394
Ind. Eng. Chem. Res. 1989,28,394-400
Performance of a Honeycomb Monolith Bioreactor in a Gas-Liquid-Solid Three-phase System Koei Kawakami,* Kyosuke Kawasaki, Fumihide Shiraishi, and Koichiro Kusunoki Department of Chemical Engineering, Kyushu University, Fukuoka 812, Japan
The gas-liquid and liquid-solid mass-transfer rates in a monolith reactor were measured with three different channel sizes, i.e., 12, 31, and 62 cells/cm2, and a t various gas and liquid flow rates. In the range of measurements, the volumetric coefficients for both mass-transfer steps were higher in the case of gas-liquid cocurrent upflow operation than in the downflow operation. The effect of channel size was less pronounced. The oxidation of glucose by immobilized glucose oxidase was investigated, and the characteristics of the monolith reactor were analyzed. The conversion of glucose was higher with the monolith reactor involving a larger number of channels, due to thinner walls leading to a reduction of resistance to internal diffusion of oxygen. Model calculation indicated that, in the upflow operation, the overall effectiveness factor of the reactor loaded with monolith 62 cells/cm2 was more than 0.3, while in the downflow operation the overall effectiveness factor was less than 0.1. It turned out that, in the latter operation, the conversion was influenced by incomplete wetting of the monolith wall. Recently, the use of a ceramic honeycomb monolith has been widely investigated as a support for nobel metal catalysts such as platinum (El Sawi et al., 1985) and palladium (Hatziantoniou and Andersson, 1984; Hatziantoniou et al., 1986) and for biocatalysts such as enzymes (Benoit and Kohler, 1975; Shiraishi et al., 1983,1986,1988), microorganisms (Ghommidh et al., 1982; Ariga et al., 1986; Shiraishi et al., 1989), and animal cells (Lydersen et al., 1985). The monolith reactor may have several potential advantages over a conventional particulate packed-bed reactor (Satterfield and Ozel, 1977): (1)high mechanical strength, (2) large pore size and large specific surface area, (3) thin walls, (4) better liquid distribution at low liquid flow rates, and ( 5 ) low pressure drop. In addition, the mass-transfer to the channel wall from the liquid was found to be much greater for a segmented gas-liquid flow than for a continuous liquid flow (Horvath et al., 1973; Hatziantoniou and Andersson, 1982). Particular interest was thus focused on the possible application of the honeycomb monoliths to three-phase catalytic reactors with gas-liquid cocurrent flow through a bed of solid catalyst. Hatziantoniou and Andersson (1984) and Hatziantoniou et al. (1986) studied liquid-phase hydrogenation of nitro compounds in a monolithic palladium catalyst reactor and reported that the direct mass transfer of hydrogen from the gas plugs to the channel wall was the dominating transport step. Those authors evaluated that this direct transport corresponded to more than 70% of the total transport of hydrogen to the channel wall. The ceramic honeycomb monolith has also been utilized as a support for immobilization of certain microorganisms. In particular, its open structure makes it possible to perform bioconversions in gas-liquid-solid three-phase systems. Ariga et al. (1986) investigated the characteristics of the monolith bioreactor in which Escherichia coli or Saccharomyces cerevisiae was immobilized within a thin K-carrageenan gel film on a channel wall. Those authors indicated that the monolith reactor was effective in the reaction systems accompanied by gas evolution such as ethanol fermentation. Ghommidh et al. (1982) studied the production of acetic acid by immobilized Acetobacter cells in a monolith reactor with a pulsed flow. It was shown that the high oxygen transfer capability of the reactor enabled a very high productivity of acetic acid. In our recent study
* To whom correspondence should be directed. 0888-5SS5/89/262~-0394$01.50/0
Table I. Structural Properties of the Ceramic Honeycomb Monoliths Used 80 400 no. of cells per square inch 200 12.4 62.0 channel size, cells/cm* 31.0 0.11 side length of square channel, cm 0.24 0.15 0.045 0.030 0.018 wall thickness, cm 1.51 1.46 1.56 apparent density, g/cm3 0.44 0.44 0.50 bulk density, g/cm3 11.8 26.9 18.5 specific surface area, cm-I 0.71 0.68 0.70 volume fraction of channel 0.39 0.37 0.41 porosity of wall
(Shiraishi et al., 1989), a strain of Gluconobacter was immobilized on a ceramic honeycomb monolith, and aerobic transformation of glucose to gluconic acid was extensively investigated. The high productivity of 26.1 kg.m-3.h-1 was achieved with stable activity of the immobilized cells during a long-term operation. In a previous report (Kawakami et al., 1987), some fundamental characteristics of the three-phase monolith reactor were examined, and it was shown that the gasliquid cacurrent upflow operation was superior to the cocurrent downflow operation, due to higher mass-transfer rates in the former operation. This paper presents a further study on mass-transfer characteristics and a model of the reactor performance using oxidation of glucose by immobilized glucose oxidase as the test reaction.
