Development of a Membrane-Assisted Fluidized Bed Reactor. 2

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Development of a Membrane-Assisted Fluidized Bed Reactor. 2. Experimental Demonstration and Modeling for the Partial Oxidation of Methanol S. A. R. K. Deshmukh, J. A. Laverman, M. van Sint Annaland,* and J. A. M. Kuipers Department of Science and Technology, University of Twente, P. O. Box 217, 7500 AE Enschede, The Netherlands

A small laboratory-scale membrane-assisted fluidized bed reactor (MAFBR) was constructed in order to experimentally demonstrate the reactor concept for the partial oxidation of methanol to formaldehyde. Methanol conversion and product selectivities were measured at various overall fluidization velocities, reactor temperatures, methanol and oxygen overall feed concentrations, ratios of gas fed via membranes relative to gas fed via the bottom distributor, and aspect ratios of the fluidized bed. High methanol conversions and high selectivities to formaldehyde were achieved with safe reactor operation (isothermal reactor conditions) at very high methanol inlet concentrations, much higher than currently employed in industrial processes. It was experimentally demonstrated that with distributive feeding of oxygen in a MAFBR the overall formaldehyde yield and throughput could be increased without a pronounced and undesirable conversion of formaldehyde to carbon monoxide. Furthermore, a one-dimensional two-phase phenomenological reactor model has been developed with which the experimentally observed conversion and selectivity as a function of the operating conditions could be well described. 1. Introduction Highly exothermic heterogeneously catalyzed gas phase partial oxidation reactions are an industrially important class of chemical transformations and require carefully designed reactors because of the large amount of reaction heat liberated and the high selectivity requirement for the intermediate product of interest. A new reactor concept is proposed which offers improved reactor safety and higher product throughput via the concept of multifunctionality. Our concept aims to combine the advantages of excellent heat transfer and gas-solid mixing characteristics of fluidized beds with the controlled dosing capability of membrane reactors to achieve an inherently safe and isothermal reactor operation. Moreover, this membrane-assisted fluidized bed reactor (MAFBR) is especially interesting for highly exothermic oxidation reactions, where the explosion region of the reactants or products imposes restrictions on the feed concentration of the reactants. Reactor operation can be rendered inherently safe via distributive addition of oxygen, enabling lower local oxygen concentrations in the fluidized bed and thus operation outside the flammability limits of the hydrocarbons. The application of fluidized bed membrane reactors for reactions of industrial importance has been investigated in the recent past. Adris et al.1,2 demonstrated that for steam reforming of natural gas the thermodynamic equilibrium restrictions can be overcome by in situ separation and removal of hydrogen via permselective thin-walled palladium-based membranes leading to increased synthesis gas yields in comparison to the industrial fixed bed steam reformer. Using simula* To whom correspondence should be addressed. Tel.: 0031-53-4894478. Fax: 0031-53-4892882. E-mail: [email protected].

tions Abdalla and Elnashaie3 showed for the catalytic dehydrogenation of ethyl benzene to styrene and Ostrowski et al.4 showed for the catalytic partial oxidation of methane to synthesis gas that with fluidized bed membrane reactors higher product selectivities could be realized compared to fixed bed reactors. Furthermore, gas withdrawal through the membranes decreases the superficial gas velocities in the top section of the bed, resulting in smaller gas bubbles, which increases the interphase gas exchange favoring high conversions (Mleczko et al.5). In these studies the insertion of permselective hydrogen membranes in a fluidized bed was investigated. In this paper we aim to demonstrate the advantages of the membrane-assisted fluidized bed for controlled dosing of one of the reactants via the membranes. In part 1 the effects of the presence of the membranes and the permeation of gas through the membranes on the gas back-mixing and the bubble-to-emulsion phase mass transfer rates have been investigated. In part 2 this study is extended for a chemically reactive system. In a lab-scale membrane-assisted fluidized bed demonstration unit, the integration of porous nonselective ceramic membranes for controlled dosing of air in a fluidized bed for the partial oxidation of methanol to formaldehyde is studied experimentally. The reaction scheme for the partial oxidation of methanol and its side reactions have been listed in Table 1. The two-phase model accounting for permeation through the membranes, developed in part 1, is extended to include the effects of gas production due to chemical reactions. With this model the experimental results are interpreted to facilitate understanding of the experimental observations. Before presenting and discussing the experimental results, the experimental setup and experimental procedure are shortly described in the next section.

10.1021/ie049092h CCC: $30.25 © 2005 American Chemical Society Published on Web 02/23/2005

Ind. Eng. Chem. Res., Vol. 44, No. 16, 2005 5967 Table 1. Reaction Scheme of Main Reaction and Side Reactions during Partial Oxidation of Methanol, Taken from Deshmukha

a

Reference 6.

