Znd. Eng. Chem. Res. 1991,30, 1413-1418 Waller, K. V.; Makila, P. M. Chemical Reaction Invariants and Variants and Their Use in Reactor Modeling, Simulation, and Control. Znd. Eng. Chem. Process Des. Dev. 1981,20,1-11. Williams, G. L.; Rhinehart, R. R.; Riggs, J. B. In-Line ProcessModel-Based Control of Wastewater pH Using Dual Base Injection. Znd. Eng. Chem. Res. 1990,29, 1254-1259. Wong, S.K.P.Control and Modeling Strategies for Nonlinear Systems with Time Delays. Ph.D. Thesis, University of California, Santa Barbara, 1985. Wong, S, K. P.; Seborg, D. E. Low-Order, Nonlinear, Dynamic Models for Distillation Columns. In Proceedings of the 1986 American Control Conference,Seattle, WA, 19&, pp 1192-1198. Wong, S. K. P.; Seborg, D. E. Control Strategies for Nonlinear Multivariable Systems and Time Delays. In Proceedings of the 1986 American Control Conference, Seattle, WA, 1986b; pp 1023-1024. Wright, R. A,; Kravaris, C. Nonlinear Control of pH Processes Using the Strong Acid Equivalent. Znd. Eng. Chem. Res. 1991,in press.
1413
See also Wright, R. A.; Kravaris, C. Nonlinear pH Control in a CSTR. Proceedings of the 1989 American Control Conference, Pittsburgh, PA, 1989; pp 1540-1544. Wright, R. A.; Kravaris, C. Nonminimum Phase Compensation for Nonlinear Processes. Presented at the AIChE Annual Meeting, Chicago, IL, 1990; paper 238e. Wright, R. A.; Soroush, M.; Kravaris, C. Strong Acid Equivalent Control of pH Processes: An Experimental Study; In Proceedings of the 1991 American Control Conference, Boston, MA, 1991;in press. Yeo, Y. K.; Williams, D. C. Bilinear Model Predictive Control. Znd. Eng. Chem. Res. 1987,26,2267-2274. Zafiiiou, E. Robust Model Predictive Control of Procesees with Hard Constraints. Comput. Chem. Eng. 1990,14,359-371.
Received for review August 8, 1990 Revised manuscript received January 15, 1991 Accepted February 1, 1991
KINETICS AND CATALYSIS Continuous Operation of the Berty Reactor for the Solvent Methanol Process Chandrasekhar Krishnan and J. Richard Elliott, Jr.* Department of Chemical Engineering, The University of Akron, Akron, Ohio 44325-3906
Jozsef M.Berty Berty Reaction Engineers Ltd., 1806 Bent Pine Hill, Fogelsville, Pennsylvania 18051-9712
In the solvent methanol process (SMP), an inert and selective solvent removes methanol as soon as it is formed from syngm. Conversion in the conventional vapor-phase methanol synthesis is limited because of equilibrium limitations due to the reverse reaction, but data presented in this paper demonstrate that high conversions can be obtained in the SMP. Rate data have been collected for the SMP a t operating conditions typical of the vapor-phase process (7.8-10 MPa, 493-513 K). Single-pass H2 and CO conversions range from 30 to 80%. In some cases, conversions are higher than those predicted by vapor phase equilibrium calculations based on the feed composition, providing that SMP is able to overcome the equilibrium barrier. Rates are 2-3 times lower than those encountered in the vapor-phase process owing to pore diffusion limitations from the presence of the liquid. 1. Introduction Methanol is synthesized catalytically from H2,CO, and COO(synthesis gas). The following are the main reactions: CO + 2H2 * CHSOH (1) C02 + H2 + CO
+ H2O
(2)
The catalyst typically preferred is Cu/ZnO/A1203,which exhibits optimum activity in the temperature range 473-543 K (Chinchen et al., 1988). The strong exothermicity of reaction 1 (-91 kJ/mol) implies that the equilibrium CH30H content decreases with increasing temperature. In order to obtain reasonable rates and conversions, operating pressures in the range 5-10 MPa (50-100atm) are required. Single-pass conversions of CO and H2 are still only around 15-30%, resulting in the recycle of large quantities of unconverted reactants. This
* T o whom correspondence should be addressed.
