Energy & Fuels 2007, 21, 3593–3598
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Sequential Simulation of a Fluidized Bed Membrane Reactor for the Steam Methane Reforming Using ASPEN PLUS A. Sarvar-Amini,† R. Sotudeh-Gharebagh,*,†,‡ H. Bashiri,† N. Mostoufi,† and A. Haghtalab§ Process Design and Simulation Research Centre, School of Chemical Engineering, UniVersity of Tehran, P.O. Box 11155-4563, Tehran, Iran, and Chemical Engineering Department, Faculty of Engineering, Tarbiat Modares UniVersity, Tehran, Iran ReceiVed June 22, 2007. ReVised Manuscript ReceiVed August 17, 2007
A simulation model is developed using ASPEN PLUS to predict the performance of a fluidized bed membrane reformer. Because there are physical and chemical phenomena interacting in the fluidized bed membrane reformer, two submodels seem necessary in the model. These submodels are the hydrodynamic and reaction submodels. The hydrodynamic submodel is based on the dynamic two-phase model, and the reaction submodel is derived from the literature. The reformer is divided into two regions: a dense bed and freeboard. The dense bed is divided into several sections. At each section, the flow of the gas is considered as the plug flow through the membrane and bubble phases and perfectly mixed through the emulsion phase. The sets of the experimental data were used from the literature to validate the model. Close agreement was observed between the model predictions and experimental data. This model can be used for the simulation of nonideal fluidized bed membrane reactors inside the ASPEN PLUS process simulator.
Introduction Steam reforming of hydrocarbons is the leading industrial process for producing hydrogen and synthesis gases for ammonia and methanol production, hydrocracking and hydrotreating, Fischer–Tropsch synthesis, and other important processes in petroleum refining and petrochemical industries.1 Typical industrial reformers are formed of hundreds of fixed bed tubes packed with nickel catalyst particles within gas-fired furnaces.2 Industrial fixed-bed steam reformers have several drawbacks, such as thermodynamic equilibrium limitations, carbon formation on the catalyst, a large temperature gradient, and low heattransfer rates, which seriously affect their operation and performance. In recent years, new membrane-based fluidized bed reformers have been under development to overcome these shortcomings. Considerable attention has been paid to the fluidized bed membrane reformers (FBMRs) as multifunctional reformers.3 The main advantages of FBMR are a uniform temperature in the bed, an increase in the hydrogen production because of changes in thermodynamic equilibrium conditions, an enhancement of methane conversion, a simultaneous reaction and separation of hydrogen, an elimination of intercatalyst diffusion limitations, a good heat-transfer capability, and a more compact * To whom correspondence should be addressed. Fax: (0098) 21-66461024. E-mail:
[email protected]. † University of Tehran. ‡ On sabbatical leave (2006–2007) at the Chemical Engineering Department, Qatar University, Doha 2713, Qatar. § Tarbiat Modares University. (1) Adris, A. M.; Lim, C. J.; Grace, J. R. The fluidized-bed membrane reactor for steam methane reforming: Model verification and parametric study. Chem. Eng. Sci. 1997, 52, 1609–1622. (2) Chen, Z.; Prasad, P.; Yan, Y.; Elnashaie, S. S. E. H. Simulation for steam reforming of natural gas with oxygen input in a novel membrane reformer. Fuel Process. Technol. 2003, 83 (1–3), 235–252. (3) Abashar, M. E. E. Coupling of steam and dry reforming of methane in catalytic fluidized bed membrane reactors. Int. J. Hydrogen Energy 2004, 29, 799–808.