2. Experimental Section The partial oxidation of methanol was performed over a commercial Fe-Mo catalyst in a square (0.05 m × 0.05 m × 1 m) membrane-assisted fluidized bed reactor (for details see part 1). The original catalyst was made available by Perstorp in the form of hollow cylindrical pellets having a height of 2.4-2.8 mm, an internal diameter of 2.4-2.6 mm and an outer diameter of 4.8-5.2 mm, and a surface area of 6 m2/g. The catalyst was crushed to the size of 75-150 µm particles. Typically about 400 g (for 0.15 m packed bed height) of this crushed catalyst was packed into the reactor. The detailed experimental setup is shown in Figure 1. The reactor temperature was controlled by heaters fixed on the wall of the reactor. Temperatures inside the reactor were measured at four different axial locations in the catalyst bed and at one location in the freeboard region above the catalyst bed. The reactor was fed with an air, methanol, and nitrogen mixture. Methanol was fed via the bottom distributor by saturating nitrogen in a liquid methanol bubbler. Additional oxygen and nitrogen were also fed via the bottom distributor to achieve the desired flow and an oxygen inlet concentration of 9-10%, just under the flammability limits. In the case of experiments with membrane permeation, additional oxygen and nitrogen were fed via the membranes. Isothermal conditions in the bubbler were maintained by circulating hot water on the shell side. The pressure and temperature of the bubbler were measured and controlled digitally in order to provide constant methanol concentrations in the feed, which were measured with a Varian 3400 gas chromatograph (GC) equipped with a thermal conductivity detector. A Molsieve-5A column was used for the separation of oxygen, nitrogen, and carbon monoxide, and a Hayesep-T column was used to separate formaldehyde, methanol, water, dimethyl ether (DME), dimethoxymethane (DMM), and methyl formate (MF), where helium was used as the carrier gas and a temperature-programmed analysis was applied at 60-160 °C. The detector signals (chromatograms) were stored in a PC using the data acquisition package Chromatographic Separation for Windows (CSW 32). Special precautions were taken to avoid condensation of methanol and polymerization of formaldehyde to para-formaldehyde by heating the freeboard and outlet of the reactor and the lines to the GC, as well as the outlet of the gas chromatograph at 135 °C. The partial oxidation of methanol to formaldehyde was studied in the temperature range of 250-300 °C. All the experiments were performed at near atmospheric pressure.

Before the experiments were initiated, the reactor was preheated using a heater attached to the membranefree walls of the reactor. Air was used to fluidize the reactor, avoiding temperature gradients in the reactor. The reactor was heated to a temperature well below the desired reactor temperature, and the methanol oxidation reaction heat was utilized to reach the final reactor temperature. Subsequently methanol was fed by bubbling nitrogen through the saturator and the air flow rate was slowly increased to avoid the formation of explosive mixtures and a sudden rise in reactor temperature. For each set of experiments, the feed concentrations were measured by bypassing the reactor and sending the feed directly to the GC. In the next section, the experimental results obtained for the partial oxidation of methanol in the membrane assisted fluidized bed reactor will be presented. These results will subsequently be interpreted using a twophase phenomenological reactor model. 3. Experimental Results To demonstrate the benefits of the novel membraneassisted fluidized bed reactor concept for the partial oxidation of methanol over an Fe-Mo catalyst, experiments were conducted over a wide range of methanol and oxygen concentrations, fluidization velocities, permeation ratios (i.e., amount of gas fed via the membranes relative to the total amount of gas fed), catalyst bed heights and bed temperatures. For all the experiments, the maximum observed temperature difference throughout the active part of the reactor was only 1-2 °C, indicating nearly perfect isothermal conditions. 3.1. Effect of the Temperature. Figure 2 shows the effect of the reactor temperature on the methanol conversion and formaldehyde selectivity and yield in a MAFBR at a fluidization velocity of 18umf at reactor temperature conditions (10umf at room temperature), a packed bed height of 0.3 m, and 20% methanol and 12% oxygen overall inlet concentrations based on the total feed, where 50% of the total flow was fed via the membranes. It is important to mention that with these overall inlet concentrations safe and isothermal reactor operation was achieved, despite the fact that this overall feed composition is well inside the flammability region. Figure 2 shows that the methanol conversion increases drastically when the temperature increases from 250 to 275 °C. With a further temperature increase, the methanol conversion hardly increases because of mass transfer limitations preventing unreacted methanol from being transferred from the bubble phase to the emulsion phase. Second, the formaldehyde selectivity decreases with higher temperatures because the formaldehyde in the emulsion phase is more susceptible to further oxidation at higher temperatures, because of the more pronounced mass transfer limitations and higher reaction rates to CO at higher temperatures. 3.2. Effect of the Methanol Inlet Concentration. In Figure 3 the effect of the methanol inlet concentration on the methanol conversion and formaldehyde selectivity and yield is shown at 275 °C and 18umf overall fluidization velocity, where 50% of the total gas was fed via the membranes. The methanol conversion decreases with an increase in the methanol inlet concentration, keeping the overall inlet oxygen concentration the same at 10%. Assuming that the methanol mass transfer rate increases linearly with the methanol concentration, the decrease in the methanol conversion indicates that the

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Figure 1. Schematic of the experimental setup of the MAFBR for the partial oxidation of methanol to formaldehyde.

