. Detailed kinetic modeling of NOx storage and reduction with hydrogen as the reducing agent and in the presence of CO2 and H2O over a Pt/Ba/Al catalyst. J. Catal. 2008, 258 (1), 273−288. 9375
dx.doi.org/10.1021/ie5011228 | Ind. Eng. Chem. Res. 2014, 53, 9365−9376
Industrial & Engineering Chemistry Research
Article
(11) Olsson, L.; Sjovall, H.; Blint, R. J. Detailed kinetic modeling of NOx adsorption and NO oxidation over Cu-ZSM-5. Appl. Catal., B 2009, 87 (3−4), 200−210. (12) Metkar, P. S.; Balakotaiah, V.; Harold, M. P. Experimental study of mass transfer limitations in Fe- and Cu-zeolite-based NH3-SCR monolithic catalysts. Chem. Eng. Sci. 2011, 66 (21), 5192−5203. (13) Cheng, X.; Bi, X. T. Reaction kinetics of selective catalytic reduction of NOx by propylene over Fe/ZSM-5. Chem. Eng. J. 2012, 211−212, 453−462. (14) Cheng, X. X.; Bi, X. T. T. Modeling NOx Adsorption onto Fe/ ZSM-5 Catalysts in a Fixed Bed Reactor. Int. J. Chem. Reactor Eng. 2013, 11 (1), 1−12. (15) Richard, D.; Nunez, M. D. D.; Schweich, D. Adsorption of complex phenolic compounds on active charcoal: Breakthrough curves. Chem. Eng. J. 2010, 158 (2), 213−219. (16) Whitman, W. The two film theory of gas absorption. Chem. Metal. Eng. 1923, 29, 146−148. (17) Delgado, J. M. P. Q. A critical review of dispersion in packed beds. Heat Mass Transfer 2006, 42 (4), 279−310. (18) Seader, J. D.; Henley, E. J. Separation Process Principles; John Wiley & Sons, Inc.: New York, 1998; p 835. (19) Cheng, X. X.; Bi, X. T. T. Modeling and simulation of nitrogen oxides adsorption in fluidized bed reactors. Chem. Eng. Sci. 2013, 96, 42−54. (20) Cai, P.; Michele, G. D.; Gradassl, A. T. A generalized method for predicting gas flow distribution between the phases in FBC. Fluidized Bed Combustion 1993, 2, 991−1002. (21) Clift, R.; Grace, J. R. Continuous bubbling and slugging. Fluidization; Academic Press: London, 1985; Chapter 3. (22) Grace, J. R. Fluid Beds as Chemical Reactors; John Wiley and Sons Ltd.: Chichester, 1986. (23) Cheng, X. X.; Bi, X. T. T. Hydrodynamics of an i-CFB deNOx reactor. Powder Technol. 2014, 251, 25−36. (24) Sit, S. P.; Grace, J. R. Effect of bubble interaction on interphase mass transfer in gas fluidized bed. Chem. Eng. Sci. 1981, 36, 327−335. (25) Grace, J. R. Generalized models for isothermal fluidized bed reactors. Wiley: New Delhi, 1984. (26) Gunther, A.; Helmrich, H.; Schugerl, K. Axial-dispersion of solid and properties of gas-bubbles of a continuous fluidized-bed. Chem. Ing. Tech. 1980, 52 (10), 841−841. (27) Weislehn, G. Determination of solid mixing coefficients in a gas fluidized bed. Chem. Ing. Tech. 1970, 42 (8), 524−&. (28) Morooka, S.; Kato, Y.; Miyauchi, T. Holdup of gas bubbles and longitudinal dispersion coefficient of solid particles in fluid-bed contactors for gas−solid systems. J. Chem. Eng. Jpn. 1972, 5 (2), 161. (29) Abba, I. A. A generalized fluidized bed reactor model across the flow regimes. PhD Thesis, University of British Columbia, Vancouver, 2001. (30) Bi, H. T.; Ellis, N.; Abba, I. A.; Grace, J. R. A state-of-the-art review of gas−solid turbulent fluidization. Chem. Eng. Sci. 2000, 55 (21), 4789−4825. (31) Chen, L. H.; Wen, C. Y. Model of solid gas reaction phenomena in the fluidized-bed freeboard. AIChE J. 1982, 28 (6), 1019−1027. (32) Chen, H. Y.; Voskoboinikov, T.; Sachtler, W. M. H. Reaction intermediates in the selective catalytic reduction of NOx over Fe/ ZSM-5. J. Catal. 1999, 186 (1), 91−99. (33) Dasgupta, S.; Agnihotri, V.; Gupta, P.; Nanoti, A.; Garg, M. O.; Goswami, A. N. Simulation of a fixed bed adsorber for thiophene removal. Catal. Today 2009, 141 (1−2), 84−88. (34) Wakao, N.; Funazkri, T. Effect of fluid dispersion coefficients on particle-to-fluid mass-transfer coefficients in packed-bedsCorrelation of Sherwood numbers. Chem. Eng. Sci. 1978, 33 (10), 1375−1384.
9376
dx.doi.org/10.1021/ie5011228 | Ind. Eng. Chem. Res. 2014, 53, 9365−9376
