Energy & Fuels 2006, 20, 463-472
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Performance Enhancement of Steam Methane Reforming Using Tubular Packed-Bed Microreactors and Dilution by Adsorbent R. Rajasree, V. Ravi Kumar,* and B. Dattatraya Kulkarni Chemical Engineering and Process DeVelopment DiVision, National Chemical Laboratory, Pune, 411 008, India ReceiVed July 8, 2005. ReVised Manuscript ReceiVed December 21, 2005
For an endothermic reaction such as steam methane reforming (SMR), temperature gradients in the packed bed can play a significant role in process performance. A reactor design strategy that considers the combination of tubular packed-bed microreactor (TPBM) operation and catalyst dilution by adsorbent is proposed for enhancement of the SMR process. Simulation studies have been performed with the proposed strategy to quantitatively study the effect of heat transfer limitation and process performance in terms of methane conversion and hydrogen purity. This is carried out using a dynamic 2D pseudohomogeneous model incorporating kinetic and reactor/reaction mechanisms describing the process. The results of these studies show that the temperature gradient in TPBM is better controlled with less catalyst loading in comparison with traditionally used larger diameter reactors. It is seen that improved heat transfer characteristics of the TPBM combined with the effects of catalyst dilution and adsorption complement each other, and enhanced performance in terms of methane conversion and hydrogen purity is observed. The findings based on this study give an insight into the performance improvement of the SMR process and show that cost reduction and process miniaturization of fuel processors for fuel cell applications are possible.
1. Introduction The key issue regarding the widespread commercialization of fuel cells is the ability to reduce cost and reactor dimensions to acceptable levels for both stationary and transport applications. The fuel processor that produces H2-rich streams from hydrocarbon-based feedstock is a major component whose design and operation play an important role in achieving overall system efficiency. In fuel processing, research has a long way to go in the development of cost-effective strategies and its overall optimization.1-10 Traditionally, these processes have been carried out in fixed bed reactors packed with catalysts where both exothermic and endothermic reactions occur together. Controlling the reactor temperature is very crucial for optimum reactor performance, and a number of methods have been studied for this aim.11-24 For an exothermic reversible reaction, the temperature may be * Corresponding author. Phone: (91) 20-25902161. Fax: (91) 20-2589 3041. E-mail:
[email protected]. (1) Park, G. G.; Seo, D. J.; Park, S. H.; Yoon, Y. G.; Kim, C. S.; Yoon, W. L. Chem. Eng. Technol. 2004, 101, 87-92. (2) Ouyang, X.; Besser, R. S. J. Power Sources 2005, 141, 39-46. (3) Lee, S. H.; Han, J.; Lee, K. W. J. Power Sources 2002, 109, 394402. (4) Shah, K.; Ouyang, X.; Besser, R. S. Chem. Eng. Technol. 2005, 28, 303-313. (5) Tonkovich, A. Y.; Zilka, J. L.; LaMont, M. J.; Wang, Y.; Wegeng, R. S. Chem. Eng. Sci. 1999, 54, 2947-2957. (6) Srinivas, S.; Dhingra, A.; Im, H.; Gulari, E. Appl. Catal., A 2004, 274, 285-293. (7) Choi, Y.; Stenger, H. G. J. Power Sources 2003, 124, 432-439. (8) Avcı, A. K.; Trimm, D. L.; O ¨ nsan, Z. I. Chem. Eng. J. 2002, 90, 77-87. (9) Avcı, A. K.; Trimm, D. L.; O ¨ nsan, Z. I. Appl. Catal., A 2001, 216, 243-256. (10) Ji, P.; Feng, W.; van der Hooi, H. J.; de Swaan Arnos, J. Ind. Eng. Chem. Res. 2004, 43, 2005-2016.