leads to substantial fixed and operating costs in a CH30H plant. The initial concept of a three-phase CH30H process was put forward by Chem Systems Inc. (Sherwin and Blum, 1975). In that process, the use of a high heat capacity, inert paraffinic oil in the synthesis loop affords good temperature control of reaction 1 but product CH30H is removed from the vapor phase and single-pass conversions of H2 and CO are only slightly better than in the vapor-phase process. Recently, the use of two novel converters to improve syngas conversion to CH30H has been reported. In the gas-solid-solid trickle flow reactor (GSSTFR) (Westerterp et al., 19871, a fine adsorbent powder selectively picks up CH30H as soon as it is formed, thereby decreasing the equilibrium limitation on the forward reaction. The reactor system with interstage product removal (RSIPR) (Westerterp et al., 1988),on the other hand, achieves a high conversion per pass by using a series of Lurgi type reactors with CH30H adsorbers installed after each reactor.
0888-5885/ 91/ 2630-1413$02.50/ 0 0 1991 American Chemical Society
1414 Ind. Eng. Chem. Res., Vol. 30, No. 7, 1991 SEPARATOR GAS.L1oU1D
TC DIFFERENTIAL PRESSURE CELL
TC
8PR
FPR
GAS TO VENT
MFC PUMP
SYNGAS STORAGE
FLOW METER
PO
PRESSURE GAUGE RECEIVER
FPR FORWARD PRESSURE REQULAT'OR BPR BACK PRESSURE REGULATOR MFC MASS FLOW CONTROLLER TC
THERMOCOUPLES
GC
GAS CHROMATOGRAPH
Figure 1. Schematic of the experimental system.
In the SMP, a solvent flows directly over the catalyst bed and preferentially absorbs CH30H. This improves single-pass conversion by substantially reducing equilibrium limitation on the forward reaction. In addition, the high heat capacity of the solvent enables efficient removal of the exothermic heat and facilitates excellent temperature control of the reaction. The strongest solvent that we have identified and tested for the SMP is tetraethylene glycol dimethyl ether (TEGDME) or tetraglyme. TEGDME (molecular weight 222.3kg/kmol) has a normal boiling point of 548 K (275 OC) and possesses excellent thermal stability. Vapor-liquid equilibrium (VLE) experiments on the syngas, CH30H, HzO, TEGDME system under synthesis conditions have shown that the desired partitioning of the species (gases mostly in the vapor phase and CH30H predominating in the liquid) is attainable (Khosla et al., 1991). Simulation of an industrial-scale three-phase reactor with cocurrent upflow of syngas and solvent indicated that up to 95% of the gas feed is converted per pass (Krishnan et al., 1991). The simulation was carried out at physical conditions found in the conventional vapor-phase methanol process. Kinetics for the simulation were estimated by making simple modifications to known vapor-phase kinetic expressions (Berty et all. 1981). According to the simulation (for a stoichiometric synthesis gas feed), almost all of the gas fed was consumed showing that gas recycle is not necessary. In this paper, results of continuous reaction experiments carried out in a Berty internal recycle reactor are presented. Methanol production rates and syngas feed conversion have been measured as a function of four important parameters: reaction temperature, total pressure, syngas flow rate, and the syngas feed stoichiometry given by R [=(HZ - CO,)/(CO = CO,)]. These parameters have been considered at two levels giving rise to 16 data points for a single catalyst pellet size. The effect of pore diffusion on process performance has been analyzed by carrying out
another set of eight experiments with a smaller catalyst pellet size. 2. Experimental Section 2.1. Experimental System. The experimental system used in this study is shown in Figure 1 and may be conveniently divided into three sections: the feed section, the reactor, and the vent section. In the feed section, synthesis gas is prepared by metering individually and then mixing pure gases Hz, CO, CH,, and CO, (supplied by Linde at 99.9% purity) before being compressed into a storage tank up to 1700 psig. The gas mixture is admitted into the reactor by operating a forward pressure regulator (Tescom Corporation), and its flow is metered and controlled by a precision mass flow controller (Brooks 58503). The heart of the experimental system is the Berty internal recycle reactor (Berty, 1974), which is a l-L autoclave. The reactor houses three concentric draft tubes (about 650 cm3dead volume),the innermost of which holds the catalyst. Internal recycle of the reactants is provided by an impeller situated just below the draft tubes. Thermocouples (iron-constantan) provided above and below the