Experimental Section The ceramic monolith, made of cordierite, was supplied by NGK Insulators Ltd. (Nagoya, Japan). In this experiment, monoliths with three different channel sizes, i.e., 80, 200, and 400 channels/in.2, were used. Each of them was cut into small rectangular blocks (20 mm X 20 mm x 5C-150 mm long) containing 49,121, and 256 channels per cross section. The structural properties of those specimens, referred to as monolith 80, monolith 200, and monolith 400, are listed in Table I. The enzymes were purchased from Sigma Chemical Co. The glucose oxidase was G 6500 from Aspergillus niger, the catalase was C 3515 from Aspergillus niger, and the trypsin was T 0134 from bovine pancreas. The method used to immobilize the enzymes was as follows (Shiraishi et al., 1988): The monolith support was first treated in an 1.0 k m ~ l - mH,S04 -~ solution under refluxing for 5 h and then soaked in an 8% solution of y(aminopropy1)triethoxysilane in toluene under refluxing for 5 h, washed with 1989 American Chemical Society
Ind. Eng. Chem. Res., Vol. 28, No. 4, 1989 395 0.2 I
1
GAS& LIQUID OUTLET
-
D D’ Cross Section i
;ID’
0.083 0.080
...._..,
-
0.5
C C’ Cross Section
1.0
2.0 5.0 ug [cm.s-’]
10.0
. .-
B -6’Cross Section
-
v)
LD m
I 1 Block of Monolith 80
I
=
4LlQUlD INLET
&
/ lw
across
Spction
.L
0.1
--
- 0.05v
)
E
x
-
-
0.02 0.01 :
q?t?T
49 Stainless Capillary Tubes
0.005 LGI
b
Figure 1. Configuration of honeycomb monolith reactor in cocurrent upflow operation.
toluene, and dried overnight. The enzyme solution containing 110 gm-3 of glucose oxidase and 110 gm-3 of catalase in a 100 m o l ~ m -phosphate ~ buffer solution (pH 6) was contacted with 2.5 % glutaraldehyde for cross-linking at 279 K for 2 h. The silanized monolith support was then soaked in the solution containing the cross-linked enzymes. The immobilized-enzyme monolith was prepared by stirring the solution for 24 h at 279 K. The immobilization of trypsin was performed in the same way, except that the concentration of enzyme was 5 kgm-3 and the pH of the solution was 8. The immobilized-enzyme monolith was loaded in a rectangular acrylic column with an inside dimension of 22 mm X 22 mm and a length of 220-330 mm. Figure 1shows the configuration of a monolith reactor for gas-liquid cocurrent upflow operations. In this operating mode, the lower half of the column serves as a gas distributor unit. Forty-nine stainless steel capillary tubes (0.25-mm i.d.) were inserted into all the channels at the bottom of a monolith 80 block (35-mm height). The gas was first introduced uniformly through 49 channels of monolith 80 and then was redistributed through 20 slices of monolith 400 (5 mm thickness and 256 channels each). The latter stack consisted of an alternate arrangement of monolith 400, with two different arrays of channels deviating by n/4 rad from each other around the main axis. The substrate solution was supplied continuously through the liquid inlet by a peristaltic pump. In the cocurrent downflow configuration,the whole apparatus was inverted; the substrate solution was supplied through the stainless steel capillaries, while the gas entered through the inlet used for the liquid feed in the upflow operation. The oxidation of a 20 mol~m-~ aqueous solution of glucose was carried out at 298 K and pH 6 (100 phosphate buffer solution). The concentration of glucose in the effluent was determined by the Somogyi-Nelson method. The hydrolysis of N-benzoyl-L-arginine ethyl ester (BAEE) to N-benzoyl-L-arginine (BA) by immobilized trypsin was used to determine the liquid-solid masstransfer coefficient in the monolith reactor. The experi-
0.005 0.01
0.02 u,
0.05 0.1
0.2
0.5
[cm.s-’I
Figure 2. Effects of superficial (a) gas and (b) liquid velocities on volumetric gas-liquid mass-transfer coefficient in monolith reactor with different channel sizes.