controlled directly by bypassing cold feed around main feed preheater and distributing it along the length of the reactor.12,13 Another method is cold stream injection that cools the bulk reacting fluid.16 The above operations require an optimal policy in rate of addition of cold feed such that the reaction rate is maximized with respect to variation of the mass flow rate at each point of the reactor. Tailoring catalyst characteristics such as nonuniform catalysts and catalyst size can be applied to control the reaction rate.20 Novel reactor designs, for example, multibed, multitubular reactor,21 and packed-bed membrane reactor22 strategies have also been employed, and further improvements in the design and analysis of these strategies are (11) Lee, K.; Aris, R. Ind. Eng. Chem. Process Des. DeV. 1963, 2, 300306. (12) Dyson, D. C.; Horn, F. J. M. J. Optim. Theory Appl. 1967, 1, 4052. (13) Dyson, D. C.; Graves, J. R. Chem. Eng. Sci. 1968, 23, 435-446. (14) Caldwell, A. D.; Calderbank, P. H. Brit. Chem. Eng. 1969, 14, 1199-1201. (15) Sadhukan, P.; Petersen, E. E. AIChE J. 1976, 22, 808-810. (16) Burghardt, A.; Patzek, T. Chem. Eng. J. 1978, 16, 153-164. (17) Danziger, D.; Popovic, D.; Schulz-Ekloff, G. Can. J. Chem. Eng. 1983, 61, 126-128. (18) Quina, M. M. J.; Ferreria, R. M. Q. Ind. Eng. Chem. Res. 1999, 38, 4615-4623. (19) Quina, M. M. J.; Ferreria, R. M. Q. Chem. Eng. Sci. 2000, 55, 38853897. (20) Hwang, S.; Linke, P.; Smith, R. Chem. Eng. Sci. 2004, 59, 42454260. (21) Papageorgiou, J. N.; Froment, G. F. Chem. Eng. Sci. 1996, 51, 2091-2098. (22) Pena, M. A.; Carr, D. M.; Yeung, K. L.; Varma, A. Chem. Eng. Sci. 1998, 53, 3821-3834. (23) Xiu, G. H.; Li, P.; Rodrigues, A. E. Chem. Eng. Sci. 2003, 58, 34253437. (24) Xiu, G. H.; Li, P.; Rodrigues, A. E. Chem. Eng. Res. Des. 2004, 82 (A2), 192-202.
10.1021/ef050205z CCC: $33.50 © 2006 American Chemical Society Published on Web 02/03/2006
464 Energy & Fuels, Vol. 20, No. 2, 2006
required. Recently, a new generalized strategy using the principle of temperature-induced equilibrium shift by controlling the subsection wall temperature has been developed by Xiu et al.23,24 This strategy is an efficient way for controlling the temperature along the reactor. However, other subsection-controlling parameters such as number of subsections, the subsection packing ratio of adsorbent and catalyst, and temperature in each subsection need to be well-optimized. The heat transfer rates in the packed-bed reactor are known to depend weakly on internal field intraparticle temperature gradients in comparison with external field interparticle gradients.25 The external field interparticle temperature gradients can be minimized by decreasing the reactor diameter, by diluting the solids, or by diluting the reactant. Due to volume constraints, reactant dilution may not be feasible in commercial reformers. Therefore, it has been suggested that controlling external field interparticle heat transfer in a packed-bed reactor by decreasing the reactor diameter and by solids dilution may prove to be advantageous.26,27 Tubular packed-bed reactors with small diameters have the advantage of providing higher surface-to-volume ratios. This results in decreasing mass transport and heat transport resistances and makes them attractive for achieving efficient thermal and mass transfer rates. Karim et al.27 applied this concept for methanol steam reforming on a commercial CuO/ZnO/Al2O3 catalyst and showed its potential in the area of fuel processing for fuel cell applications. Their study points out that there exist heat transfer limitations in the bed and that the majority of activity measurements that have been reported for methanol steam reforming could be influenced by temperature gradients within the reactor. Catalyst dilution by solids increases the wall area to catalyst volume for heat transfer. It establishes specific temperature profiles (both in the radial and axial directions) in the packedbed reactor and thereby improves the isothermicity while simultaneously suppressing other effects that adversely affect performance (e.g., nature of axial dispersion, channeling loss). The strategy of dilution by solids has been studied to improve packed-bed reactor performance and is now being increasingly applied.17-19,28-31 The cost of catalyst forms a major component in reactor operation, and dilution can help in this aspect. In addition to packed-bed catalytic reactors, catalyst dilution by solids has been applied to fluidized-bed reactors32,33 and tricklebed reactors.34,35 Care should be taken, however, to prevent bypass of flow inside a reactor while employing the dilution strategy.29-31 Catalyst dilution studies reported in the literature use inert solid-catalyst mixtures. In the present study, we assume that the catalyst is diluted with an adsorbent that acts (25) Kirillov, V. A.; Kuzin, N. A.; Mescheryakov, V. D.; Droboshevich, V. I. Chem. Eng. Sci. 2001, 56, 381-386. (26) Mears, D. E. J. Catal. 1971, 20, 127-131. (27) Karim, A.; Bravo, J.; Datye, A. Appl. Catal., A 2005, 282, 101109. (28) Taniewski, M.; Lacjowicz, A.; Skutil, K.; Czechowicz, D.; Ferreria, R. M. Q. Chem. Eng. Sci. 1996, 55, 4271-4278. (29) Hwang, S.; Smith, R. Chem. Eng. Sci. 2004, 59, 4229-4243. (30) Berger, R. J.; Pe´rez-Ramı´rez, J.; Kapteijn, F.; Moulijn, J. A. Chem. Eng. J. 2002, 90, 173-183. (31) van den Bleek, C. M.; van der Wiele, K.; van de Berg, P. J. Chem. Eng. Sci. 1969, 24, 681-694. (32) Irani, R. K.; Kulkarni, B. D.; Doraiswamy, L. K. Ind. Eng. Chem. Process. Des. DeV. 1979, 18, 648-655. (33) Irani, R. K.; Kulkarni, B. D.; Doraiswamy, L. K.; Hussain, S. Z. Ind. Eng. Chem. Process Des. DeV. 1982, 21, 192-195. (34) Van Klinken, J.; Van Dongen, R. H. Chem. Eng. Sci. 1980, 35, 59-66. (35) Al-Dahnan, M. H.; Dudukovic, M. P. AIChE J. 1996, 42, 25942606.
