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Ind. Eng. Chem. Res. 2009, 48, 1090–1095
Simulation of a Membrane Reactor for the Catalytic Oxidehydrogenation of Ethane Marı´a L. Rodriguez,*,† Daniel E. Ardissone,‡ Angeliki A. Lemonidou,§ Eleni Heracleous,§ Eduardo Lo´pez,†,| Marisa N. Pedernera,† and Daniel O. Borio† PLAPIQUI (UNS-CONICET), Camino La Carrindanga, km.7, 8000 Bahı´a Blanca, Argentina, FICES (UNSL), AVenida 25 de Mayo 384, 5730 Villa Mercedes (San Luis), Argentina, Chemical Process Engineering Research Institute (CERTH/CPERI), Department of Chemical Engineering, Aristotle UniVersity of Thessaloniki, P.O. Box 1517, UniVersity Campus, GR-54006 Thessaloniki, Greece, and Institut de Te`cniques Energe`tiques, UniVersitat Polite`cnica de Catalunya, AVenida Diagonal 647, Ed. ETSEIB, 08028 Barcelona, Spain
Industrial-scale ethylene production using a novel membrane multitubular reactor for the ethane oxidative dehydrogenation process over a Ni-Nb-O catalyst is proposed. A theoretical study was performed by means of a pseudohomogeneous model of the tube and shell sides. The feasibility and convenience of using this novel design, as well as the influence of the main operating variables on the reactor performance, were analyzed. The introduction of the membrane leads to lower oxygen partial pressures inside the catalyst tubes when compared with a conventional multitubular reactor. This leads to very good ethylene selectivities, good temperature control as a result of lower heat generation rates, and reasonable production rates. The reactor performance appears to be highly affected by the balance between the rate of oxygen consumption by the chemical reaction and its rate of permeation through the membrane. Under certain operating conditions leading to lower reaction rates, an undesired accumulation of oxygen inside the tubes is observed. A minimum amount of O2 injected at the tube mouth appears beneficial to overcoming this accumulation phenomenon. The membrane reactor shows a nonconventional inverse parametric sensitivity with respect to the inlet temperature. When the reactor is operated at conditions where the reaction is controlled by the permeation flow of O2 through the membrane, it is possible to reach high selectivities to ethylene, significant ethane conversions, and mild temperature profiles. Introduction Ethane oxidative dehydrogenation (ODH) appears to be the most promising alternative to the conventional process of thermal cracking (pyrolysis) of naphtha or ethane for industrial ethylene production. The ODH reaction is carried out in the presence of a suitable catalyst, competing with the undesired combustion of ethane and ethylene. In contrast to the highly endothermic and equilibrium-limited thermal cracking process, in the ODH process, the reactions are irreversible and exothermic. Ethane ODH has been studied over a wide range of catalysts.1,2 Recently, Heracleous and Lemonidou3,4 reported the development of bulk Ni-Nb-O mixed oxides as both highly active and selective catalytic materials for the desired reaction. Moreover, the only byproduct formed over these mixed oxide catalysts is CO2. The lack of CO production significantly simplifies the product separation downstream from the reactor, with all the associated economic benefits. The system of reactions taking place over the Ni-Nb-O mixed oxides is represented by reactions 1-3 1 C2H6 + O2 f C2H4 + H2O 2
(1)
7 C2H6 + O2 f 2CO2 + 3H2O 2
(2)
C2H4 + 3O2 f 2CO2 + 2H2O
(3)
* To whom correspondence should be addressed. E-mail:
[email protected]. † PLAPIQUI (UNS-CONICET). ‡ FICES (UNSL). § Aristotle University of Thessaloniki. | Universitat Polite`cnica de Catalunya.