ment was carried out by flowing nitrogen as the inert gas and 1 BAEE solution at 298 K and pH 8 (100 mol~m-~ phosphate buffer solution). The concentration of the substrate in the effluent was measured spectrophotometrically at 253.6 nm. The volumetric gas-liquid mass-transfer coefficient in the monolith reactor was determined from the measurement of the desorption rate of oxygen from water into the nitrogen stream.
Results and Discussion In a previous report (Kawakami et al., 1987),we showed that, at low liquid flow rates, the two-phase flow in the monolith channels was a slug flow in the case of cocurrent upflow and was a liquid film flow in the case of cocurrent downflow. The residence-time distribution of the monolith reactor, as determined by the pulse-response method with glucose as the tracer, indicated almost perfect mixing in the liquid phase during the gas-liquid cocurrent upflow operation. Higher gas flow rate as compared to liquid flow rate might cause vigorous mixing of the liquid in the axial direction of each monolith channels. Gas-Liquid Mass Transfer. Parts a and b of Figure 2 show the effects of superficial gas and liquid velocities, respectively, on the volumetric gas-liquid mass-transfer coefficients, klab, in the cocurrent upward and downward flow operations. The upflow operation gave the values of k l a b as 2-5 times higher than the downflow operation. This might be attributable to a larger gas-liquid interfacial area in the upflow operation. The effect of channel size was less pronounced in both operations. The shaded area in Figure 2b indicates the range of k l a b in the trickel-bed reactor (Ramachandranand Chaudhari, 1983) packed with particles of 2-mm diameter, which has a specific surface area equivalent to that of the honeycomb monolith used. The monolith reactor exhibited a much higher k l a b than
396 Ind. Eng. Chem. Res., Vol. 28, No. 4, 1989 10
MONOLITH LW
-e-.
g.-*-*
MONOLITH 400 DOWNFLOW u, :0029 cmslI
0 0.2
0.5
1
2
5
3
1
2
3
4
5
ug [ c m s ” ]
19
ug I c m s-’I
Figure 5. Effect of Superficial gas velocity on conversion of glucose in immobilized-enzyme monolith reactor.
DOWNFLOW
I A A C.
05 0002
O W 5 001
002
u,
n
5 08 005 0 1 [cm s ’1
02
Figure 3. Effects of superficial (a) gas and (b) liquid velocities on Sherwood number of liquid-solid mass transfer in monolith reactor with different channel sizes.
1/2O - 0.02 < I 0.01 ; 1
0.005~
t
0.002 0.002
0
2
0.02
uI
0.05
01
02
0.5
[crn.~-~l
Figure 4. Effect of superficial liquid velocity on volumetric coefficient for liquid-solid mass transfer of oxygen.
the usual particulate packed-bed reactor. Liquid-Solid Mass Transfer. Horvath and Solomon (1972) and Horvath et al. (1973) studied the hydrolysis of N-benzoyl-L-arginine ethyl ester (BAEE) by immobilized trypsin on porous walls in an open tubular heterogeneous enzyme reactor and found that the conversion was controlled by the film diffusion of the substrate. The same reaction was utilized to measure the liquid-solid masstransfer coefficients in the monolith reactor. The Sherwood number, Sh (=k,d/D), was calculated from the measured conversion of BAEE, assuming perfect mixing of the liquid phase in the case of cocurrent upflow operation and plug flow for the downflow operation (Kawakami et al., 1987). Parts a and b of Figure 3 show the effects of superficial gas and liquid velocities, respectively, on Sh. The upflow operation is again advantageous, giving a 4-5 times higher mass-transfer rate than the downflow operation. Sh was less influenced by the superficial gas velocity but was in-
6
8
Lkg.s.mol-’l
Figure 6. Conversion of glucixe versus W/F in immobilized-enzyme monolith reactor with different channel sizes.