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not only as a diluent for the catalyst but also at the same time as an adsorbent for CO2. Steam methane reforming (SMR) reactions are equilibrium-controlled, and selective removal of CO2 from the reaction zone increases methane conversion according to Le Chatelier’s principle. Hydrotalcite adsorbent, which selectively removes CO2 from the reaction zone and is stable under wet gas and high temperature, is particularly known to be suitable for the SMR process.36,37 Temperature gradients in SMR in a packed bed can adversely affect process performance, and this article focuses on the use of tubular packed-bed microreactors (TPBM) and employs a catalyst dilution strategy for performance enhancement of SMR for controlling the temperature gradient in the bed. As discussed above, these two aspects of reactor design (viz., microreactor operation and catalyst dilution by solids) have generally been studied separately. In the present work, we consider a strategy where the above two features can take place simultaneously. For the above objective, we have studied the SMR process performance in TPBM for diluted and undiluted conditions using a dynamic 2D pseudohomogeneous model. The model incorporates mass and energy balances with nonlinear adsorption isotherm together with a linear-driving force (LDF) model for adsorption rate and this is coupled with the SMR reactions. The results of simulation are presented and discussed in detail in section 3 and show the improvement in SMR performance pertaining to effects of varying reactor diameter, extent of catalyst dilution by adsorbent, and solids packing density. 2. 2D Mathematical Model for the SMR Process The key chemical reactions of the SMR process are given by:
CH4 + H2O S CO + 3H2, ∆H298 ) 206.2 kJ/mol (1) CH4 + 2H2O S CO2 + 4H2, ∆H298 ) 164.9 kJ/mol (2) CO + H2O S CO2 + H2, ∆H298 ) -41.1 kJ/mol (3) The mathematical model used to describe the SMR reactor with adsorbent dilution is a dynamic 2D pseudohomogeneous model that considers the nonisothermal, nonadiabatic, and nonisobaric nature of operation. The model assumptions used are: 1. Axial dispersed plug flow prevails in the bed. 2. Mass dispersion in the axial direction is considered. 3. Mass dispersion in the radial direction is assumed to be negligible. 4. The system is nonisothermal. Thermal dispersion in both axial and radial directions is considered. 5. The reaction kinetic model employed is that proposed by Xu and Froment.38 This kinetic model has been shown to be useful to study the SMR process.39-41 6. Volumetric change of flow due to adsorption and reaction is taken into account in the overall material balance. 7. The gas is assumed to be an ideal gas. 8. The adsorbent and catalyst particles are the same size and spherical in shape. (36) Zou, Y.; Mata, V.; Rodrigues, A. E. Ind. Eng. Chem. Res. 2001, 40, 204-209. (37) Ding, Y.; Alpay, E. Chem. Eng. Sci. 2000, 55, 3461-3474. (38) Xu, J.; Froment, G. F. AIChE J. 1989, 35, 88-96. (39) Ding, Y.; Alpay, E. Chem. Eng. Sci. 2000, 55, 3929-3940. (40) Xiu, G. H.; Soarse, J. L.; Rodrigues, A. E. AIChE J. 2002, 48, 28172832. (41) Xiu, G. H.; Li, P.; Rodrigues, A. E. Chem. Eng. J. 2003, 95, 8393.