catalyst bed serve to monitor the reaction temperature as well as the gradientless condition during runs. Temperature control of the reactor is achieved by use of an electric furnace that has three heating zones. As shown in Figure 1,the solvent is admitted through the side of the reactor. Precise metering of the solvent feed flow is obtained through the use of a positive-displacement pump (Eldex A-6043) that can operate against significant back pressures. The vent section begins with a gas-liquid separator. The gas-liquid foam that exits at the top of the reactor is admitted into a separator (500-cm3Whitey stainless steel bomb) that is inclined upward at a slight angle to the horizontal so as to provide a large surface area for the gas to disengage from the liquid. The gas leaves the separator
Ind. Eng. Chem. Res., Vol. 30, No. 7, 1991 1415 a t the top while the liquid exits through the bottom (see Figure 1). Gas-liquid separation is carried out at reaction temperature and pressure. Subsequent chromatographic analysis of the vapor and liquid phases should then yield kinetic as well as VLE information. A sensitive differential pressure (DP) cell (Validyne Corporation; P305D) located immediately downstream of the separator detecta the d difference in pressure between the liquid and gas streams equivalent to the static head of liquid between the separator and the DP cell outputs a corresponding analog voltage signal. The voltage is read from a digital voltmeter. Liquid level in the separator is controlled by operating a bellows valve (Nupro Company) that is placed downstream of the DP cell on the liquid side. The valve is opened so as to control the signal from the DP cell (and thus the liquid level in the separator) to a predetermined setpoint. The gas stream from the top of the separator is cooled in a double pipe condenser by running cold water, so as to collect the condensibles like CH30H, HzO, and solvent. In all our experiments, we observed that only a small fraction of the total CH30H produced was present in the outlet gas stream. The high selectivity of TEGDME for CH30H ensured that over 90% of the product was present in the discharge liquid stream. A dome loaded backpressure regulator (Grove Regulator Company) that controls the reactor pressure throttles the discharge gas down to atmospheric pressure. The vent gas is then let through a wet test meter (Precision Scientific Company) for flow rate measurement after being analyzed for its composition chromatographically. A thermal conductivity type gas chromatograph (Perkin Elmer Model 8410) equipped with a 10-port gas sampling valve and a 4-port liquid sampling valve (Valco International) was used for compositional analysis of the feed and discharge gas streams as well as the discharge liquid stream. Two packed GC columns, Porapak Q and Tenax, were used in series to resolve all the components of the system. The gases, CH30H, and H 2 0 were resolved by Porapak Q while the heavy solvent was eluted by backflushing the Tenax column. The liquid stream was sampled at reactor temperature and pressure while the gas streams (feed and discharge) were sampled a t 5-10 psig. A mixture of HpHe in the volume ratio 8.691.4 was used as carrier gas for the GC. The experimental procedure was as follows: The reactor pressure was taken up to about 350 psig with syngas. Solvent pumping was started and continued until around 350 cms of the solvent was admitted into the reactor. At this point, the reactor was nearly filled with liquid (the empty internal volume is about 350 cm3). Solvent pumping was stopped. The reactor pressure was then raised to the final operating value and syngas flow was eatablished. The heating of the reactor was started next and the impeller was switched on. The temperature was taken to the final value, and during the heatup period the DP cell signal was carefully monitored. There was a slight rise in the signal due to the solvent overflowing into the separator by thermal expansion. At this point, solvent pumping was resumed and maintained during the entire course of the experiment. The DP cell signal was allowed to rise to about 1.5 V dc (corresponding to about 11 in. of liquid head). The bellows valve (downstream of the DP cell) was then opened just enough to let the liquid flow out so as to maintain the signal at around 1.5 V dc. The flow rate of discharge liquid was conveniently measured by using a buret. Chromatographic analyses of the exit gas stream and the liquid stream were carried out after allowing sufficient time for steady-state conditions to be reached.