Multitubular reactors are commonly used in industry to carry out these exothermic processes, with control of the reaction temperature being a key factor for maintaining good selectivity levels.5 An attractive alternative to these fixed-bed reactors is the use of a membrane reactor to axially distribute the oxygen injection in the reaction media (see Figure 1). As reported by Heracleous and Lemonidou,3,4 lower oxygen partial pressures increase the selectivity, decrease the generated heat, and increase the ethylene yield. The oxygen distribution also allows better heat management, as it leads to a decrease of the local reaction rates and, consequently, of the rate of heat generation. Furthermore, in the case of flammable mixtures (as in this case), the oxygen distribution through the membrane prevents the local generation of dangerous concentrations and allows operation under normally prohibitive concentrations in the reaction mixture. To this end, an inorganic porous membrane was selected. This membrane is not permselective to oxygen but presents higher permeation fluxes than dense membranes. Dense membranes are completely selective to oxygen; however, the permeation fluxes provided are insufficient for the ignition of the reaction in the present study. Moreover, dense membranes show limited thermal and chemical stability upon aging.
Figure 1. Scheme of the ethane ODH membrane reactor.
10.1021/ie800564v CCC: $40.75 2009 American Chemical Society Published on Web 12/19/2008
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In the present contribution, a theoretical study of a multitubular membrane reactor for the ODH of ethane to ethylene over a Ni-Nb-O mixed oxide catalyst is presented. By means of a pseudohomogeneous model, the feasibility and convenience of using this novel design, as well as the influence of the main operating variables on the reactor performance, were analyzed.
dTS UπdTnT(T - TS) ) 2 dz FSj Cpj
∑ j)1
where the permeation flux is quantified by the expression7
(
Mathematical Model A one-dimensional, pseudohomogeneous, steady-state model was used to represent the ODH of ethane to ethylene in a multitubular packed-bed membrane reactor. Based on operation with high total flow rates through the packed bed, axial mass and energy dispersions and external transport limitations were assumed to be negligible. Internal mass- and energy-transport limitations were also not considered based on the use of a thin washcoat over the catalyst spheres (eggshell type). The use of tubes of small diameter (1 in.) supports the assumption of relatively flat mass and temperature radial profiles. The reactor shell was assumed to be adiabatic, and isobaric conditions were selected for the air flowing on the shell side. The friction factor f proposed by Ergun was employed to predict the pressure drop inside the tubes. The catalyst particles were assumed to be 6-mm-diameter spheres of the eggshell type. The washcoat covering the spheres, where the active components were impregnated, was assumed to have a mean thickness of 177 µm.. The use of washcoated particles resulted in a low bed density of FB ) 200 kgcat/m3, favoring the moderation of the heat generation rate per unit volume and, consequently, achieving milder operating conditions. The inert membrane tubes were filled with catalyst particles, and oxygen was dosed from the shell side as air through the membrane into the tubes. The model is represented by the following equations: Reaction Side (Catalyst Tubes) Mass Balance 3 dFj ) ATFB νijri dz i)1
∑
j ) C2H6,C2H4,CO2,H2O
3 dFj ) ATFB νijri + JjπdT dz i)1
∑
j ) O2,N2
(4)
(5)
Energy Balance
∑
∑
j)1
(6)
Momentum Equation
Dej,k ) K0
)
j ) O2,N2 (10)
8RTm πMj Fj ) Fj0 FSj ) FSj0 T ) T0 TS ) TS0 P ) P0
(11)
The power-law kinetic model proposed by Heracleous and Lemonidou4 for reactions 1-3 over a Ni0.85Nb0.15O catalyst (see eq 12) was used for the simulations. Table 1 reports the corresponding parameters for ri (TR ) 543 K, pj in Pa).