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9. The pressure distribution in the packed bed adsorptive reactor is described by the Ergun equation.42 10. The effect of changing bed voidage with variation of reactor-to-particle diameter ratio43 on methane conversion, hydrogen purity, and in the Ergun equation has been considered. 11. The gas phase and the catalyst/adsorbent particles are assumed to be in local mass/thermal equilibrium at all times. 12. Five chemical species are considered, viz., CH4, H2O, CO, H2, and CO2. 13. The nonlinear Langmuir model is used to describe the multicomponent adsorption equilibrium isotherm for CO2. 14. The LDF model is used for the adsorption rate mechanism. In the present study, hydrotalcite37 is used as adsorbent for CO2 and the Langmuir isotherm model, and the LDF rate model was found to give an adequate description of CO2 adsorption and desorption behavior. 15. The catalyst is assumed to be uniform and small in size, so that the effectiveness factor is considered to be unity and is the same for all reactions. 16. External field interparticle heat transfer effects (both radial and axial temperature gradients in the framework of packed-bed reactor taken as a whole) in the reactor are considered. 17. There is complete mixing of the adsorbent and catalyst particles with negligible catalyst deactivation. 18. The feed stream and column wall are maintained at the same constant temperature. For the above assumptions, the governing equations with initial and boundary conditions are summarized below. 1. The overall mass balance equation:
∂C
t
+
∂(uC)
∂t
∂z
n
+ Fad
∂qji
n III
- Fcat∑∑υijηjRj ) 0 ∑ i)1 ∂t i)1 j)I
(4)
t
∂t
+
∂(uCi) ∂z
+ Fad
∂qji ∂t
III
- Fcat
∂
( ) ∂Ci
υijηjRj ) b DL ∑ ∂z ∂z j)I
(5)
where DL is the axial dispersion coefficient and is evaluated by the correlation given by Edwards and Richardson.44 3. Mass transfer rates:
∂qji ) kfi(q/i - qji) ∂t
(6)
where kfi is the LDF mass transfer coefficient and may be evaluated by the correlation reported by Ding and Alpay.37 4. Adsorption equilibrium:
q/i )
mibiPi n
1+
n
Fad
(
∂qji
-∆Hadi ∑ ∂t i)1
)
∂t
∂T
+ CCpgu
-
∂z
n III
- Fcat
∑ ∑υijηjRj∆HRj ) i)1 j)I
( ) ( )
1 ∂ ∂T ∂T kz + krr (8) ∂z ∂z r ∂r ∂r ∂
Because the effective axial and radial thermal conductivities do not vary by orders of magnitude, we choose for purpose of simulation studies kr ) kz and use an estimated value obtained from the correlation given by Yagi et al.45 For the simulation of the undiluted bed, the adsorbent parameters (mi, bi, kfi, ∆Hadi,Fad) are zero-valued. 6. Initial conditions: The initial conditions (at t ) 0) used in the present study are as follows:
T ) Tf, qji ) 0, PH2 ) Pf, CH2 )
PH2
, Pi ) 0, RTf Ci ) 0 (i ) CH4, H2O, CO, CO2) (9)
where Tf and Pf are feed gas temperature and pressure, respectively. 7. Boundary conditions: The following boundary conditions are used in the simulations (i) Reactor inlet (z ) 0, 0 e r e R0)
-bDL
( )
∂Ci ) uf(Cfi - Ci) ∂z
(∂T∂z ) ) u CC
-kz
f
pg(Tf
- T)
u ) u f, P ) P f
2. The component mass balance for component i in the gas phase:
∂Ci
∂T
[tCCvg + (Fad + Fcat)Cps]
(7)
bjPj ∑ j)1
5. The energy balance for the bed-volume element: (42) Ergun, S. Chem. Eng. Sci. 1952, 48, 89-94. (43) de Klerk, A. AIChE J. 2003, 49, 2022-2029. (44) Edwards, M. H.; Richardson, J. F. Chem. Eng. Sci. 1968, 23, 109123.
(10a) (10b) (10c)
(ii) Reactor outlet (z ) L, 0 e r e R0)
∂Ci ∂T ) 0, )0 ∂z ∂z
(10d)
∂u ∂P ) 0, )0 ∂z ∂z
(10e)
(iii) At the center of reactor (r ) 0, z g 0)
∂Ci ∂T ) 0, )0 ∂r ∂r
(10f)
∂u ∂P ) 0, )0 ∂r ∂r
(10g)
(iv) At the reactor wall (r ) R0, z g 0)
∂Ci )0 ∂r -kr
∂T ) hw(Tw - T) ∂r
∂u ∂P ) 0, )0 ∂r ∂r
(10h) (10i) (10j)
where Tw is the wall temperature (assumed to be equal to Tf for purposes of simulation) and hw is the wall heat transfer coefficient. (45) Yagi, S.; Kunii, D.; Wakao, N. AIChE J. 1960, 6, 543-546.
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Table 1. Parameter Values Used in the Simulations parameter a
bco2 dpb Cpga Cpsa ∆HadCO2a mco2a Pfc hwb ufc Tfc H2O/CH4b bb pb tb µa Fadb Fcatb ηb a
value diluted bed: 2.36 × 10-4 Pa-1; undiluted bed: 0 5 × 10-4 m 42 J/mol K 850 J/kg K diluted bed: -17000 J/mol; undiluted bed: 0 diluted bed: 0.65 mol/kg; undiluted bed: 0 445.7 kPa 71 J/m2‚K 0.008 m/s 773 K 6 0.48 0.24 0.64 2.87 × 10-5 Pa-s diluted bed: 498 kg/m3, undiluted bed: 0 249 kg/m3 1.0
Data from refs 37 and 39. b Data from ref 40. c Present study.