Figure 2. Effect of reactor temperature on the methanol conversion and selectivities to the reaction products at an overall fluidization velocity of 18umf, initial packed bed height of 0.3 m, and 20% methanol and 12% oxygen overall inlet concentrations (50% permeation through the membranes).

overall conversion rate is at least partly influenced by reaction kinetics, where lower oxygen concentrations result in lower reaction rates. The selectivity to formaldehyde increases with an increase in the methanol inlet concentration, because at higher methanol inlet concentrations the stoichiometric feed composition is approached, which decreases the excess of oxygen and thereby favors the formation of formaldehyde rather than carbon monoxide. It is also important to notice that the selectivity toward dimethyl ether increases at higher methanol concentrations, where the catalyst surface becomes more saturated with methoxy-groups, resulting in higher dimethyl ether formation rates. These experiments also demonstrate that the formaldehyde productivity of the MAFBR can be increased almost linearly by increasing the methanol inlet concentration, since the formaldehyde yield remains almost constant.

Figure 3. Effect of the methanol feed concentration on the methanol conversion and selectivity to the reaction products at an overall fluidization velocity of 18umf, where 50% of the total gas was fed via the membranes, using an overall oxygen concentration in the feed of 10%, an initial packed bed height of 0.3 m, and a reactor temperature of 275 °C.

3.3. Effect of the Oxygen Concentration. From Figure 4 it can be concluded that feeding oxygen in excess of the stoichiometric value, the methanol conversion is strongly increased because of the higher local oxygen concentrations in the emulsion phase, which increase the reaction rates. Also the CO selectivity increases and the formaldehyde selectivity decreases because of the higher local oxygen concentration. Concluding, a stoichiometric methanol-oxygen (2:1) feed composition and a reactor temperature of 275 °C produce the highest formaldehyde yield. These conditions have been selected in the subsequent study on the effects of the fluidization velocity, the flow rate of the gas added via the membranes, and the initial catalyst packed bed height.

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Figure 4. Effect of the oxygen overall feed concentration on the methanol conversion and selectivity to the reaction products at an overall superficial velocity of 18umf, a catalyst packed bed height of 0.3 m, a reactor temperature of 275 °C, 15% overall methanol concentration in the feed, and 50% of the total gas added via the membranes.

3.4. Effect of the Fluidization Velocity. Figure 5 shows the effect of the dimensionless fluidization velocity (u/umf) (corresponding to room temperature) on the methanol conversion (C), formaldehyde selectivity (S), formaldehyde yield (Y), and carbon monoxide selectivity for overall inlet methanol and oxygen concentrations of 17% and 9% respectively (i.e., nearly stoichiometric feed composition) for different membrane permeation ratios varying between 0 and 40%. The following important observations can be discerned from the experimental data: 1. The methanol conversion decreases with an increase in the fluidization velocity for all permeation ratios and bed heights investigated. This can be attributed to a more pronounced mass transfer limitation of methanol from the bubbles, formed at the distributor, to the emulsion phase. The average bubble diameter

Figure 5. The effect of the dimensionless overall fluidization velocity (u/umf) (at room temperature) on the methanol conversion (C), formaldehyde selectivity (S), formaldehyde yield (Y), and carbon monoxide selectivity for overall inlet methanol and oxygen concentrations of 17% and 9%, respectively, at 275 °C and for different membrane permeation ratios: (a) packed bed height of 0.15 m (L/D ) 3), (b) packed bed height of 0.20 m (L/D ) 4), and (c) packed bed height of 0.25 m (L/D ) 5).

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Figure 6. Effect of catalyst packed bed height (L/D ) aspect ratio) on the methanol conversion (C), formaldehyde selectivity (S), formaldehyde yield (Y), and carbon monoxide selectivity for overall methanol and oxygen concentrations of 17% and 9%, respectively, at 275 °C for different membrane permeations ratios: (a) u/umf ) 8 and (b) u/umf ) 10, all corresponding to room temperature. Table 2. Effect of Imposing Different Axial Oxygen Concentration Profiles via Distributing Different Flow Rates through the Three Bundles on the Methanol Conversion and Product Selectivities, Where 20% of the Total Gas Was Added via the Membrane, with a Reactor Temperature of 275 °C, Overall Inlet Methanol and Oxygen Concentrations of 17% and 9%, Respectively, a Catalyst Packed Bed Height of 0.25 m, and an Overall Fluidization Velocity of 8umf Corresponding to Room Temperaturea feed to bundle 1 [%]

feed to bundle 2 [%]

feed to bundle 3 [%]

CH3OH conversion [%]

HCHO selectivity [%]

HCHO yield [%]

CO selectivity [%]

DME selectivity [%]

33 50 66 33 50 40 50

33 50 33 66 0 40 30

33 0 0 0 50 20 20

95.40 94.05 93.32 94.22 92.69 94.01 92.98

89.63 90.30 91.02 90.65 91.21 91.42 91.38

85.51 84.93 84.94 85.41 84.54 85.41 84.96

8.43 7.23 6.60 7.11 6.14 6.86 6.36

1.43 1.62 1.63 1.53 1.75 1.73 1.74

a The centers of bundles 1, 2, and 3 were at 6.4, 12.8, and 19.2 cm above the distributor, respectively, and each bundle consisted of 6 membranes.