design.3 Despite these advantages, very few attempts have been reported in the literature on the industrial commercialization of FBMRs because of technological difficulties, complexity in the scale up, and economics.4 In addition, there are some difficulties with experiments to study the effects of operational conditions on the performance of FBMRs, such as the controllability of operating variables, proper maintaining of highly expensive membrane tubes, complex operational concerns, and tedious experiments involved. Furthermore, the planning and implementation of experiments, the analysis of the full array of interactions between system components and bed materials, internal surfaces, and measurement devices, and interpretation of results to diagnose abnormal operating conditions present operators with a range of difficult intellectual challenges. When the reformer is modeled, these effects could be investigated while saving time without a further need to generate highly expensive data. Modeling the FBMR as a nonideal system is useful to investigate the effect of key operating parameters, optimization, and scale up of such reformers. Because two types of phenomena, i.e., physical and chemical, coexist in FBMRs, two submodels, describing these two phenomena, should simultaneously be included in the modeling: the hydrodynamic submodel, which explains the physical phenomena, and the reaction submodel, which describes the chemical reactions occurring in the reformer. Several approaches have been introduced in the literature for modeling FBMRs that can be contributed to high complexity in the hydrodynamic and nonideality of these reformers. Adris et al.1 developed a two-phase bubbling bed model for steam methane reforming with the gas permeation via the membranes. They developed a model in which a plug flow reactor (PFR) for both the bubble and dense phases is (4) Deshmukh, S. A. R. K.; Heinrich, S.; Morl, L.; Vn Sint Annaland, M.; Kuipers, J. A. N. Membrane assisted fluidized bed reactors: Potentials and hurdles. Chem. Eng. Sci. 2007, 62, 416–436.
10.1021/ef7003514 CCC: $37.00 2007 American Chemical Society Published on Web 09/28/2007
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Figure 2. Features of a dense bed in the fluidized bed membrane model.
Figure 1. Typical FBMR.
assumed. This model was later extended by Dogan et al.5 and Rakib et al.6 for the autothermal steam reforming with oxygen to investigate the effect of oxygen. Mleczko et al.7 developed a bubble assemblage model based on the work of Kato et al.8 for the FBMR. In their model, emulsion and bubble phases are considered to be composed of several continuously stirred tank reactors (CSTRs). Chen et al.9 developed a steady-state onedimensional PFR model for the circulating FBMR. Various process simulators, such as ASPEN PLUS and HYSYS, are employed for industrial process simulation. These process simulators only include standard, ideal reactors, such as the PFR and CSTR. Despite the vast application of the commercial simulator in the chemical industries, a fluidized bed membrane reactor unit does not exist in the simulators. The ideal reactors representing different hydrodynamic phases are available in the commercial process simulators; therefore, it would possible to introduce the nonideal fluidized bed membrane reactors in such simulators by the proper combination of these ideal reactors with the external or internal subroutine. In this work, the FBMR is modeled using the sequential modular approach by dividing it into several sections in series. At each section, the flow of gas is considered as a plug flow through the bubbles and membrane and perfectly mixed through the emulsion phase. The standard modules available in ASPEN (5) Dogan, M.; Posarac, D.; Grace, J.; Adris, A. M.; Lim, C. J. Modeling of autothermal steam methane reforming in a fluidized bed membrane reactor. Int. J. Chem. Reactor Eng. 2003, 1, 1–12. (6) Rakib, M. A.; Alhumaizi, K. I. Modeling of a fluidized bed membrane reactor for the steam reforming of methane: Advantages of oxygen addition for favourable hydrogen production. Energy Fuels 2005, 19 (5), 2129–2139. (7) Mleczko, L.; Ostrowski, T.; Wurzel, T. A fluidized bed membrane reactor for the catalytic partial oxidation of methane to synthesis gas. Chem. Eng. Sci. 1996, 51, 3187. (8) Kato, K.; Wen, C. Bubble assemblage model for fluidized bed catalytic reactors. Chem. Eng. Sci. 1969, 24, 1351. (9) Chen, Z.; Elnashaie, S. S. E. H. Steady-state modeling and bifurcation behavior of circulating fluidized bed membrane reformer–regenerator for the production of hydrogen for fuel cells from heptane. Chem. Eng. Sci. 2004, 59 (18), 3965–3979.
Figure 3. Schematic simulation diagram of the reformer.