[ ( )]
ri ) kR,i exp -
Ei 1 1 p aip bi R T TR HCi O2
(12)
The overall heat-transfer coefficient (U) was evaluated using the guidelines suggested by Froment and Bischoff.6 The heattransfer coefficient for the process gas side was calculated by means of the equation of Leva,8 and that corresponding to the shell side was calculated using the guidelines reported by Kern.9 To avoid the multiplicity of steady states inherent to countercurrent operation, a cocurrent flow configuration between the process gas and the coolant (air) was selected. Additionally, the cocurrent scheme lowers the hot spots, which leads to longer catalyst life, and reduces the parametric sensitivity, which ensures safe operating conditions.10 The geometrical parameters and main operating conditions are listed in Table 2. An inert ceramic porous membrane was selected for this work. Its geometric parameters11 and thermal conductivity12 are listed in Table 3. The reactor was assumed to be cooled both by convective heat exchange between reactants and coolant and by cold-shot of air permeated through the membrane.
1 2 3
kR,i (kmol kgcat-1 s-1 Pa-(a+b)) -10
4.177 × 10 1.272 × 10-13 9.367 × 10-11
Ei (J/mol)
ai
bi
96 180 76 210 98 420
0.520 0.547 0.475
0.213 0.829 0.319
Table 2. Geometrical Parameters and Operating Conditions
fFgus2 dP )dz dp
(7)
Permeation Side (Shell) Mass Balance dFSj ) -JjπdTnT dz Energy Balance
B0 Dej,k + P RTmδ RTmδµi m
The initial conditions at z ) 0 are given by
reaction
j)1
JO2CpO2) + U]
Jj ) ∆P
Table 1. Kinetic Parameters
ATFB 3 4 (T - TS) dT ) 6 ri(-∆Hri) [(JN2CpN2 + 6 dz d T i)1 FjCpj FjCpj
∑
(9)
j ) O2,N2
(8)
parameter
value
tube length, L internal tube diameter, dti shell diameter, dS number of tubes, nt feed pressure, P0 shell-side pressure, PS total molar feed flow rate, FT0 bed density, FB S air molar flow rate, FT0 tube-side inlet temperature, T0 inlet air temperature (shell side), TS0
4m 2.54 cm 3.93 m 12 000 5 atm 6 atm 8400 kmol/h 200 kg/m3 20 000 kmol/h 340, 360, 380 °C 25, 100, 200 °C
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Figure 2. Tube-side (solid lines) and shell-side (dashed lines) temperatures vs axial position for T0 ) 340, 360, and 380 °C. Inlet O2 ) 1%, TS0 ) 25 °C. Table 3. Geometric Parameters and Thermal Conductivity of the Membrane11,12 parameter
value
K0 × 10 (m) B0 × 1017 (m2) δ (m) ka [J/(s m °C)]
23.1 11.5 0.00175 16.06
10
a
At TM ) 300 °C.
The global selectivity (SG) was calculated as the ratio between the amount of ethylene produced and the amount of ethane consumed, from the reactor inlet to the desired axial coordinate. The instantaneous selectivity (SI), evaluated as the ratio of the local net ethylene production rate to the net ethane consumption rate, represents the actual situation in a particular point of the reactor. Global selectivity and instantaneous selectivity were calculated using the expressions SG )
FC2H4,z - FC2H4,0 FC2H6,0 - FC2H6,z
SI )
r1 - r3 r1 + r2
(13)
(14)
Results and Discussion Influence of the Inlet Temperature on the Tube Side (T0). Figure 2 shows the temperature evolution along the reactor length, for three different inlet temperatures at the permeate side (T0). To favor the reaction ignition, a low amount of oxygen (1%) was also fed to the reactor mouth. It is clear that, for T0 ) 340 °C, the reaction rates were still very low; therefore, the process stream was overcooled by the gas being circulated through the shell. Moreover, an accumulation of O2 along the membrane tubes was detected (see Figure 3); i.e., for T0 ) 340 °C, the mass-transfer rate of O2 by permeation through the membrane was higher than the rate of oxygen consumption by the chemical reaction. An analogous accumulation behavior was observed for T0 ) 360 °C. However, the temperature was high enough to ignite the reactor at z ≈ 2 m. For these conditions, a sudden temperature increase occurred (Figure 2), and the oxygen profile inside the tubes showed a maximum and then decreased to very low values (see Figure 3 for z > 2.4 m). Finally, when T0 ) 380 °C was selected,
Figure 3. Axial profiles of oxygen molar flow (FO2) and global selectivity (SG) for the conditions of Figure 2.