Table I. Experimental Plan for Kinetic Study
high
temperature, K pressure, MPa gas feed flow, mol/h feed gas stoichiometryb [R = (H2- C02)/(C0 +
low value 493.00 7.85 0.67 1.69
catalyst size,’ (volume/surface ratio)
0.0143 0.0313
parameter
cod1
value 513.00 10.13 1.34 2.40
a The effect of catalyst size waa observed at only the lower value for R. *The feed compositions in vol % for H2/CO/CH,/C02 were 62.7/26.2/4.2/6.9(R = 1.69) and 70.0/15.0/5.0/10.0(R = 2.4),respectively. e See text for explanation.
2.2. Experimental Plan. Table I shows the experimental plan followed for this study. Operating conditions were chosen to closely match those encountered in commercial vapor-phase methanol synthesis so as to provide a sound basis for comparison of process performance. In CH30H synthesis, which is strongly exothermic and is also accompanied by a decrease in moles (see eq 11, temperature and pressure play a key role not only in dictating reaction equilibrium but also in determining the kinetics. Accordingly, reaction temperatures were fixed at 493 and 513 K while total pressures were chosen to be 7.85 and 10.1 MPa (77.5 and 100 atm, respectively). In the SMP, these variables take on an added significance in that they also dictate the VLE characteristics of the process (Khosla et al., 1991). CH,OH solubility in the solvent is favored at high pressures and low temperatures. Consequently, we can expect a superior process performance a t these conditions. The third important variable is syngas feed flow rate, which affects the single-pass conversions of Hz and CO. Values were fixed at 15 and 30 standard liters per hour (0.67 and 1.34 mol/h, respectively). The effect of reactant composition on the kinetics can be followed by varying the parameter R, which represents the syngas feed stoichiometry. A low value such as 1.7 used in this study roughly corresponds to syngas generated by steam reforming naphtha while a high value of 2.4 is typically obtained as a result of steam reforming natural gas. The initial design set of 16 experiments was executed with 3/16 in. X 3/16 in. catalyst pellets (volume to surface ratio of 0.0313 in.). A supplementary set of eight experiments was carried out with a catalyst batch prepared by breaking each 3/ 16-in.pellet into approximately eight equal portions (volume to surface ratio of 0.0143 in.). The variables for this set were temperature, pressure, and syngas feed flow with R held constant at 1.7. The catalyst employed in this study (Cu/ZnO/Alz03) was prereduced as received, and only a mild reduction procedure was required to remove the surface oxide layer and expose the active copper sites. Reduction was carried out over a period of 4 h with syngas flowing at 15 standard liters per hour at 473 K and a total pressure of 7.85 MPa. Other operating conditions (which were kept constant for all experiments) were catalyst amount (25.7 g), solvent feed rate (0.75 mol/h), and impeller speed (750 rpm). 3. Results and Discussion The results of the kinetics experiments are summarized in Table 11. Experiments 1-8 were carried out with the smaller catalyst particles and experiments 9-24 with 3/ 16 in. X 3/16 in. pellets. Single-pass feed conversions shown in Table I1 (H, and CO) were computed from the number of moles fed and the unreacted gases present in both phases at the reactor outlet. It should be noted that the values in Table I1 reflect contributions to conversion from both reactions.