Figure 2. Radial temperature profiles at z ) 0.03 m (where the temperature is the highest) for a 6-mm-diameter reactor at different feed temperatures. Values of other parameters are given in Table 1.
Figure 1. Axial temperature profiles at the center (r ) 0.0) for a 6-mmdiameter reactor at different feed temperatures. Values of other parameters are given in Table 1.
de Klerk43 has reported a model that obtains a good description of the influence of column-to-particle diameter ratio (D/dp) on the average bed voidage (b). This model is valid for equally sized spheres and for D/dp > 2. In the present study, we consider D/dp ) 4-50 with catalyst and adsorbent of uniform spherical particles of the same diameter. The above model may therefore be used in the present study for calculating the bed voidage with D/dp and is given by
(
D b ) i + 0.35 exp -0.39 dp
)
(11)
where i is defined as the average bed voidage of an infinite diameter column. de Klerk43 showed that the effect of D/dp on b is more pronounced for small values of D/dp (15). In the present study, the value of average voidage for the infinite diameter column is assumed to be the situation when D/dp ) 5040 (i.e., i ) 0.48). The average bed voidage b can then be calculated by using eq 11. The effect of change of bed voidage with D/dp has therefore been incorporated in the SMR model formulation. 3. Results and Discussion It may be noted that the model considered is sufficiently rigorous for analysis of SMR system behavior and is based on
Figure 3. Effect of steady-state SMR performance as a function of reactor diameter. Values of other parameters are given in Table 1.
a first principles approach. It involves both reaction and adsorptive separations occurring simultaneously. The adsorption equilibrium isotherm is nonlinear, and the values of its constants depend on temperature and pressure. Also, the above equations are highly coupled and nonlinear, and this can even cause steep axial composition gradients in the reactor depending on parameter values and operating conditions. Equations 4-8 with initial and boundary conditions (eqs 9 and 10) are numerically solved by a finite difference formulation, and the solutions obtained predict the SMR process performance for undiluted and diluted bed conditions. For simulation purposes the finite difference formulation employed a sufficient number of grid points along the reactor length and radius so that numerically stable and accurate solutions that describe the system dynamics in time are obtained. The overall mass balance equation (eq 6) and Ergun equation42 are solved to obtain the velocity and pressure profiles in the reactor along with eqs 5 and 8 for the effluent mole fractions and reactor temperature profiles. The process performance is then quantitatively compared by studying the conversion of methane XCH4 and hydrogen purity yH2(dry basis) defined as, respectively,40
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XCH4 )
Energy & Fuels, Vol. 20, No. 2, 2006 467
Feed of CH4 (mol/s) - Effluent of CH4 (mol/s)
[ (
) 1-
Feed of CH4 (mol/s)
)( )
RTf ufPfyCH4
uPyCH4
feed
RT
yH2(dry basis) )
]
outlet
y H2 1 - yH2O
(12)
(13)
Temperature Gradients in a Tubular Packed-Bed Microreactor. To check whether the SMR performance is affected by the existence of temperature gradients, simulation studies for a 6-mm diameter packed-bed reactor for varying feed temperature (Tf) was first carried out. The catalyst loading (Wc) was kept constant at 6.45 g. The values of other parameters used in the simulation studies are listed in Table 1. The steady
state temperature profiles were obtained by solving the dynamic model for a sufficiently long time. The steady-state axial temperature profiles at the center (at r ) 0.0 m) and the radial temperature profiles where the temperature is the highest at z ) 0.03 m are plotted in Figures 1 and 2, respectively. For varying Tf, Figure 1 shows that steep axial temperature gradients exist that move further downstream for higher Tf. With increase in Tf, higher values of conversion of methane (XCH4) are possible. In turn, this leads to higher heat removal rates by reaction, and consequently, steeper temperature gradients are observed near the reactor inlet. Figure 2, however, shows that even for a 6-mmdiameter reactor radial temperature gradients can exist due to interparticle heat transfer, and this becomes more dominant at higher Tf. Mears26 developed a criterion to assess the presence of radial heat transfer resistance by evaluating R0/RP. According to the criterion, if R0/RP > 100 then the resistance may be neglected. In our study, R0/RP ) 12, and the results obtained in
Figure 4. Temperature profiles: (a) d ) 2 mm and (b) d ) 15 mm. Values of other parameters are given in Table 1.