creased in proportion with the 0.3 power of the superficial liquid velocity in the upflow operation and with the 0.4-0.8 power in the downflow operation. Figure 4 shows the volumetric liquid-solid mass-transfer coefficients of oxygen, k,a,, redrawn from Figure 3b assuming that Sh is proportional to S C ~ ’The ~ . channel size also had a minor influence on the liquid-solid mass-transfer rates. Oxidation of Glucose by Immobilized Glucose Oxidase. The oxidation of glucose (G) by glucose oxidase in the presence of catalase in sufficient excess proceeds without a gross accumulation of hydrogen peroxide and is represented by
G 0.005 0.01
4 W / F x lo-‘
05
+ f/zO,
-
L
+ H,O
(1)
where L stands for glucono-&lactone. The intrinsic reaction rates were measured in a closed stirred slurry reactor involving a suspension of the crushed immobilized enzyme monolith, by following the decrease in the concentration of dissolved oxygen. Calculation (Lee et al., 1981) indicated that the intraparticle diffusion effect was negligible with catalytic effectiveness factors larger than 0.9 under the experimental conditions studied. The kinetic data obtained for different concentrations of glucose and oxygen were expressed by the following rate expression (Gibson et al., 1964; Weibel and Bright, 1971; Linek et al., 1980):
The kinetic parameters were determined as k,, = 1.44 and KG , = 0.134 mmol-kg-l-s-’, KmO = 1.03 mol~m-~, kmol~m-~. Figure 5 shows the effect of superficial gas velocity, ug, on the conversion of glucose in the monolith 400 reactor. The conversion in both the upflow and downflow operations was unaffected by ug over 2 ems-'. Figure 6 illustrates the relation between the conversion measured as a function of liquid flow rate and the ratio of mass of the
Ind. Eng. Chem. Res., Vol. 28, No. 4, 1989 397 immobilized enzyme monolith to molar flow rate of glucose, WIF, in the monolith 80,200, and 400 reactors. For every monolith, the upflow operation gave a higher conversion than the downflow operation. This is due to higher mass-transfer rates between gas and liquid and between liquid and solid in the upflow than in the downflow. In both operations, the monolith reactor with a smaller channel size or a larger number of channels led to higher conversion. Probably, this may be ascribed to the reduction of internal diffusion resistance of oxygen, because the monolith with a smaller channel size has thinner walls. Simulation of Conversion in Honeycomb Monolith Reactor. To explain the experimental results mentioned above more quantitatively, the simulation was executed based upon the mathematical model. For this purpose, the batchwise oxidation of glucose was performed in a rotating basket reactor loaded with small fragments of the immobilized-enzyme monolith. The effective diffusivity of oxygen, DeO,was estimated as 1.1X cm2.s-l, from the analysis of the time-course data for the combined action of the enzyme reaction and internal diffusion of oxygen. A model for the monolith reactor for cocurrent upflow and downflow of gas and liquid is developed based on the following assumptions: 1. Oxygen is the limiting reactant. 2. The reaction rates on the immobilized-enzyme monolith are expressed by eq 2. 3. For upflow operation, the gas stream in each channel is in plug flow and the liquid phase behaves as axially well-mixed flow. 4. For downflow operation, both the gas and liquid phases move in plug flow. 5. The concentration of oxygen in the gas phase is constant throughout the reactor. 6. For plug flow of both the gas and liquid with downflow, the concentration variation of dissolved oxygen in the liquid phase along the reactor is negligible; i.e., dynamic equilibrium between gas and liquid holds. 7. The external surface of monolith wall with liquid-filled pore is completely wetted. Then, the steady-state mass balance equation for oxygen can be derived for both the upflow and downflow operations as follows (Appendix 1): klab(Co* - cia) = k,a,(Cio - cso) = (W/ud)?l,(-ro) (3)
For the cocurrent upflow (liquid well mixed), the mass balance equation for glucose is given by (Appendix 1) U](CGi - cG,)/z = 2(W/Ud)qo(-rO) (4) while the mass balance for the cocurrent downflow (liquid in plug flow) results in the following equation (Appendix 1):