the accumulation behavior disappeared, and the oxygen molar flow decreased continuously along the tube length. Figure 3 also includes the global selectivity to ethylene (SG) along the reactor for the same operating conditions. The highest outlet selectivity corresponds to the case T0 ) 380 °C, for which the average partial pressure of O2 along the catalyst tube was lowest. This result is consistent with the experimental evidence reported by Heracleous and Lemonidou.3,4 Conversely, for lower feed temperatures, the accumulation of oxygen along the reactor tube caused the selectivity to deteriorate and could cause undesired thermal effects. It is interesting to note that the maximum reactor temperature for T0 ) 380 °C was lower than that corresponding to T0 ) 360 °C (Figure 2). This kind of “inverse” parametric sensitivity with respect to the feed temperature, unusual for conventional multitubular reactors, is a distinctive characteristic of this type of membrane reactor. It is clear that, for both T0 ) 360 and T0 ) 380 °C, the outlet oxygen conversion is almost complete. That is, approximately the same total amount of the limiting reactant (O2) is being converted. However, the total amount of generated heat is higher for T0 ) 360 °C, because of the poorer selectivity to ethylene (SG) at the reactor outlet (combustion reactions 2 and 3 are much more exothermic than the ODH reaction). This selectivity lost is the cause of the observed increase in the maximum temperature when the feed stream is cooled from 380 to 360 °C. Influence of the Oxygen Feed Content. As shown in Figure 1, a specified amount of air can be fed to the reactor mouth to increase the reaction rate in the first part of the membrane reactor. This action leads to an increase in ethylene production, but also causes selectivity to deteriorate. As seen in Figure 4, the presence of O2 in the feed stream has an important influence on the temperature profiles. For the selected inlet temperatures, the reaction does not ignite when no O2 is fed at z ) 0. Under these operating conditions, the heat-transfer rate toward the shell side is always higher than the heat generation rate, and therefore, the temperature decreases monotonically along the axial direction. Because the reaction rates are very low, the permeation flux through the membrane contributes to increase the oxygen content inside the reaction mixture from zero to a maximum value at the reactor outlet. Consequently, the ethylene selectivity decreases monotonically (Figure 5). When 1% oxygen is fed, the temperature shows a flat profile in the first half of the reactor up to the region where the reaction ignites. Beyond the axial position 2.4 m, the temperature shows a slight increase. For the case of 5% O2 in the feed, the temperature rises sharply, and the oxygen content inside the membrane tubes decreases rapidly from its inlet value toward a very low value for z >
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Figure 4. Tube-side (solid lines) and shell-side (dashed lines) temperatures vs axial position for different inlet oxygen contents. T0 ) 360 °C, TS0 ) 25 °C.
Figure 5. Axial profiles of oxygen molar flow (FO2) and global selectivity (SG) for the conditions of Figure 4.
0.7 m (Figure 5). It is important to note that the selectivity profiles corresponding to 1% and 5% O2 show minima that are located at axial positions where the oxygen content is decreasing sharply. That is, when the oxygen partial pressure of O2 diminishes, the instantaneous selectivity (SP) shows a rapid increase, and the global selectivity (SG) tends to improve gradually. The corresponding conversion profiles of ethane are presented in Figure 6. When no O2 is added to the feed stream, the outlet conversion is very poor (around 3%). The addition of 1% O2 at the reactor mouth renders possible the conversion of almost 15% of the ethane feed, associated with an ethylene selectivity of 92% (see Figure 5), which is a very good value. The reactor selectivity is a key parameter for an ODH plant, because, at conditions of total recycle of the unconverted ethane (the usual operation mode), the ethylene yield of the entire plant (reactor and separation section) coincides with the value of the reactor selectivity. That is, a selectivity of 92% means that 1.087 kmol of C2H6 would be necessary to produce 1 kmol of C2H4 (1.164 tons of C2H6 per ton of C2H4). At first glance, the case with 5% O2 seems to be the best of the three alternatives, because the outlet selectivity is the highest (Figure 5) and the ethane conversion is also the highest (Figure 6). However, a sharp temperature rise in the first 70 cm of the tube is observed in Figure 4, with a hot spot of around 550 °C followed by a slight overcooling. Such operating conditions would not be feasible,
Figure 6. Axial profiles of ethane conversion for different inlet oxygen contents. Same conditions as in Figure 4.