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Figure 5. Velocity profiles: (a) d ) 2 mm and (b) d ) 15 mm. Values of other parameters are given in Table 1.
Figures 1 and 2 showing the presence of heat transfer limitations are supported by the Mears criterion. The above results suggest that the SMR reactor diameter is an important parameter that affects external field interparticle heat transport mechanisms. We therefore studied the effects of decreasing reactor diameter on performance, and the results obtained are discussed next. SMR Process Performance in Tubular Packed-Bed Microreactors. Studies are carried out by reducing the reactor diameter (d) systematically from 25 mm to 2 mm, keeping a catalyst loading of 6.45 g. The reaction was assumed to be carried out at Tf ) 773 K. The other parameter values are chosen as given in Table 1. The steady-state process performance in terms of methane conversion (XCH4) and H2 purity (yH2) is plotted in Figure 3. As seen, there is a significant improvement in the process performance with decrease in d. For decreasing d, the reactor length L needs to be altered to maintain constant catalyst loading. To assess the results, the needed reactor length with decreasing reactor diameter is also shown in Figure 3. For a catalyst loading of 6.45 g, the methane
conversion increases from 20 to 100% upon reducing d from 25 to 2.8 mm. A reactor with d ) 2.8 mm with the corresponding length can therefore be designed for full methane conversion. It is seen that, even at lower d, process miniaturization by appropriate design and integration is feasible for the calculated values of reactor length and chosen diameter. Figure 4 compares the axial and radial temperature profiles plotted for reactors having d ) 2 mm and (Figure 4a) and 15 mm (Figure 4b). It is clearly seen that the temperature gradients are controlled for d ) 2 mm when compared to d ) 15 mm, especially in the radial direction. It has been reported that the advantages resulting from isothermality by reducing the reactor diameter are significantly higher than the effects of channeling for R0/RP ratios as low as even 4.26,31 In the present study, for a 2-mm-diameter reactor where isothermality is approached, R0/ RP ratio is 4. This suggests that choosing an appropriate tubular microreactor diameter is possible with negligible channeling effects.
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Energy & Fuels, Vol. 20, No. 2, 2006 469
Figure 6. Pressure profiles: (a) d ) 2 mm and (b) d ) 15 mm. Values of other parameters are given in Table 1.
We also studied the effects of velocity change and pressure drop due to adsorption and reaction in the reactor. The nature of the velocity profiles obtained is illustrated in Figure 5 for d ) 2 mm (Figure 5a) and d ) 15 mm (Figure 5b). The results indicate that, for the similar superficial velocities, smaller-diameter reactors can provide good radial flow features in the extended reaction zone. Hence, it may now be possible to control the hydrodynamics and thereby the reaction behavior in the reaction zone. Likewise, the pressure drop for the lower d ) 2 mm reactor is in the range of 2-3 kPa, does not seems to be significant. It may also be seen in Figure 6 that the magnitudes and nature of pressure profiles for both d ) 2 mm and d ) 15 mm reactors are not considerably different. This is again advantageous from design considerations as the effects of pressure drop may not be important for the SMR process taking place in a TPBM.
The dynamics of the SMR process in time at the exit of the reactor is illustrated in Figure 7a,b in terms of XCH4 and yH2 for d ) 2, 4, 10, and 25 mm. It is observed from Figure 7a,b that the transient period is longer as the reactor diameter becomes smaller from 25-mm to 4-mm reactors. For a 2-mm-diameter reactor, 100% conversion at the reactor outlet is seen throughout the operation of the process. This suggests that, for a TPBM with lower diameter, a lower Wc and therefore of shorter length would suffice for achieving complete conversion. Our calculation showed that a 2-mm-diameter reactor and length corresponding to a Wc of 5.5 g yield 100% conversion. Hence, operating with TPBM will help in running the SMR process at low Wc than conventionally employed larger-diameter reactors for the same performance parameters. Effect of Adsorbent Dilution on the SMR Process. Simulation for effects of catalyst dilution is carried out by assuming hydrotalcite as the diluent that also acts as adsorbent for CO2
470 Energy & Fuels, Vol. 20, No. 2, 2006
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Figure 7. Transient profiles of SMR performance as a function of reactor diameter at the exit: (a) Methane conversion (XCH4) and (b) hydrogen purity (yH2). Values of other parameters are given in Table 1.