(and thus the bubble-to-emulsion phase mass transfer resistance) increases with an increase in the fluidization velocity and bed height. 2. The selectivity toward formaldehyde increases at higher fluidization velocities, mainly because of the lower formaldehyde residence time. Bubbles help to increase the selectivity by trapping the formaldehyde formed lower in the bed, which is not readily transferred back to the emulsion phase in the top of the bed due to the increased mass transfer resistance because of bubble growth. 3. The formaldehyde yield decreases with an increase in the fluidization velocity, because of the stronger decrease in the methanol conversion compared to the increase in formaldehyde selectivity. 4. Addition of part of the total gas via the membranes (distributive addition of oxygen) increases the methanol conversion, first because of the higher methanol resi-

dence time caused by the reduction in the gas flow rate fed via the distributor, and second because of the smaller bubble size which improves the methanol transfer rate to the emulsion phase. However, the selectivity to formaldehyde decreases because of the increased formaldehyde residence time in the bed and also because of the addition of oxygen at the top of the bed via the ceramic membranes where the formaldehyde concentrations are the highest. 5. It is striking to notice that the carbon monoxide selectivity remains at about 5% at higher overall fluidization velocities of 10umf even with addition of gas via membranes. Hence low selectivities to carbon monoxide can be maintained while achieving higher conversions by using distributive feeding of oxygen via the membranes. The byproducts in our experiments were dimethyl ether and methyl formate and occasionally dimethoxy methane, especially at high inlet methanol

Ind. Eng. Chem. Res., Vol. 44, No. 16, 2005 5971 Table 3. Mass Conservation Equations for the Bubble and Emulsion Phase of Compartment n Total Mass Balance s s s s ub,n-1 AT - ub,n AT + ue,n-1 AT - ue,n AT + utotATF

+

Nb

∑(∑ν

RT

i,j)Rjmcat

j

i

P

)0

where s ue,n AT ) ue,nAT(1 - δb,n), usb,0AT ) (1 - F)utotATδb,0, use,0AT ) (1 - F)utotAT(1 - δb,0) Bubble Phase Component Mass Balances utotATFCp,n,iδb,n + Nb

s s ub,n-1 ATCb,n-1,i - ub,n ATCb,n,i +

[

s s (ue,n-1 AT - ue,n AT) +

∑(∑ν j

]

RT

i,j)Rjmcat

i

P

for i ) 1...nc

utotATF(1 - δb,n) + Nb

Ce,n,i - Kb,e,nVb,n(Cb,n,i - Ce,n,i) ) 0

Emulsion Phase Component Mass Balances s s ue,n-1 ATCe,n-1,i - ue,n ATCe,n,i +

[

utotATFCp,n,i(1 - δb,n) Nb

s s (ue,n-1 AT - ue,n AT) +

∑(∑ν j

i

]

RT

i,j)Rjmcat

P

utotATF(1 - δb,n) + Nb

Ce,n,i + Kb,e,nVb,n(Cb,n,i - Ce,n,i) +

∑ν

i,jRjmcat

)0

j

for i ) 1...nc

concentrations above 15%, which can be recovered and recycled to be converted back to formaldehyde and CO eventually. 6. Finally, the same trends for the conversion and selectivities as a function of the fluidization velocity and membrane permeation ratio were observed for all of the three bed heights studied. 3.5. Effect of the Catalyst Bed Height. In Figure 6 the effect of the initial packed bed height on the methanol conversion (C), formaldehyde selectivity (S),