-ui dCG/$z = ~ ( W / U ~ ) V , ( - ~ O )
(5) In eq 4 and 5, qo indicates the overall effectiveness factor involving the effects of external diffusion of oxygen from the gas-liquid interface to the catalyst surface as well as internal diffusion within the porous monolith wall and is given by Yamane’s approximate expression (Yamane, 1981) for the Michaelis-Menten kinetics on the immobilized biocatalyst with infinite slab geometry (Appendix 2). The conversions of glucose at different liquid flow rates were computed by solving eq 4 or 5 numerically. The calculated curves are shown by the dotted lines in parts a-c of Figure 7. In the case of upflow operation, the calculated curves agreed rat,her well with the experimental results except for the higher conversion region. However, in the case of downflow operation, the observed conversions were much higher than the calculated values (dotted lines). A possible reason for the enhanced conversion might be incomplete wetting of the catalyst surface (invalidity of assumption 7), as often encountered in trickle-bed reactors
MONOLITH 80
-
UPFLOW
1
MONOLITH 400
0‘
10)
10‘ W/F I k g ~ s . m o l ” 1
io5
I
Figure 7. Comparison between calculated curves and experimental results for conversion of glucose in (a) monolith 80, (b) monolith 200, and (c) monolith 400 reactors.
operated at low liquid flow rates (Tsukamoto et al., 1982; Goto and Mabuchi, 1984). If part of the outer surface not covered by liquid rivulets exists on the honeycomb walls, the gaseous reactant is more easily supplied through the nonwetted fraction of the catalyst surface. For such a situation, the overall effectiveness factor involving influence of incomplete wetting is approximated by the weighted average value for the wetted and nonwetted surface fractions, f and 1- f , respectively (Appendix 1): In this equation, the effectiveness factor for the wetted surface fraction, T,,, is defined in terms of Co* by taking the effects of gas-liquid and liquid-solid mass-transfer and intraparticle diffusion into account and is given by Yamane’s approximate equation (Yamane, 1981). Meanwhile, qd is based upon CEO= CO* (by assuming negligible resistance to gas-solid mass transfer on the nonwetted surface) and is expressed by the approximate equation (Appendix 2) of Kobayashi et al. (1976). Thus, 7, in eq 3 is replaced by fq,,, and qo in eq 5 is expressed by eq 6 (Appendix 1). The resultant equations were solved numerically with different values off. For monoliths 80,200, and 400, the calculated conversionswith f = 0.9,0.93, and 0.93, respectively, agreed reasonably well with the observed conversions (the solid lines for the downflow in parts a, b, and c of Figure 7). Even in the case of upflow operation, the observed values for monoliths 200 and 400 in the low conversion region tended to slightly exceed the calculated curves (the dotted lines for the upflow operation in parts b and c of Figure 7). Naturally, the difference between them was small and may be within the range of errors. However, in this case, there is the possibility of direct contact between gas
398 Ind. Eng. Chem. Res., Vol. 28, No. 4, 1989
bubbles and the honeycomb wall (Goto and Mabuchi, 1984; Hatziantoniou and Andersson, 1984; Hatziantoniou et al., 1986). The effect of incomplete wetting was likewise examined on the conversion in the upflow operation. As a result, it was found that the calculated curves with f = 0.93 seemed in better agreement with the data (the solid lines for the upflow operation in parts b and c of Figure 7 ) . On the contrary, in the region of higher conversion, the observed conversions tended to be less than the calculated values. A possible reason for this would be a drop in pH due to the formation of gluconic acid. In fact, an increase in the concentration of phosphate buffer solution from 100 to 500 m0l.m" resulted in an increase in glucose conversion by 7-870 at the lower liquid flow rates. The model calculation indicated that in the case of cocurrent upflow operation with the monolith 400 reactor, the overall effectiveness factor, qo, was more than 0.3, while in the case of downflow operation, qo was less than 0.1, owing to severe external mass-transfer limitation.