PLUS could be combined together to represent the real behavior of the fluidized bed membrane reactors. Moreover, several calculator blocks are provided along with the reactor modules for the calculation of hydrodynamic parameters of the reformer. Furthermore, the model has to be easily applicable in the ASPEN PLUS process simulator. Modeling A schematic diagram of a typical FBMR is shown in Figure 1. As shown in this figure, methane is premixed with steam
Sequential Simulation of a FBMR
Energy & Fuels, Vol. 21, No. 6, 2007 3595 Table 1. Correlations and Equations Used in the Modeling
parameter 1
formula
Archimedes number Ar )
2
minimum fluidization velocity
3
bubble velocity
4
bubble rise velocity
5
emulsion velocity
Fgdv3(Fp - Fg)g µ2
Wen and Yu17
Rem ) [(33.7)2 + .0408Ar]0.5 - 33.7 Ub ) U0 - Ue + ubr Kunii and Levenspiel ubr ) 0.711(gdb)
U0 - δUb 1-δ
bubble mean diameter
0.3H D
(
db ) dbm - (dbm - db0)exp -
) Mori and Wen
7
bubble maximum diameter
8
bubble initial diameter
9
bubble-phase fraction for Geldart B particles14
10
11
12 13
emulsion-phase porosity for Geldart B particles14
18
dbm ) 1.64[A(U0 - Umf)]
0.4
[
1.38 A(U0 - Umf) db0 ) 0.2 Nor g
(
[
δ ) 0.534 1 - exp -
(
(
bubble-phase porosity for Geldart B particles14
b ) 1 - 0.146exp -
solid entrainment in the freeboard and expansion zone interphase mass-transfer coefficient
]
U0 - Umf 0.413
e ) mf + 0.2 - 0.059exp -
Miwa et al.16
)]
U0 - Umf 0.429
U0 - Umf 4.439
)
Cui et al.10
) Kunii and Levenspiel
kiq )
[
4DiemfUb umf + 3 πdb
]
mole flow rate of component “i” from the emulsion to bubble phase
Ni ) Askiq(Ci,emulsion - Ci,bubble)
15
mole flow rate of component “i” from the bubble to emulsion phase
Ni ) Askiq(Ci,bubble - Ci,emulsion)
16
hydrogen permeation rate QH2 ) Qeff effective permeability constant of hydrogen
and fed into the bottom of the bed (lower region), where the reforming reactions take place. Inside the bed, a number of hydrogen perm-selective membrane tubes are inserted. Among these membrane tubes, the catalyst is fluidized and the steam reforming of methane takes place. The produced hydrogen selectively permeates through hydrogen membranes and is then carried away by sweep gas, such as steam in the hydrogen membrane tubes.20 The reforming reactions are then completed in the upper region (the freeboard and expansion zone) of the reformer, where the entrainment and carry-over of catalyst particles may have a considerable effect on reforming. The rate of reforming in these regions is directly affected by hydrodynamics. Consequently, these two regions should be considered in the modeling: the dense bed, which is operating under the bubbling fluidization regime, and a more dilute upper region,
( )( )[√ FH MH
Am δ
13
0.5
14
17
13
1⁄2
Ue ) 6
reference
√ ]
Sit and Grace
15
PHR 2 - PHM2
Katsuta et al.19
Qeff ) ηpQ0exp(-Ep/RT)
i.e., the freeboard and expansion zone. The following is a summary of the steps involved in the modeling of these regions: Lower Region. The lower region of the reformer (dense bed) consists of three phases: emulsion, bubble, and membrane phases, as shown in Figure 2. To model the dense bed by considering the complexity of hydrodynamics, the movement of gas through the bubble and membrane phases are considered as the plug flow, while the emulsion phase is assumed as the completely mixed phase. The assumptions made in developing the equations of the model for the dense bed are summarized as follows: (1) The hydrodynamics of both phases are characterized by a dynamic two-phase theory of fluidization.10 (2) Reactions occur in both bubble and emulsion phases only. (3) (10) Cui, H. P.; Mostoufi, N.; Chaouki, J. Characterization of dynamic gas– solid distribution in the fluidized beds. Chem. Eng. J. 2000, 79, 135–143.