formaldehyde yield (Y), and carbon monoxide selectivity is shown for an overall methanol and oxygen concentration of 17% and 9%, respectively, for different membrane permeation ratios (0-40%) and fluidization velocities (u/umf). The methanol conversion increases with higher catalyst bed heights. The selectivity toward formaldehyde decreases slightly, but the selectivity toward the undesired carbon monoxide increases significantly, mainly because of the increased residence time of methanol and formaldehyde, respectively. Moreover, increasing the part of the total fluidizing gas fed via the membranes increases the conversion at low aspect ratios. However, the effect of the permeation ratio on the conversion becomes negligible at a bed height of 0.25 m (L/D ) 5) at both overall velocities studied. Apparently distributing the fluidizing gas via the membranes to increase the methanol conversion is most effective at lower aspect ratios because of the relatively large residence times required to achieve near complete conversion of methanol. 3.6. Effect of Influencing the Axial Oxygen Concentration Profile via Axially Varying Membrane Permeation Flows. Table 2 shows the effect of imposing different oxygen concentration profiles created by distributing different air flow rates via the membrane bundles. These experiments were carried out at an overall superficial velocity of 8umf, overall inlet methanol and oxygen concentrations of 17% and 9%, respectively, a packed bed height of 0.25 m, and a reactor temperature of 275 °C, with 20% of the total gas added via the membranes at various distribution ratios within the three membrane bundles. The experimental results clearly show that there is little effect of imposing different axial oxygen concentration profiles in the reactor by feeding different air flow rates via each of the bundles. The overall formaldehyde yield was always about 85%. Little effect on the product selectivities was also to be expected since the reaction order in oxygen is identical for both reactions, the methanol partial oxidation to formaldehyde and the formaldehyde partial oxidation to carbon monoxide.6 However, the results show that by feeding relatively more oxygen at the bottom of the bed and simultaneously feeding a corresponding lower flow rate at the

Table 4. Kinetic Rate Equations for the Methanol Partial Oxidation Reaction System, Taken from Deshmukha parameter

a

equation Kads,CH3OHPg,CH3OH

surface fraction of methanol

θCH3OH )

surface fraction of oxygen

θO2 )

surface fraction of formaldehyde

θHCHO )

surface fraction of DME

θDME )

reaction rate toward formaldehyde reaction rate toward CO

RHCHO ) kHCHOθCH3OHθO2 RCO ) kCOθHCHOθO2

reaction rate toward DME

RDME ) kDME,fPg,CH3OH -

reaction rate toward DMM

RDMM ) kDMM,fPg,HCHOPg,CH3OH -

reaction rate toward formaldehyde from DME

RDME - HCHO ) kDME•HCHOθDMEθO2

Reference 6.

1 + Kads,CH3OHPg,CH3OH + Kads,H2OPg,H2O

Kads,O2Pg,O21/2 1 + Kads,O2Pg,O21/2 Pg,HCHO 1 + Kads,CH3OHPg,CH3OH + Kads,H2OPg,H2O

Kads,DMEPg,DME 1 + Kads,DMEPg,DME

kDME,f Pg,DMEPg,H2O KDME,eq Pg,CH3OH kDMM,f Pg,DMMPg,H2O KDMM,eq Pg,CH3OH

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Table 5. Reaction Rate Constants (Pre-exponential Constants and Activation Energies) and Equilibrium Constants for the Methanol (RHCHO) and Formaldehyde Partial Oxidation (RCO), Dimethyl Ether Formation (RDME), and Dimethoxy Methane Formation (RDMM), Taken from Deshmukha (RHCHO)

(RCO) (RDME) (RDMM) (RDME-HCHO) a

KCH3OH [atm-1] KO2 [atm-1/2] KH2O [atm-1] kHCHO [mol kg-1 s-1] kCO [mol kg-1 atm-1 s-1] kDME,f [mol kg-1 atm-1 s-1] Keq DME ) exp(-2.2158 + 2606.8/T) kDMM,f[mol kg-1 atm-2 s-1] Keq DMM ) exp(-20.416 + 9346.8/T) KDME [atm-1] kDME-HCHO [mol kg-1 s-1]

2.6 × 10-4 1.423 × 10-5 5.5 × 10-7 1.5 × 107 3.5 × 102 1.9 × 105

ECH3OH EO2 EH2O EHCHO ECO EDME,f

[kJ‚mol-1] [kJ‚mol-1] [kJ‚mol-1] [kJ‚mol-1] [kJ‚mol-1] [kJ‚mol-1]

-56.78 -60.32 -86.45 86.0 46.0 77.0

4.26 × 10-6

EDMM,f

[kJ‚mol-1]

46.5

5 × 10-7 6.13 × 105

EDME EDME-HCHO

[kJ‚mol-1] [kJ‚mol-1]

-96.72 98.73

Reference 6.