1416 Ind. Eng. Chem. Res., Vol. 30, No. 7, 1991 Table 11. Kinetic Data for the Solvent Methanol Processo-" production rates, mol/(kg of catalyst h) CHBOH HZ0 3.77 0.78 3.73 0.93 0.93 8.02 5.37 0.66 8.60 0.74 5.84 1.01 8.40 0.89 6.50 0.51 8.56 0.62 2.71 1.08 5.18 0.92 4.35 0.84 6.22 0.92 5.02 0.68 6.58 0.84 5.90 0.48 8.25 0.96 3.71 1.87 3.59 1.74 4.07 1.72 4.02 1.69 5.35 1.12 7.59 1.59 3.85 0.68 6.83 1.31 1.15 6.28 1.63 8.54
conversion, % run
T,K
P, MPa
1
493 493 493 493 493 513 513 513 513 493 493 493 493 513 513 513 513 493 493 493 493 493 493 513 513 513 513
7.85 7.85 7.85 10.13 10.13 7.85 7.85 10.13 10.13 7.85 7.85 10.13 10.13 7.85 7.85 10.13 10.13 7.85 7.85 7.85 7.85 10.13 10.13 7.85 7.85 10.13 10.13
1R 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 17R 18 18R 19 20 21 22 23 24
feed gas flow rate, mol/h 0.67 0.67 1.34 0.67 1.34 0.67 1.34 0.67 1.34 0.67 1.34 0.67 1.34 0.67 1.34 0.67 1.34 0.67 0.67 1.34 1.34 0.67 1.34 0.67 1.34 0.67 1.34
R
co
HZ
1.69 1.69 1.69 1.69 1.69 1.69 1.69 1.69 1.69 1.69 1.69 1.69 1.69 1.69 1.69 1.69 1.69 2.40 2.40 2.40 2.40 2.40 2.40 2.40 2.40 2.40 2.40
78.4 77.2 57.0 73.1 64.3 75.3 67.4 74.1 74.1 44.6 46.6 53.7 46.5 64.4 49.7 72.4 56.0 64.3 63.2 39.7 40.1 71.3 63.5 71.4 58.3 74.2 66.5
66.8 64.3 52.3 70.3 61.6 70.2 63.4 71.5 67.8 60.7 41.1 51.1 42.5 61.4 43.7 68.3 51.5 49.7 50.0 29.0 28.9 64.6 55.2 57.7 51.0 67.8 55.8
"Runs 1-8 with smaller catalyst size (volume/surface = 0.0143 in.); runs 9-24 with 3/16 in. x 3/16 in. pellets (volume/surface = 0.0313 in.). bFor all runs catalyst weight was 25.7 g. CForall runs solvent flow rate was 0.75 mol/h and impeller speed was 750 rpm. d R in run number denotes repeated experiment.
These values were computed on the basis of the measured compositions of all the components at the outlet, not on subtraction assuming 100% material balances. Since the closures on our material balances were usually between &lo%, assuming 100% material balances can be misleading. In experiments 1-8, the conversions obtained ( 5 0 4 % ) are fairly high considering that they correspond to a single equilibrium stage operation. At comparable conditions, typical single-pass conversions in the vaporphase process on the lab scale range from 30 to 50% (Gandhi, 1983; Krishnan, 1987). In an industrial-scale reactor which may be considered to consist of several differential stages, the overall conversion can be expected to be very high. This however is not the case with conventional vapor-phase operation where single-pass conversions are limited to 15-30% due to excessive heat generation and the need to keep temperature excursions to a minimum from both equilibrium and reactor thermal stability standpoints. In the SMP,the presence of the solvent in the reactor has two major advantages: effective product removal and efficient temperature control of the reaction owing to the high heat capacity of the solvent. It is therefore possible to envision very high single-pass conversions to the point of eliminating recycle on the industrial scale. Table I11 shows vapor phase equilibrium conversions based on the syngas feed composition (given by R), reaction temperature, and total pressure employed in this study. A comparison of the actual conversions in Table I1 with the equilibrium values in Table I11 should yield a useful measure of the process performance. It can be seen for some of the runs (5-8,15,17,19-24) that the H2 conversions (or H2and CO conversions) obtained are greater than the corresponding vapor phase equilibrium conversions. In particular, experiment 5 showed greater conversions of Hz and CO than the equilibrium values. This aspect of the process becomes noticeable at the higher temperature
Table 111. Equilibrium CO and H2Conversions at Experimental Conditions conversion, %
temp, K 493.0 493.0 513.0 513.0 493.0 493.0 513.0 513.0
press., MPa 7.85 10.13 7.85 10.13 7.85 10.13 7.85 10.13
R
co
1.69 1.69 1.69 1.69 2.40 2.40 2.40 2.40
82.6 86.8 72.1 78.2 91.7 94.3 82.9 88.1
HI 70.9 74.6 62.2 67.4 46.7 50.0 41.5 45.3