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Table 2. Descriptions of ASPEN PLUS Calculators Used in the Simulation calculator name
purpose
Calc1
to calculate volumes, the voidage, and the flow rate for ideal reactors (CSTR and PFR) in the bed to calculate voidage in the freeboard to calculate the average flow rate of components separated from the PFR added to the membrane tube and CSTR to calculate the average flow rate of components separated from the CSTR and added to the membrane tube and PFR
Calc2 Calc3 Calc4
equations used 1–11 (Table 1) 12 (Table 1) 13 and 15–17 (Table 1) 13, 14, 16, and 17 (Table 1)
The bubbles reach their equilibrium size quickly above the distributor. (4) The temperature is assumed to be constant in the dense bed. (5) Hydrogen is the only species transferred to the membrane. (6) The plug flow is considered for sweep gas and hydrogen in the membrane. To develop a sequential modular model, the dense bed is axially divided into several sections, where each section consists of a membrane tube and two ideal reactors: a PFR to represent the gas flow through the bubbles and a CSTR to represent the gas flow through the emulsion. Following reactions taking place in PFR and CSTR:11 CH4 + H2O S CO + 3H2
(1)
CO + H2O S CO2 + H2
(2)
CH4 + 2H2O S CO2 + 4H2O
(3)
As illustrated in Figure 3, the reactions take place at each section in CSTR and PFR followed by the mass transfer among the effluents of each section (bubble, emulsion, and the membrane tube). The ASPEN PLUS simulator was used as a platform to solve the reaction and hydrodynamic submodels simultaneously. The intrinsic capabilities of the simulator were used in developing the model, i.e., at each section, the PFR and CSTR were used to model the gas flow behavior through the bubble, membrane and emulsion phases, respectively. Along with these ideal unit operation blocks provided by ASPEN PLUS, calculator blocks have also been introduced for the calculation of the voidage, volume, and flow rate of gas through the different phases at each section, as described in Table 2. Upon the determination of hydrodynamic parameters for individual phases at each section, ASPEN PLUS solves material and energy balance equations and predicts physical and thermodynamic properties. Effluents of the phases at the end of each section (i.e., unreacted CH4, H2O, and reforming byproducts, such as CO, CO2, and H2O) are then entered into the hierarchical block called “MASS TRANSFER”, where the mass transfer among the phases are calculated, as illustrated in Figure 4. As the mass-transfer calculations are completed among the phases, the corresponding flow distribution is then properly calculated to move to the next section (i + 1). The calculation is continued upward until the top of the bed is reached. Upper Region. The upper region of the reformer consists of two different zones, including the freeboard and expansion zone. Entrainment of solid particles in these regions has a considerable effect on the overall conversion of FBMR. The variation of the (11) Kaihu, H.; Hughes, R. The kinetics of methane–steam reforming over Ni/R-Al2O catalyst. Chem. Eng. J. 2001, 82, 311–328. (12) Adris, A. M.; Lim, C. J.; Grace, J. R. The fluidized bed membrane reactor (FBMR) system: A pilot scale experimental study. Chem. Eng. Sci. 1994, 49, 5833–5843. (13) Kunni, D.; Levenspiel, O. Fluidization Engineering, 2nd ed.; Butterworth-Heineman: Boston, MA, 1991. (14) Grace, J. R. Fluid beds as chemical reactors. In Gas Fluidization Technology; Geldart, D., Ed.; Wiley: Chichester, U.K., 1986; pp 285– 339.
Figure 4. Schematic diagram of the MASS TRANSFER block.