top, the methanol conversion is slightly decreased, the formaldehyde selectivity slightly increased, and the CO selectivity slightly decreased. Thus the CO losses can be somewhat reduced at the same formaldehyde yield. The lower methanol conversion indicates a slightly more pronounced bubble-to-emulsion phase mass transfer limitation due to the less optimal flow distribution over all the membranes resulting in a slightly larger average bubble size. On the other hand, the mass transfer limitations for the formaldehyde help to keep the CO selectivity low. 4. Modeling 4.1. Model Equations. A one-dimensional two-phase model for the MAFBR was developed in part 1, which accounts for gas permeation through the membranes. In this model the fluidized bed is divided in the axial direction into a series of well mixed compartments (CISTRs) for the bubble phase as well as a series of CISTRs for the emulsion phase, where the number of compartments for each phase, each of equal volume, is related to the extent of axial gas back-mixing. From tracer injection experiments, it was concluded in part 1 that both phases in the experimental MAFBR exhibit approximately plug flow behavior, so that the number of CISTRs in the emulsion phase is specified at 5 (large enough to exclude any influence of the number of CISTRs, which was concluded from a parametric sensitivity analysis) and for the bubble phase at 20. A detailed description of the model assumptions and a complete list of the constitutive equations used can be found in part 1. This model has been extended to include the partial oxidation of methanol over an Fe-Mo catalyst and its main consecutive and parallel reactions. On the basis of the model assumptions detailed in part 1, the steadystate bubble and emulsion phase total and component mass conservation equations have been formulated for each compartment, taking chemical transformations in the emulsion phase and a net gas production due to the chemical reactions into account (see Table 3). For an explanation of the symbols used, the reader is referred to the symbol list. Detailed reaction scheme and kinetic rate expressions, based on measurements carried out over wide concentration ranges, have been taken from Deshmukh6 and have been summarized in Tables 1, 4, and 5. A schematic representation of the gas flows between the compartments of the bubble and emulsion phases that constitute the fluidized bed is depicted in Figure 7. 4.2. Parametric Sensitivity. Before actually comparing the predictive capabilities of the developed

model, simulations were performed to assess the sensitivity of the two main model parameters, namely, the number of compartments in series for the emulsion phase (i.e., the extent of gas back-mixing in the emulsion phase) and the bubble diameter factor fdb (i.e., the factor with which the bubble diameter as calculated with the correlation proposed by Mori and Wen7 needs to be modified due to the presence of the membranes and gas permeation through the membranes). Both these parameters have been determined by independent tracer injection experiments and residence time distribution (RTD) measurements using ultrasound in part 1 as a function of the operating conditions. The sensitivity of these parameters on the methanol conversion and product selectivities is discussed below. 4.2.1. Effect of the Number of Emulsion Phase Compartments. Figure 8 shows the sensitivity of the number of compartments in the emulsion phase on the methanol conversion and product selectivities for a base

Figure 7. Schematic representation of the two-phase reactor model; B ) bubble phase, E ) emulsion phase.


Figure 8. Sensitivity of the predicted methanol conversion and product selectivities with respect to the number of CISTRs in the emulsion phase. Conditions: 17% methanol and 9% oxygen overall inlet concentration, at 275 °C reactor temperature, an overall fluidization velocity of 10umf (corresponding to room temperature), and a packed bed height of (a) 0.2 m and (b) 0.15 m. As determined from RTD experiments: fdb ) 1.

Figure 9. Sensitivity of the predicted methanol conversion and product selectivities with respect to the bubble diameter factor fdb. Conditions: 17% methanol and 9% oxygen overall inlet concentration, at 275 °C reactor temperature, a packed bed height of 0.2 m, and an overall fluidization velocity of 10umf (corresponding to room temperature). The number of CISTRs in the emulsion phase was fixed at 5.

case experiment. From the RTD studies reported in part 1, it was concluded that for these operating conditions fdb equals unity. From this figure it is very clear that the number of emulsion phase CISTRs hardly affects the methanol conversion if this number exceeds 3, while the formaldehyde and CO selectivities are basically independent of the number of emulsion phase CISTRs. These simulations demonstrate that the number of emulsion phase CISTRs is not a very critical parameter for this reaction system. (Note that the oxygen dependency for the partial oxidation of methanol to formaldehyde and the oxidation of formaldhyde to CO is identical and almost zero order under the reaction conditions investigated.) 4.2.2. Effect of the Bubble Diameter Factor. In Figure 9 the effect of the bubble diameter factor on the methanol conversion and product selectivities is shown keeping the number of emulsion phase CISTRs at five. From this figure it is very clear that the methanol conversion strongly depends on the bubble diameter factor, as expected because of the prevailing bubble-toemulsion phase mass transfer limitations, which become more pronounced for bigger bubbles. However, the selectivity to formaldehyde and CO is not significantly influenced by the bubble diameter. Smaller bubbles improve the transfer of the formaldehyde formed in the emulsion phase to the bubble phase, which thereby increases the selectivity to formaldehyde by trapping the formaldehyde inside the bubbles and hence prevent-

Figure 10. Comparison of methanol conversion in the MAFBR as a function of the gas residence time by varying the catalyst bed height (0.15-0.25 m) and the fluidization velocity (10-20umf at room temperature) for different bubble diameter factors. Conditions: 17% methanol and 9% oxygen overall inlet concentration, at 275 °C reactor temperature and permeation ratios of (a) 0% and (b) 40%.

ing CO formation (see Figure 9 at lower fdb values); however, this effect is very small. Moreover, the figure shows that the measured methanol conversion, as well as the formaldehyde and CO selectivities, can be very well predicted with the developed two-phase model using fdb ) 1 (i.e., unmodified bubble diameter as calculated with the Mori and Wen7 correlation), as was also determined from independent RTD measurements using ultrasound for these operating conditions. 4.3. Comparison with Experiments. In this section the capabilities of the developed two-phase reactor model to describe the methanol conversion and product selectivities in the MAFBR are assessed by comparison