Figure 7. Tube-side (solid lines) and shell-side (dashed lines) temperatures vs axial position for different inlet shell temperatures. Inlet O2 ) 0%, T0 ) 360 °C.
because of the high parametric sensitivity. Moreover, the Ni-Nb-O catalyst was not tested in that temperature range.3,4 Influence of the Inlet Temperature on the Shell Side (TS0). The inlet temperature of the air stream (TS0) is an important parameter, because it determines the heat-transfer rate between the process stream and the cooling medium. Figure 7 shows the influence of TS0 on the reactor temperatures for a feed temperature T0 ) 360 °C and no oxygen at z ) 0. When the air stream is fed to the shell side at room temperature, the convective heat-transfer rate in the first part of the reactor is too high. In addition, the cold-shot effect of the air being permeated through the membrane is at a maximum, because the shell-side temperature is at a minimum (around 25 °C). Then, the reactor is overcooled (Figure 7), and an accumulation of O2 along the reactor position is detected (Figure 8). At this point, two alternatives appear possible to ignite the reactor. The first is to decrease the air flow rate through the shell side. This option would be convenient to decrease the power requirements of the air compressor. (Certainly, the total air flow rate given in Table 2 is quite high.) However, a reduction of the air flow rate could substantially decrease the heat-transfer coefficients on the shell side (the controlling side for the heat transfer), leading to lower overall heat-transfer coefficients and higher parametric sensitivity. The second alternative is to maintain the air flow rate and increase the value of TS0. In fact, the reaction can be ignited
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Figure 8. Axial profiles of oxygen molar flow (FO2) and global selectivity (SG) for the conditions of Figure 7.
Figure 10. Ethylene production for different inlet operating conditions.
temperatures and the permeation flows must be selected carefully to avoid undesired temperature rises and/or reaction extinction. Conclusions
Figure 9. Axial profiles of ethane conversion for different inlet oxygen contents. Same conditions as in Figure 7.
for TS0 ) 100 and 200 °C, in which case the undesired oxygen accumulation inside the tubes tends to disappear and the ethylene selectivity tends to improve (Figure 8). As shown in Figure 9, the outlet ethane conversion for TS0 ) 100 and 200 °C is around 13%. Figure 9 shows another important characteristic: The conversion profiles are almost linear for axial positions beyond the point where the oxygen depletion takes place (z ≈ 1.5 m for TS0 ) 200 °C and z ≈ 2.5 m for TS0 ) 100 °C; see Figure 8). For all of these axial positions, the ethane conversion rate (i.e., the slope of the straight lines in Figure 9) is determined by the permeation rate of oxygen through the membrane pores. At most of the studied operating conditions, this linear behavior of the ethane conversion is associated with slightly increasing and almost linear axial temperature profiles (Figures 2, 4, and 7). Figure 9 also suggests that the permeations flows are not strongly dependent on the temperature in the shell side, because the straight portions of the conversion profiles are almost parallel. This result is consistent with the adopted model for the permeation flows, which are not activated by temperature. Last, the influence of the main operating variables on the ethylene production is shown in Figure 10. As can be seen, production rates of around 250 000 tons/year would be feasible with the proposed multitubular membrane reactor. As shown before, the presence of a certain amount of O2 in the reactor mouth tends to increase the reactor production. Both the inlet