where the equilibrium conversion is adversely affected. The fact that vapor phase equilibrium conversions can actually be exceeded may seem unusual, but in the SMP, the presence of a highly selective solvent that removes CH30H continuously, thereby overcoming chemical equilibrium limitations, makes this possible. It should be pointed out in this connection that Zabor et al. (1960), in their study of the catalytic hydration (by tungstic oxide) of propylene to 2-propanol in the presence of excess water, obtained propylene conversions that were much higher than those predicted from vapor-phase thermodynamic equilibrium calculations. Conversions obtained in our preliminary semibatch experiments (Berty et al., 1990) were as high as 93%, but the syngas feed flow was much lower than those employed in this study. Reaction rates were calculated on the basis of the total quantity of CH30H and H20 found in the liquid and vapor phases. The runs with the smaller catalyst particle size yielded CH30H rates that were, on an average, higher than those with the 3/16-in. tablets by a factor of 1.5. It may be noticed for some runs that the H 2 0 conversion rate appears to be higher for the larger catalyst, contrary to expectations. This anomalous result can be attributed to the relatively small conversions for the second reaction and
Ind. Eng. Chem. Res., Vol. 30, No. 7, 1991 1417 the resultant sensitivity to errors in the material balances. The ratio of the size of the larger to smaller catalyst particle on the other hand, is about 2.19 (0.0313/0.0143). This clearly shows that the process does not operate in the asymptotic region of the Thiele modulus-effectiveness factor plot (Froment and Bischoff, 1979) and hence is limited by pore diffusion. In general, the reaction rates were found to be lower than vapor-phase CH30H rates by a factor of 2.5. This observation is in line with the results obtained from our preliminary semibatch experiments (Berty et al., 1990). In the current study, it was observed that an increase in temperature from 493 to 513 K consistently resulted in an increase in CH30H formation rate. This increase however was not as large as it would have been if pore diffusion limitations were absent. For example, the CH30H rate for experiment 13 (5.02 mol/(kg-h)) is greater than that for experiment 9 (2.71 mol/(kg.h)) by a factor of 1.9. On the other hand, for an intrinsic activation energy of 155 kJ/mol (37 kca/mol) reported for the Cu/ZnO/ A1203 catalyst used in the current work (Berty et al., 1981), the increase in the rate should be roughly a factor of 4. Over the entire set of experiments carried out on this study, total pressure had a near linear effect on CH30H rate. An increase in pressure increases the driving force available for CH30H synthesis. Berty et al. (1981) have defined the driving force (DF) as DF PH - PH,eq = PH - PM:/KPpH& (3) where p represents component partial pressure; subscripts H, C, and M denote H2, CO, q d CH30H; and K is the equilibrium constant for reaction 1. In the S d P , the driving force takes the same form but the partial pressures would have to be replaced by either concentrations of the species in the liquid or the liquid-phase fugacities. We prefer to use fugacities, which describe nonideal solution behavior better than concentrations, which are usually computed from Henry’s constants. The Henry’s law holds for dilute solutions but in the SMP, this approach may not be valid especially for CH30H, which is present to an appreciable degree in the solvent. Both the feed gas flow rate and the syngas feed stoichiometry (R in Table 11) favor the rate of formation of CH30H. Increasing the gas flow rate decreases the residence time in the reactor and thus decreases feed conversions. However, it also gives a lower CH30H partial pressure leading to an enhancement in the driving force (see eq 3). Operation is farther away from equilibrium, and hence r a m are higher at higher flow rates. The effect of increasing R is the same as increasing H2 partial pressure, which serves to increase CH30H production rates. On the basis of a mathematical analysis of the rate data given in Table 11, we propose the following kinetic models for reactions 1 and 2: 19541