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Figure 11. Comparison of the predicted methanol conversion and product selectivities by the two-phase model with the experimental results as a function of the catalyst packed bed height for a 17% methanol and 9% oxygen overall feed composition at 275 °C reactor temperature, an overall fluidization velocity of 10umf corresponding to room temperature, and three different permeation ratios: (a) 0% permeation, fdb ) 0.9; (b) 20% permeation, fdb ) 0.9; and (c) 40% permeation, fdb ) 0.8.

with the experimental results obtained with the experimental demonstration unit. 4.3.1. Methanol Conversion as a Function of Gas Residence Time. First, the methanol conversions measured as a function of the (average) gas residence time (by varying both the fluidization velocity and the catalyst bed height) have been compared for two different permeation ratios (0 and 40%) with the computed methanol conversions using the developed two-phase reactor model, where different bubble diameter factors have been assumed (see Figure 10). Figure 10 clearly shows a strong influence of bubble-to-emulsion phase mass transfer limitations on the methanol conversion. The measured methanol conversions can be well described with a bubble diameter factor fdb of 0.9 in case of 0% permeation and with fdb ) 0.8 in case of 40% permeation of gas through the membranes for all the experiments with different fluidization velocities and

catalyst bed heights. From the RTD curves measured with the ultrasound technique described in part 1, bubble diameter factors of 0.9-1.0 and 0.7-0.8 were deduced for 0 and 40% permeation for the relatively high fluidization velocities investigated here. Concluding, the bubble diameter factors derived from the measured methanol conversions correspond nicely with the bubble diameter factors deduced from the RTD measurements, especially considering the effect of the temperature on the bubble size. At higher temperatures somewhat smoother fluidization and smaller bubbles are expected.8 4.3.2. Methanol Conversion and Product Selectivities as a Function of Bed Height. In Figure 11 the model predictions for the methanol conversion and the product selectivities at various bed heights and various permeation ratios are compared with the experimental results. It is very clear from Figure 11 that


the computed and measured product selectivities to formaldehyde and carbon monoxide match quite nicely. Small discrepancies in the formaldehyde and carbon monoxide selectivities are mainly due to a slight overprediction of the DME selectivity, most probably caused by the fact that the kinetic rate expression for the DME formation reaction was not based on experimental data for high methanol concentrations, as employed in this study. Concluding, all the observed trends can be qualitatively and quantitatively described by the developed two-phase reactor model. 5. Summary and Conclusions The partial oxidation of methanol to formaldehyde over an industrial Fe-Mo catalyst has been experimentally investigated in a membrane-assisted fluidized bed reactor in order to demonstrate the reactor concept. A small laboratory-scale membrane-assisted fluidized bed reactor was constructed equipped with horizontal porous ceramic membranes and cooling tubes, in which the methanol conversion and the product selectivities toward formaldehyde, carbon monoxide, dimethyl ether, and dimethoxy methane were measured at various overall fluidization velocities, reactor temperatures, methanol and oxygen overall feed concentrations, ratios of gas added via the membranes, and distributor and aspect ratios of the fluidized bed. High methanol conversions and high selectivities to formaldehyde were achieved with safe reactor operation at very high methanol inlet concentrations, much higher than currently employed in industrial processes. It was experimentally demonstrated that distributive feeding of oxygen in a MAFBR produces an increased overall formaldehyde yield and throughput without pronounced undesirable conversion of formaldehyde to carbon monoxide for low aspect ratios of the catalyst bed. A one-dimensional two-phase phenomenological reactor model has been developed considering both the bubble phase and the emulsion phase as a series of ideally stirred tank reactors, where the number of tank reactors is related to the extent of gas back-mixing. From independent steady-state tracer injection experiments, plug flow behavior for both phases was concluded for the investigated lab-scale MAFBR (see part 1). In the two-phase model, the addition of gas via the membranes and the gas production due to chemical reactions have been accounted for. For a MAFBR neither experimental data nor literature correlations are available for the prediction of the bubble size as a function of the axial position in the bed and superficial gas velocity, which accounts for the permeation of gas through the membranes. A bubble diameter factor was defined as the factor with which the bubble diameter as calculated with the correlation proposed by Mori and Wen7 for fluidized beds without internals changes for the MAFBR. From the measured methanol conversion as a function of the gas residence time (varied by changing the fluidization velocity and the bed height), a bubble diameter factor of 0.9 in the case of 0% permeation and 0.8 in the case of 40% permeation of gas through the membranes was fitted, with which all the experimental results (methanol conversion and formaldehyde and CO selectivities) for different fluidization velocities and catalyst bed heights could be very well described. These bubble diameter factors match nicely with the bubble diameter factors deduced from independent RTD measurements using the ultrasound