voidage with the height in the upper region of the reformer is dependent upon the transport disengagement height (TDH) and can be calculated according the procedure given by Kunni and Levenspiel.13 The assumptions made in developing the model for the upper region are summarized as follows: (1) PFR is used for the modeling of the freeboard and expansion zone. (2) Hydrogen is permeated to the membrane tube at the exit of the reactor. (3) The amount of voidage in the freeboard is calculated using a calculator block shown in Table 2. The calculation procedure is the same as the lower region by considering two phases in the calculations. Results and Discussion Results of the simulation model presented in this work are compared to the experimental data reported by Adris12,21 in terms of methane conversion, hydrogen yield, and hydrogen permeation through the membrane. In Figure 5, the model prediction is compared to the experimental conversion data at different temperatures for 1–5 sections. As seen in the figure, the methane conversion is increased with an increase in the reformer temperature because of the endothermic behavior of the reforming reactions (except reaction number 2). Moreover, the permeability of hydrogen (15) Sit, S. P.; Grace, J. R. Effect of bubble interaction on interphase mass transfer in gas fluidized beds. Chem. Eng. Sci. 1981, 36, 327–335. (16) Miwa, K.; Mori, S.; Kato, T.; Muchi, I. Behaviour of bubbles in gaseous fluidized beds. Chem. Eng. J. 1972, 12, 187–194. (17) Wen, C. Y.; Yu, Y. H. A generalized method for predicting the minimum fluidization velocity. AIChE J. 1996, 12, 610–612. (18) Mori, S.; Wen, C. Y. Estimation of bubble diameter in gaseous fluidized beds. AIChE J. 1975, 21, 109–115. (19) Katsuta, H.; Farraro, R. J.; McLeUan, R. B. The diffusivity of hydrogen in palladium. Acta Metall. Sin. (Eng. Lett.) 1979, 27, 1111–1114. (20) Chen, Z.; Yan, Y.; Elnashaie, S. S. E. H. Catalyst deactivation and engineering control for steam reforming of higher hydrocarbons in a novel membrane reformer. Chem. Eng. Sci. 2004, 59, 1965–1978. (21) Adris, A. M. A fluidized bed membrane reactor for steam methane reforming: Experimental verification and model validation. Ph.D. dissertation, University of British Columbia, Vancouver, Canada, 1994.
Sequential Simulation of a FBMR
Figure 5. FBMR conversion as a function of the temperature at different numbers of sections (steam/carbon ratio ) 2.4; P ) 0.98 MPa; sweep gas flow rate ) 80 mol/h; sweep gas pressure ) 0.4 MPa; methane flow rate ) 74.2 mol/h).
through the membrane is increased by an increase in the temperature, leading to a shift in thermodynamic equilibrium and, consequently, to an increase in the methane conversion. This figure also shows that by dividing the reformer into 4–5 sections, a close agreement between the predicted and experimental conversion can be found. It is worthy to mention that, by further increasing the number of sections, the rate of the mass transfer among the emulsion phase, bubble phase, and membrane becomes close to the actual reformer. However, there is a limitation in the number of sections, because the increase in the number of sections alters the hydrodynamics of the emulsion phase from well mixed to plug flow. Under these conditions, the model overpredicts the experimental data as expected. It is also important to mention that the number of sections is not the same for other FBMRS and may be adjusted depending upon the hydrodynamics and the reactions involved for a given system. Figure 6 shows a comparison between predicted and experimental hydrogen yields at different temperatures for 1–5 sections. As shown in this figure, the hydrogen yield is increased by an increase in the temperature because of an increase in the methane conversion. This figure also shows that the reformer with 4 or 5 sections has a good agreement with the experimental data, as also confirmed in Figure 5. The predicted methane conversion is compared to the experimental data at different steam/methane ratios for 1–5 sections in Figure 7. As shown in this figure, the methane conversion is increased with an increase in the steam/methane ratio. This is due to the fact that an increase in the steam/ methane ratio causes the reforming reactions to use the excess steam and could be the reason of an increase in the methane conversion. However, with an increase in the steam flow rate, the hydrogen partial pressure in the reformer is reduced, leading to a decrease in the permeability of hydrogen and in the methane conversion. Because of the lower membrane capacity used in the experimental study reported by Adris et al.,12,21 the changes in the methane conversion as a result of the reduction in the permeability are not significant. This figure also shows that the reformer with 4 or 5 divisions has a good agreement with the experimental data in terms of the steam/methane ratio, confirming the findings reported in Figures 5 and 6.
Energy & Fuels, Vol. 21, No. 6, 2007 3597
Figure 6. Hydrogen yield as a function of the temperature at different numbers of sections (steam/carbon ratio ) 2.4; P ) 0.98 MPa; sweep gas flow rate ) 80 mol/h; sweep gas pressure ) 0.4 MPa; methane flow rate ) 74.2 mol/h).