RT
The liquid-phase fugacities denoted by f in the above models were computed from the Soave equation of state (Soave, 1972) with temperature, pressure, and liquid-phase mole fractions as input variables. The rate equations were expressed in this way because the liquid-phase fugacities provide the closest available measure of the component activities just outside the catalyst pellet and the Soave
Table IV. Statistical Information on Kinetic Models (4)
(5)
parameter
A1 B1
A2 BZ
value 37.56 19541 21.19 10071
It1
6.07 6.26 5.72 5.91
model R2,%
77 71
equation has been shown to provide an accurate representation of the effects of temperature, pressure and composition for a large number of mixtures (Elliott and Daubert, 1987). Statistical information on the parameters in the kinetic models and the goodness of fit are given in Table IV. Equilibrium constants Kl and K2were estimated by use of literature correlations (Thomas and Portalski, 1958; Hyman, 1968). q1 and q2 are the effectiveness factors that account for the observed mass transfer resistances for the CH30H synthesis reaction and the water-gas-shift reaction. The general form of the effectiveness factor (for a pore-diffusion-limited case) is the one corresponding to a first-order reversible reaction, namely
S,/V, is the particle surface to volume ratio; kv . is the rate constant given by k,, = exp(Ai - Bi/V (Table b;K is the equilibrium constant; subscript i refers to the ith reaction. De is the effective diffusivity given by De = ( € / 7 ) D ~ s
(7)
Here, e is the catalyst internal void fraction (0.5 in our case) and 7 is the tortuosity factor (7.2 for the 3/16-in. pellet used in this study). Dw is the molecular diffusion coefficient of CH30H in solvent (Reid et al., 1987). Defining Dw in this manner gives us a conservative estimate for the effective diffusivity. The general form of the model equations in (4) and (5) was originally outlined in an earlier work on vapor-phase CH30H synthesis (Berty et al., 1981). With these models, we were able to obtain reasonable agreement between predicted and experimental rates reported in this study, once the driving forces were rewritten in terms of fugacities and liquid diffusivities were used instead of gas-phase values. A parity plot for predicted and experimental CH30H rates is given in Figure 2. The percentage average absolute deviation was found to be 15%, which seems reasonable considering the simplicity and compact form of the proposed model and the range of conditions considered. In the current study, exernal mass transfer resistances (gas-liquid and liquid-solid) were kept at a minimum by providing sufficient agitation of the gas and liquid in the reactor. The separation of the gas and liquid foam at reaction temperature and pressure enabled the measurement of VLE K values (vapor mole fraction divided by the liquid mole fraction) for all components in the system. If the reactor was operated under minimum external mass transfer resistance conditions, then the K values measured in the present study should agree with those obtained from phase equilibrium studies (Khosla et al., 1991). This indeed was the case for all the experiments, proving that external mass transfer was rapid and that the kinetic
1418 Ind. Eng. Chem. Res., Vol. 30, No. 7, 1991 f = fugacity, atm
K1,K2= thermodynamic equilibrium constants for reactions
1 and 2 kVi = intrinsic rate constant given by k,, = exp(Ai - B i / T ) R1,R2 = rates of reactions 1 and 2, mol/(m3 s) R = universal gas constant, atm m3/(mol K) T = reaction temperature, K 0
-7
0
/
Greek Symbols = catalyst internal void fraction
t 0
= tortuosity factor v1 = effectiveness factor for reaction 1 v2 = effectiveness factor for reaction 2 T
O
/
/
Figure 2. Parity plot for predicted and experimental methanol yields.