pp = apparent density, kg.m-3 4 = Thiele modulus defined by eq A-17, dimensionless 4c = critical value of C$ for qo in eq A-14, dimensionless
Subscripts d = dry (nonwetted, gas covered) e = exit G = glucose g = gas phase i = inlet 1 = liquid phase 0 = oxygen o = overall s = surface w = wet (liquid covered)
Appendix 1: Reactor Model Development For a three-phase monolith reactor with cocurrent flow of gas and liquid, a conventional model, similar to that used for a trickle-bed reactor in the case of gaseous reactant limiting and incomplete wetting of catalyst surface Nomenclature (Mills and Dudukovie, 1984), was used in this study. (i) Plug Flow of both Gas and Liquid with Cocura, = external surface area of monolith based on volume of rent Downflow. In this case, the steady-state mass channels, 4/d, m-l balance for each of the three phases in the reactor leads Bi = Biot number, dimensionless to the following equations: Cdso= concentration of oxygen on nonwetted surface,mol" CG = concentration of glucose, m ~ l . m - ~ gas phase C,o = concentration of oxygen in gas stream, m ~ l . m - ~ Co* = concentrationof oxygen at gas-liquid interface,m ~ l - m - ~ -ugcgO' - klab(CgO/HO - ClO) Hokda,(Cgo/Ho - CdsO) = 0 (A-1) Clo = concentration of oxygen in bulk liquid, mol~m-~ Cso = concentration of oxygen on wall surface of monolith, liquid phase m01.m-~ Cwso = concentration of oxygen on wetted surface, mol" -UIC]O' + k,ab(Cgo/Ho - ClO) - kcac(Cl0 - CwsO) = 0 D = molecular diffusivity, m2.s-l (A-2) DeO = effective diffusivity of oxygen, m2-s-l -U]CG' - 2f~w(w/~d)(-rO(cwsO~ CG)) d = side length of square channel, m 2(1 - f)Vd(W/Ud)(-ro(CdsO, CG)) = 0 (A-3) F = molar flow rate of glucose, moles-' f = wetting efficiency, dimensionless solid phase Ho = Henry's constant for oxygen, dimensionless (A-4) kcac(Cl0 - CwsO) = f~w(w/ud)(-rO(cwsO, C G ) ) KM = apparent Michaelis constant, m ~ l - m - ~ KmG = constant in eq 2, mol~m-~ HOkdac(CgO/HO - CdsO) = KmO= constant in eq 2, m ~ l e m - ~ (1- f)qd(W/Ud)(-rO(CdsO, CG)) (A-5) KO= overall mass-transfer coefficient, m d k, = liquid-solid mass-transfer coefficient, m d where ' means d/dz. kcat = reaction rate constant defined by eq 2, mold.kg-' If we use assumption 4 in the text and postulate that k d = gas-solid mass-transfer coefficient, m 4 ' resistance to direct mass transfer of oxygen from gas phase k p b = volumetric gas-liquid mass-transfer coefficient, sS1 to nonwetted surface of the catalyst is negligible, eq A-1 L = half wall thickness of monolith, m and A-5 become unnecessary, and CdsO in the kinetic exm = generalized modulus defined by eq A-11, dimensionless pression for eq A-3 is replaced by Co* (=Cfl/Ho). A more -ro = reaction rate of oxygen, moldkg-' simplified model is derived by assumption 5 in the text Sh = Sherwood number, k,d/D, dimensionless and then by combining eq A-2 and A-4 to give ug = superficial gas velocity based on cross-sectional area of klab(CO* - ClO) = kcac(Cl0 - CwsO) = Koac(CO* - CwsO) = channel, m-s-' ul = superficial liquid velocity based on cross-sectional area f~w(W/ud)(-ro(Cwso,CG)) (A-6) of channel, mas-' where l/Koa, = l/k$b + l/k,a,. If the catalytic effecVM = apparent maximum rate, mol-s-l-kg-l tiveness factor for