Figure 8. Effect of adsorbent dilution on SMR performance as a function of reactor diameter. Constant bulk packing density of (a) methane conversion (XCH4) and (b) hydrogen purity (yH2). Solid line: undiluted reactor. Dashed line: diluted reactor. Values of other parameters are given in Table 1.
in the reaction zone. The equilibrium and kinetic constants for hydrotalcite adsorbent have been taken from Ding and Alpay.37,39 The effects of adsorbent dilution are investigated in two ways: (1) by comparing the performance of dynamics of diluted and undiluted TPBM of varying diameters having a fixed catalyst loading (Wc) and (2) by studying the effects of dilution as a function of adsorbent loading (Wa) in a TPBM for varying bulk packing density Fb with constant catalyst loading. Microreactor Dynamics with Dilution. In carrying out the studies, we assume Wc ) 6.25 g and Wa ) 12.5 g for the diluted bed. The bulk packing density Fb(Fb ) Fcat + Fad) is assumed to be 747 kg/m3. The process performance in time at the outlet of the diluted beds in terms of XCH4 and yH2 is presented, respectively, in Figure 8a,b for varying d. The performance of
the undiluted beds for the same catalyst loading, Wc ) 6.25 g (with Wa ) 0.0 g) and Fb(Fb ) Fcat) equal to 249 kg/m3, are also given in Figure 8a,b for comparison. The solid lines are for the undiluted systems, while the dashed lines are for the diluted reactors. It is clearly seen that for smaller-diameter reactors (e.g., Figure 8a, 4-mm curve) the effects of dilution are much more pronounced than those for the larger-diameter reactor (e.g., Figure 8a, 25-mm curve). Interestingly, for smallerdiameter tubular microreactors (e.g., Figure 8a, 4-mm curve) the time taken to reach the steady state is far greater than the larger-diameter reactor (e.g., Figure 8a, 25 mm). This implies that higher conversions are possible for longer time duration for the diluted bed when compared to the undiluted bed. The radial temperature drop due to the endothermic nature of SMR
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Figure 9. Effect of adsorbent dilution on SMR performance for a 3-mm-diameter reactor: Varying bulk packing density. Values of other parameters are given in Table 1.
is more uniform in smaller-diameter reactors with dilution as seen in Figure 4. The average temperature during the initial transient regime (i.e., before the system reaches the steady state)in the smaller-diameter reactor is lower when compared to the larger-diameter reactor. Adsorption is favored at lower temperature, and this enhances the performance of adsorbent and thereby decreases the CO2 concentration in the reaction zone. This decreases the bulk-phase CO2 concentration, while increasing the temperature due to the exothermic nature of adsorption. These effects are favorable for steam reforming reactions given by eqs 1 and 2 and allow an optimal balance between reaction and adsorption. Saturation of adsorbent by adsorbate occurs in the case of catalyst dilution by adsorbent, and regeneration of adsorbent is necessary. This is usually carried out by cyclic operations such as pressure swing adsorption (PSA), temperature swing adsorption (TSA), or a combination of the above.46-50 The longer transient time observed in the case of smaller TPBM permits longer cycle time for the cyclic operation of the adsorption-enhanced SMR process. This is an added advantage in terms of operating costs. Effects of Dilution for Varying Bulk Packing Density with Constant Catalyst Loading. To get insights into the effects of increasing Wa, we carried out studies by increasing the adsorbent loading but with a constant Wc ) 6.25 g. For this, the bulk packing density needs to be altered to keep constant Wc. We now define a dilution ratio parameter, B, as
B)
Vad (Vad + Vc)
(14)
analogous to dilution with inert solids,26 where Vad is the volume of adsorbent and Vc is the volume of catalyst. The steady-state performance for 3-mm-diameter TPBM is shown in Figure 9 for varying dilution ratio B. It is evident from Figure 9 that there is an influence of dilution effect for B > 0.25, and this enhances with increase in B with complete (46) Rajasree, R.; Moharir, A. S. Comput. Chem. Eng. 2000, 24, 24932505. (47) Ruthven, D. M. Principles of Adsorption and Adsorption Processes; Wiley & Sons: New York, 1984. (48) Ruthven, D. M.; Farooq, S.; Knaebel, K. S. Pressure Swing Adsorption; VCH Publishers: New York, 1994. (49) Yang, R. T. Gas Separation by Adsorption Processes; Butterworths: Boston, 1987. (50) Tien, C. Adsorption Calculations and Modeling; Butterworth: Boston, 1994.