According to the previous results, it is possible to reach the following conclusions: (1) The proposed multitubular membrane reactor can be considered as an attractive alternative for carrying out the oxidative dehydrogenation of ethane on a highly active and selective Ni-Nb-O ODH catalyst. The low oxygen partial pressures inside the catalyst tubes lead to good ethylene selectivities, good temperature control as a result of the low heat generation rates, and reasonable production rates. (2) Depending on the selected inlet temperatures, there are convenient values for the amount of oxygen to be fed to the reactor mouth. (3) Under certain operating conditions, the membrane reactor for ODH presents inverse parametric sensitivity with respect to the inlet temperature, as a consequence of the accumulation of O2 inside the tubes and its influence on the ethylene selectivity and heat generation rates. (4) When the reactor is operated at conditions where the reaction is controlled by the permeation flow of O2 through the membrane, it is possible to reach high selectivities to ethylene, significant ethane conversions, and mild temperature profiles. (5) Some of the main characteristics of the proposed design, e.g., the convenience of using a liquid coolant medium to remove the generated heat, the use of cross flow in the shell side, and so on, could be modified to improve the reactor performance and optimize the ethylene yield. The use of pure oxygen instead of air as a cooling medium should also be considered. In this last case, the need for air separation to obtain pure oxygen should be balanced with the difficulties associated with the separation of N2 from C2H4 downstream from the reactor. (6) The possible failure of some of the membrane tubes is, of course, a safety problem that should be studied carefully, in order to avoid the possible formation of explosive mixtures within the operating range of the reactor variables. Acknowledgment Support of this work through Universidad Nacional del Sur (UNS) and Consejo Nacional de Investigaciones Cientı´ficas y Tecnolo´gicas (CONICET) is gratefully acknowledged.
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Nomenclature AT ) cross-sectional area of the tubes, m B0 ) geometric parameter, m2 Cpj ) specific heat of component j, kJ/(kmol K) dp ) equivalent diameter of the catalyst pellet (momentum equation), m dT ) internal tube diameter, m dS ) shell diameter, m Dej,k ) Effective Knudsen diffusion coefficient of component j, m2/s Fj ) molar flow of component j (reaction side), kmol/h FjS ) molar flow of component j (shell side), kmol/h ∆Hr,i ) heat of reaction of reaction i, kJ/mol Jj ) permeation flow of component j, kmol/(h m2) K0 ) geometric parameter, m Mj ) molecular weight of component j, kg/kmol nT ) number of tubes L ) tube length, m P ) total pressure (reaction side), Pa PS ) total pressure (shell side), Pa Pm ) average pressure, Pa ∆P ) pressure drop across the membrane, Pa ri ) reaction rate of reaction i, kmolHC/(kgcat s), where HC ) C2H6 (reactions 1 and 2) and HC ) C2H4 (reaction 3) R ) universal gas constant, kg m2/(kmol s2 K) T ) temperature (reaction side), K TS ) temperature (shell side), K Tm ) average temperature, K U ) overall heat-transfer coefficient, kJ/(h m2 K) us ) superficial velocity, mf3/(mr2 s) SG ) global selectivity SI ) instantaneous selectivity z ) axial coordinate, m Greek Letters νij) stoichiometric coefficient (with i reaction and j component), dimensionless δ ) membrane thickness, m Fg ) gas density, kg/mf3 FB ) bed density, kgcat/m3 µj ) viscosity of component j, Pa s Subscripts j ) component j i ) reaction i 0 ) at the axial coordinate z ) 0 2
L ) at the axial coordinate z ) L C2H6 ) ethane C2H4 ) ethylene CO2 ) carbon dioxide H2O ) water O2 ) oxygen N2 ) nitrogen Superscript S ) shell side
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ReceiVed for reView April 08, 2008 ReVised manuscript receiVed August 19, 2008 Accepted August 27, 2008 IE800564V