liquid-covered surface fraction, tW, is ud = volume of channels, m3 replaced by the overall effectiveness factor, qwo, involving W = mass of immobilized-enzyme monolith, kg both internal and external diffusion effects, Cwsoin the 2 = height of stacked monolith, m kinetic expression for eq A-3 and A-6 is replaced by Co*, z = axial coordinate, m and eq A-3 is rewritten as follows: Greek Symbols -UiCG' = 2(fqw0+ (1 - f)Vd)(W/ud)(-ro(Co*, CG)) (A-7) qd = catalytic effectiveness factor for dry surface fraction, (ii) Gas in Plug Flow and Liquid Well Mixed with dimensionless qo = overall effectiveness factor, dimensionless Cocurrent Upflow. The steady-state mass balance qw = catalytic effectiveness factor for wet surface fraction, equations for the gas and solid phases are given in the same dimensionless forms as eq A-1, A-4, and A-5, and those for the liquid qwo = overall effectiveness factor for wet surface fraction, phase can be written as dimensionless qo = effectiveness factor for zero-order reaction, dimensionless ql = effectiveness factor for first-order reaction, dimensionless K = dimensionless Michaelis constant
Ind. Eng. Chem. Res., Vol. 28, No. 4, 1989 399 -ul(cG,e
- CG,i)
- 2f~w(~/ud)(-rO(cwsO, cG))z-
Jz2(1 - f)17d(W/Ud)(-r0(Cds0, CG))dz = 0 (A-3')
For the enzyme kinetics given by eq 2 in the text, the apparent maximum rate, V M , and Michaelis constant, KM, are expressed as a function of CG,
Under assumption 4 and the assumption that the contribution of the first term in eq A-2' is negligible due to low solubility of oxygen in the liquid, eq A-6 applies again. In a similar manner as in case i, if CdsO is given by its saturated value and qwo is used in place of qw in eq A-3', one can get the following equation: =
UI(CG,i - c G , e ) / z
2(ftlw0 + (1 - f)Td)(W/ud)(-rO(Co*, CG)) (-4-7') For calculating the conversion of glucose at different liquid flow rates, eq A-7 and A-7' were solved numerically by the Runge-Kutta-Gill method and the Regula-falsi method, respectively.
Appendix 2: Determination of Effectiveness Factor of Monolith Biocatalyst The effectiveness factors included in the eq A-7 and A-7' were estimated by using the approximate expressions for infinite slab geometry developed by Kobayashi et al. (1976) and Yamane (1981) for the Michaelis-Menten kinetics. Kobayashi's approximate expression (Kobayashi et al., 1976) exhibits the highest accuracy when the external film diffusional resistance is ignored and is given by qo 2.6~~.'71 v= 1 2.6~O.~ where
+ +
(A-9) q 1 = (tanh
m=
m)/m
h - K In (1
(1 + K)(I
(A-10)
+ I/K))'/'
(A-11)
and (A-12) K
=
(A-13)
KM/CsO
Yamane (1981) extended eq A-8 so as to include the estimation of the overall effectiveness factor involving influences of both internal pore and external film diffusional resistances. In this case, the parameters in eq A-8 are defined by I 1 (A-14)
-1
(A-15)
tanh 4
GC =
(+ L1 )
2l/ BJi
2
(A-16)
(A-17) (A-18) (A-19)
(A-20) (A-21) Thus, the catalytic effectiveness factor, qd, for the gascovered surface fraction was determined from eq A-8-A-13, while the overall effectiveness factor, qwo, for the liquidcovered surface fraction was determined from eq A-8 and eq A-14-A-19. Registry No. 02,7782-44-7; glucose oxidase, 9001-37-0.