technique described in part 1, where bubble diameter factors of 0.9-1.0 and 0.7-0.8 were found for 0 and 40% permeation for the relatively high fluidization velocities investigated here, especially in view of the temperature effect on the bubble size. With the developed two-phase model, which accounts for gas permeation through the membranes, all the experimentally observed influences of the operating conditions in the MAFBR could be described well and can therefore be used for further optimization of the MAFBR. Concluding, for the partial oxidation of methanol to formaldehyde in the MAFBR with a small aspect ratio, relevant for industrial applications, the distributive feeding of oxygen via porous membranes in a MAFBR offers the clear advantage of decreased axial gas backmixing and a high formaldehyde yield at very high methanol inlet concentrations (i.e., high throughput) with an inherently safe isothermal reactor operation. Acknowledgment This research is part of the research program carried out within the Center for Separation Technology, as a cooperation between the University of Twente and TNO, The Netherlands Organization for Applied Scientific Research. The authors thank Wim Leppink for the construction and maintenance of the experimental setups. Notation AT ) bed cross section, [m2] C ) concentration, [mol‚m-3] E ) activation energy, [kJ‚mol-1] F ) permeation fraction, [-] f ) bubble diameter factor, [-] Hf ) final fluidization bed height, [m] k ) reaction rate constant, [s-1] K ) adsorption constant Kbc ) volumetric interchange coefficient between bubble and cloud phase, [s-1] Kbe ) volumetric interchange coefficient between bubble and emulsion phase, [s-1] Kce ) volumetric interchange coefficient between cloud and emulsion phase, [s-1] eq K ) thermodynamic equilibrium constant, [atm-1] m ) mass of catalyst, [kg] nc ) number of components, [-] Nb ) number of compartments in the bubble phase, [-] Ne ) number of compartments in the emulsion phase, [-] Pi ) partial pressure of component i, [atm] R ) gas constant [)8.314 J‚mol-1‚K-1] Ri ) reaction rate of component i, [mol‚kg-1‚s-1] T ) temperature, [K] or [°C] u ) superficial gas velocity to the reactor, [m‚s-1] ub ) rise velocity of cloud of bubble, [m‚s-1] ubr ) rise velocity of single bubble, [m‚s-1] usb ) superficial bubble velocity, [m‚s-1] usb,0 ) initial superficial bubble velocity, [m‚s-1] s ub,max ) maximum superficial bubble gas velocity, [m‚s-1] umf ) minimum fluidization velocity, [m‚s-1] Vi ) volume of the ith phase, [m3] x ) gas-phase fraction, [-] Greek Letters δbo ) initial bubble phase fraction [-] δb ) bubble phase fraction [-] mf ) bed voidage at minimum fluidization conditions, [-] µg ) viscosity of gas, [Pa‚s]

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νi,j ) reaction stoichiometry constant for component i for reaction j, [-] Fg ) density of gas, [kg‚m-3] Fp ) density of fluidizing particles, [kg‚m-3] θi ) surface fraction of component i, [-] Subscripts and Superscripts 0 ) initial ads ) adsorption avg ) average b ) bubble phase be ) bubble phase to emulsion phase br ) bubble rise cat ) catalyst DME ) dimethyl ether DMM ) dimethoxy methane e ) emulsion phase eq ) thermodynamic equilibrium exp ) experimental f ) forward reaction, final height g ) gas phase i ) component i in ) inlet j ) component j max ) maximum mf ) minimum fluidization n ) nth p ) permeation r ) rise s ) superficial tot ) total

Literature Cited (1) Adris, A. M.; Elnashaie, S. S. E. H.; Hughes, R. A fluidized bed reactor for steam reforming of methane. Can. J. Chem. Eng. 1991, 69, 1061. (2) Adris, A. M.; Lim, C. J.; Grace, J. R. The fluid bed membrane reactor system: A pilot scale experimental study. Chem. Eng. Sci. 1994, 49, 5833. (3) Abdalla, B. K.; Elnashaie, S. S. E. H. Fluidized bed reactor without and with selective membranes for the catalytic dehydrogenation of ethylbenzene to styrene. J. Membr. Sci. 1995, 101, 31. (4) Ostrowski, T.; Giroir-Fendler, A.; Mirodatos, C.; Mleczko, L. Comparative study of the catalytic partial oxidation of methane to synthesis gas in fixed bed and fluidized bed membrane reactors. Part II: Development of membranes and catalytic measurements. Catal. Today 1998, 40, 191. (5) 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. (6) Deshmukh, S. A. R. K. Membrane Assisted Fluidized Bed Reactor: Experimental demonstration for partial oxidation of methanol. Ph.D. Thesis, University of Twente, Enschede, The Netherlands, 2004. (7) Mori, S.; Wen, C. Y. Estimation of bubble diameter in gaseous fluidized beds. AIChE J. 1975, 21, 109. (8) Kunii, D.; Levenspiel, O. Fluidization Engineering; Wiley, New York, 1991.

Received for review September 16, 2004 Revised manuscript received January 10, 2005 Accepted January 11, 2005 IE049092H

Development of a Membrane-Assisted Fluidized Bed Reactor. 2

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