process was the slower step. The catalyst was highly selective to CH,OH formation as can be seen from a comparison wih H20 production rates from Table 11. The ratio of CH30H to H20 rates ranged from 2.5 to 14 with the average value around 7. It was also observed that the catalyst withstood the solvent even at high temperatures without significant loss of its initial activity. 4. Conclusions
Continuous experiments on the Berty reactor for the solvent methanol process have demonstrated the following (a) High single-pass conversions of H2 and CO can be obtained in a single differential stage operation. Conversions, in some cases, are higher than vapor phase equilibrium values especially at higher temperatures. (b) Methanol production rates are lower than vaporphase rates by a factor of 2.0-3.0. Solvent present in the catalyst pores leads to pore diffusion limitations. (c) Within the time frame of this study, it was found that the catalyst maintained ita initial activity. The solvent itself was thermally and chemically stable under synthesis conditions. These results demonstrate several novel features of the solvent methanol process and illustrate a general methodology for improving the efficiencies of equilibrium-limited processes. Based on these results and computer simulations of scaled-up reactors, it is possible to establish some preliminary relationships between this methodology and other alternatives. Our preliminary evaluations indicate that the SMP is best suited to large-scale commercial manufacture of methanol with syngas feed from either natural gas or naphtha feedstock. For other feedstocks, such as coal, the high CO/H2 ratio is expected to make alternative methodologies preferable. Acknowledgment We thank the Ohio Board of Regents and the Union Carbide Corporation for financial and technical support. Nomenclature A, B = parameters in the intrinsic rate constant expression (Table 11)
Subscripts H = hydrogen C = carbon monoxide D = carbon dioxide M = methanol W = water 1 = liquid phase Regietry No. CH30H, 67-56-1; CO, 630-08-0; COz, 124-38-9.
Literature Cited Berty, J. M. Reactor for Vapor Phase Catalytic Studies. Chem. Eng. Prog. 1974, 70 (5), 78. Berty, J. M.; Lee, S.; Sivagnanam, K.; Szeifert, F. Diffusional Kinetics of Catalytic Vapor Phase Reactions With Decreasing Total Number of Moles. Inst. Chem. Eng. Symp. Ser. 1981, 87. Berty, J. M.; Krishnan, C.; Elliott, J. R. Beat the Equilibrium. Chemtech 1990, 20 (lo), 624. Chinchen, G. C.; Denny, P. J.; Jennings, J. R.; Spencer, M. S.; Waugh, K. C. Synthesis of Methanol. Part 1. Catalysts and Kinetics. Appl. Catal. 1988, 36, 1. Elliott, J. R., Jr.; Daubert, T. E. Revised Procedures for Phase Equilibrium Calculations with the Soave Equation of State. Znd. Eng. Chem. Res. 1987,26, 1686. Froment, G. F.; Bischoff, K. B. Chemical Reactor Analysis and Design; Wiley: New York, 1979. Gandhi, R. Kinetics of Low-Pressure Methanol Synthesis. M.S. Thesis, The University of Akron, 1983. Hyman, M. H. Simulated Methane Reformer Reactions. Hydrocarbon Process. 1968,47 (7), 131. Khosla, P.; Krishnan, C.; Elliott, J. R.; Berty, J. M. Binary and Multicomponent Vapor-Liquid Equilibria of Synthesis Gas Components, Methanol and Water With Tetra Ethylene Glycol Dimethyl Ether (Tetraglyme). Chem. Eng. Commun. 1991,102,35. Krishnan, C. Effect of the Minor Variables on the Kinetics of LowPressure Methanol Synthesis. M.S. Thesis, The University of Akron, 1987. Krishnan, C.; Berty, J. M.; Elliott, J. R. Simulation of a Three-phase Reactor for the Solvent Methanol Process. Chem. Ens!. - Commun. 1991, in press. Reid, R. C.; Prausnitz, J. M.; Poling, B. E. The Properties of Gases and Liwids, 4th ed.: McGraw-Hill: New York. 1987. Sherwin, M. B.; Blum, D. B. Methanol Synthesis in a Three-phase Reactor. Prep. Pap.-Am. Chem. SOC.,Diu. Fuel Chem. 1976,20 (3), 146.
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Receiued for reuiew February 4,1991 Accepted February 26, 1991