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Received for review June 9, 1988 Accepted December 5, 1988
Synthesis of Phthalic and Maleic Anhydrides from n -Pentane. 1. Kinetic Analysis of the Reaction Network Gabriele Centi, Jose Lopez Nieto,+Davide Pinelli, and Ferruccio Trifirb" Department of Industrial Chemistry and Materials, University of Bologna, V.le Risorgimento 4, 40136 Bologna, Italy
The kinetics of n-pentane oxidation over a vanadyl pyrophosphate catalyst is described by a Langmuir-Hinshelwood mechanism; four parallel reactions leading to maleic anhydride, phthalic anhydride, CO, and CO,; and two consecutive reactions for the formation of carbon oxides from the two anhydrides. The rate-determining step is a surface reaction between one adsorbed n-pentane molecule and one oxygen molecules, indicating that the reaction leading to C-C bond formation in phthalic anhydride synthesis occurs after the rate-determining step. Phthalic anhydride selectivity is higher for the lower reaction temperatures and n-pentane or oxygen concentrations. The relevance of the kinetic information with regard to the analysis of the mechanism for the formation of the C8 anhydride from the C5alkane is also discussed. The C5 cut is a relatively low-cost hydrocarbon stream in the oil-refining industry, and the straight-chain fraction in particular is not specifically utilized in the petrochemical industry (Weissermel and Arpe, 1978). Therefore, there is moderately large interest in the development of new routes to upgrade the value of these hydrocarbons, especially n-pentane. This alkane is one of the principal components of the C5 cut together with cyclopentadiene and isoprene, but its functionalization is much more difficult. Heterogeneous catalytic oxidation processes are a powerful method for the functionalization of raw hydrocarbons (Chinchen et al., 1987; Hucknall, 1974), which have been successfully applied in recent years also to the conversion of light alkanes. The large-scale application of the process of n-butane selective oxidation to maleic anhydride is a typical example. The active phase for this reaction is vanadyl pyrophosphate (Centi et al., 1988; Hodnett, 1985; Busca et al., 1986). Recently we have shown (Centi et al., 1987; Centi and Trifirb, 1987) that, by the use of vanadyl pyrophosphate as the catalyst, n-pentane can be selectively transformed to phthalic and maleic anhydrides. The global selectivity is comparable to that of maleic anhydride in the oxidation of n-butane. Using different catalysts (supported Mo/V/P mixed oxides), Honicke et al. (1987a,b) also have found the formation of maleic and phthalic anhydrides from C.5 hydrocarbons, particularly from cyclopentene. The synthesis of the C8 anhydride from the C5hydrocarbon
* To whom correspondence
should be addressed. On leave from the Instituto d e Catalisis y Petroleoquimica, Serrano 119, 28006 Madrid, Spain. 0888-5885/89/2628-0400$01.50/0
is a new type of oxidation reaction involving the formation of multiple C-C bonds and aromatization, an unusual effect in the presence of gaseous oxygen and an oxidation catalyst (Chinchen et al., 1987; Hucknall, 1974). Previous investigations on the oxidation of C5 hydrocarbons have been focused on the conversion of branched- or straightchain pentenes on V205 (Butt and Fish, 1966a,b; Butt et al., 1966), of l,&pentadiene on CuO and cobalt molybdate catalysts (Mattson and Sasser, 1984), of C5 olefins on V,05-Mo03 catalysts (Seiyama et al., 1977) or on other mixed oxides of transition metals (Hucknall, 1974). In all cases, the formation of reaction products with a number of carbon atoms higher than the starting hydrocarbon was never observed. It is thus interesting to study the kinetic behavior of n-pentane oxidation on vanadyl pyrophosphate in order to learn more about the dynamic aspects of this new type of heterogeneous synthesis by selective oxidation. In addition to this specific aspect, a more general interest lies in the study of the kinetic and mechanistic characteristics of alkane functionalization, in order to understand the key factors responsible for the ability of the catalysts to activate and selectively convert paraffin feedstocks. In recent years, more attention has been focused on these problems, both for the purpose of upgrading basic knowledge regarding surface reaction mechanisms and for the more practical aspect of reducing costs in the substitution of olefins or aromatics with paraffinic feedstocks. This paper presents a systematic study of the kinetics aspect of the reactions of formation of phthalic and maleic anhydrides from n-pentane using an highly active/selective vanadyl pyrophosphate catalyst for n-butane oxidation to 0 1989 American Chemical Society