Figure 7. Methane conversion as a function of the steam/methane ratio at different numbers of sections (temperature ) 650 °C; P ) 0.69 MPa; sweep gas flow rate ) 80 mol/h; sweep gas pressure ) 0.4 MPa; methane flow rate ) 41.2, 53, and 74.2 mol/h).
Figure 8 shows the parity plot among the experimental hydrogen permeation rate, the model predictions, and the values calculated by Adris et al.12,21 As shown in this figure, the average hydrogen permeation rates calculated by the model [on the basis of the mean value of effectiveness factor (η ) 0.39) reported by Adris et al.12,21] are in close good agreement with the experimental data and the value calculated by Adris et al.12,21 The small deviation between the models developed in this work with the values reported by Adris et al.12,21 may be contributed to the different hypothesis used in the modeling between these two approaches. The results of F-test analysis also confirms that the model is providing close results in predicting the behavior of the fluidized bed reformer. Conclusions A fluidized bed reactor model was developed for predicting the performance of FBMRs. The model consists of several ideal
3598 Energy & Fuels, Vol. 21, No. 6, 2007
Figure 8. Comparison of experimental data with the models in terms of the hydrogen permeation rate through the membrane.
reactors, which are combined in an appropriate manner, so that the flow of the gas is considered as the plug flow through the bubble and membrane phases and perfectly mixed through the emulsion phase in each section. The catalytic reaction expressions and DTP parameters were adapted as reaction and hydrodynamic submodels, respectively. The ASPEN PLUS process simulator has been used to solve these two submodels simultaneously. The model prediction was compared to experimental data in terms of the methane conversion, the hydrogen yield, and the hydrogen permeation through the membrane, and close agreement was found. Using the approach proposed in this study, the prediction of the nonideal fluidized bed membrane reactor’s behavior at different operating conditions becomes possible in ASPEN PLUS. The results of this study could be used to represent the nonideal fluidized bed membrane system inside the process simulators. Acknowledgment. Financial support from the College of Engineering, Qatar University (project number 06019P) is greatly acknowledged. Dr. S. Al-Asheh, head of the research committee, and Dr. H. Alfadala, the dean of the college, are highly appreciated for their help. The authors are also grateful for the critical and helpful comments of the anonymous reviewers. Nomenclature
db ) Bubble mean diameter (m) dbo ) Bubble initial diameter (m)
SarVar-Amini et al.
dbm ) Bubble maximum diameter (m) D ) Reformer diameter (m) dF ) Particle diameter (m) B ) Bubble phase F ) Mole flow rate (mol/s) G ) Acceleration of gravity (m/s2) H ) Bed height (m) M ) Membrane tube U0 ) Superficial velocity at the onset of fluidization (m/s) Ub ) Bubble velocity (m/s) ubr ) Bubble rise velocity (m/s) Ue ) Emulsion gas velocity (m/s) Umf ) Minimum fluidization velocity (m/s) Vb ) Bubble-phase volume (m3) Ve ) Emulsion-phase volume (m3) VPFR ) PFR volume (m3) Die ) Diffusivity of component i in the gas mixture (m2/s) A ) Reformer surface area (m2) As ) Surface area between phases (m2) Am ) Membrane surface area in each section (m2) kiq ) Interphase mass-exchange coefficient (m/s) Ci,dense ) Average concentration of component i in the dense phase (kmol/m3) Ci,bubble ) Average concentration of component i in the bubble phase (kmol/m3) T ) Temperature (K) R ) Universal gas constant (kJ mol–1 K–1) Ep ) Activation energy for permeation (kJ/mol) E ) Emulsion phase η ) Permeation effectiveness factor Qeff ) Effective permeability constant of hydrogen (mol m–1 s–1 Pa–0.5) R PH2 ) Hydrogen average partial pressure in the reactor (Pa) M PH2 ) Hydrogen average partial pressure in the membrane (Pa) Ar ) Archimedes number Fg ) Gas mixture density (kg/m3) Fp ) Solid particle density (kg/m3) µ ) Gas viscosity (kg m–1 s–1) dv ) Solid particle diameter (m) i ) Section number Greek Symbols εe ) Emulsion-phase voidage εb ) Bubble-phase voidage δ ) Bubble-phase fraction EF7003514