conversion of methane observed for B ) 0.8. However, a longer reactor length is now required for accommodating the increased Wa (also shown in Figure 9). The proper choice of B and reactor diameter can therefore help in choosing a reactor of optimal length. In summary, the results obtained on studying the SMR reaction in TPBM show reduction in thermal gradients in the reactor. This makes control of the temperature the key operating parameter for the endothermic SMR reaction and thereby enhances the process performance in terms of methane conversion and hydrogen purity. The advantage of lower average temperature for smaller-diameter columns during the initial transient regime is shown to bring about further enhancement in SMR process performance for the case of dilution by adsorbent as adsorption is inherently exothermic in nature. The results obtained also indicate that the effects of pressure drop due to the use of TPBM may be controlled for the SMR reaction. The known advantages of using (a) tubular packed-bed microreactors and (b) catalyst dilution in reactor operation (as discussed in the Introduction) have been amalgamated in this study to show how their combined presence can simultaneously improve SMR reactor performance. 4. Conclusions Studies with a dynamic 2D pseudohomogeneous process model have shown that TPBM and catalyst dilution by adsorbent enhance the SMR process. Decreasing the reactor diameter significantly increases the SMR process performance, and complete conversion of methane is attainable. Also, for the same process performance, TPBM allows running the SMR process at a catalyst loading lower than that of larger-diameter reactors. Analyzing dilution of catalyst by adsorbent material in TPBM while maintaining constant bulk packing density enhances the SMR performance in the adsorption-enhanced reaction zone. This enhancement increases further with decreasing reactor diameter and also permits longer cycle time for the cyclic operation of the adsorption-enhanced SMR process. It therefore offers an added advantage in terms of operating costs. Adsorbent dilution by varying bulk packing density also has a very significant effect on SMR process performance. This study suggests that, by simultaneous optimization of reactor diameter and adsorbent loading, reduction in process costs and improving the SMR performance may be possible. It may be noted that for tubular packed-bed microreactors, the effects of higher pressure drop (although it may be controlled for the SMR reaction as seen in the present study), potential blockage, reactor packing, productivity, effects of nonuniform flow field due to voidage effects at low particle diameters, internal field intraparticle gradients, and so forth need to be analyzed in detail. The development of more sophisticated models involving flow in disordered packed beds and relating both microscopic and macroscopic properties using networked models are likely to lead to more detailed information.51 The basic trends in the simulation results obtained here from first principles modeling of tubular packed-bed microreactors with adsorbent dilution for the SMR reaction are encouraging, and this strategy needs to be validated by carrying out detailed experiments. Acknowledgment. R.R. acknowledges the Council of Scientific and Industrial Research (CSIR), New Delhi, for financial support. (51) Thompson, K. E.; Fogler, S. AIChE J. 1997, 43, 1377-1389.
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Nomenclature B ) Dilution ratio bi ) Langmuir model constant for component i, Pa-1 C ) Total molar concentration in the bulk phase, mol/m3 Cfi ) Molar concentration of component i in the feed, mol/m3 Ci ) Molar concentration of component i, mol/m3 Cpg ) Gas-phase heat capacity, J/mol‚K Cps ) Solid-phase heat capacity, J/kg‚K Cvg ) Gas-phase heat capacity at constant volume, J/mol‚K d ) Reactor diameter dp ) Particle diameter D ) Diameter of the reactor, m DL ) Axial dispersion coefficient, m2/s hw ) Wall heat transfer coefficient J/m2‚s‚K kfi ) LDF mass-transfer coefficient of component i, s-1 kr ) Effective radial thermal conductivity, J/m‚s‚K kz ) Effective axial thermal conductivity, J/m‚s‚K L ) Reactor length, m mi ) Langmuir constant for component i, mol/kg P ) Local total pressure, kPa Pf ) Feed pressure, kPa Pi ) Partial pressure of gas-phase component i, kPa q* ) Equilibrium solid-phase concentration, mol/kg qji ) Solid-phase concentration for component i (averaged over an adsorbent particle), mol/kg r ) Radial coordinate in the reactor, m R ) Universal gas constant, J/mol‚K R0 ) Inner radius of the reactor, m
Rajasree et al. Rp ) Radius of the particle, m t ) Time, s T ) Temperature in bulk gas phase, K Tf ) Feed gas temperature, K Tw ) Wall temperature, K u ) Superficial velocity, m/s uf ) Initial superficial velocity, m/s Vad ) Volume of adsorbent, m3 Vc ) Volume of catalyst, m3 Wa ) Adsorbent loading, kg Wc ) Catalyst loading, kg z ) Axial coordinate in the reactor, m Greek Letters b ) Bed voidage, dimensionless i ) Average bed voidage of an infinite diameter column P ) Particle porosity, dimensionless t ) Total bed porosity, dimensionless ηi ) Catalyst effectiveness factor, dimensionless Fad ) Mass of adsorbent per bed volume, kg/m3 Fb ) Bulk packing density of the adsorbent and catalyst, kg/m3 Fcat ) Mass of catalyst per bed volume, kg/m3 µ ) Viscosity of fluid, kg/m‚s υij ) Stoichiometric coefficient of component i in reaction j, dimensionless -∆Hadi ) Adsorption heat of component i, J/mol ∆HRi ) Reaction heat of reaction i, J/mol EF050205Z
