763
Ind. Eng. C h e m . R e s . 1989,28, 763-771 ance upon conditioning, to exploit this technique for commercial applications.
Acknowledgment We thank Dr. D. Ganguly and Dr. M. Chatterjee (Central Glass and Ceramic Research Institute, Calcutta) for their cooperation in carrying out the desorption experiments and Dr. S. Dasgupta (Jadavpur University) and Dr. S. Roy (Indian Association for Cultivation of Science, Calcutta) for the X-ray diffraction studies. We are also grateful to S. Datta (Lecturer, Jadavpur University) for some valuable suggestions.
Nomenclature Jp = total permeation rate, kg/(h m2) Ji = individual permeation rate, g/(s cm2) a = separation factor Cs = g sorbed/cm3 of dry membrane L = membrane thickness, cm Do= diffusivity at zero concentration, cm2/s D = integral diffusivity, cm2/s y = average plasticizing coefficient, cm3/g E , = apparent activation energy, kcal/mol 0 = angle of scattering Subscripts
M = methanol G = ethylene glycol
Literature Cited Ghosh, I.; Sanyal, S. K.; Mukherjea, R. N. Pervaporation of Methanol-Ethylene Glycol with Cellophane Membrane: Some Mechanistic Aspects. Ind. Eng. Chem. Res. 1988, 27, 1895-1900. Hirotsu, T. Graft Polymerized Membranes of Methacrylic Acid by Plasma for Water-Ethanol Permseparation. Znd. Eng. Chem. Res. 1987,26, 1287-1290. Larchet. C.: Brun. J. P.: Guillou. M. SeDaration of Benzene-n-HeDtane 'Mixture by Pervaporation with Elastomeric Membrane: Performance of Membranes. J. Membrane Sci. 1983,15,81-96. Mark, H. F., Gaylord, N. G., Bikales, N. M., Eds. Cellulose. In Encyclopedia of Polymer Science and Technology; Interscience: New York, 1965; Vol. 111, pp 152-169. Michaels, A. S.; Vieth, W.; Hoffman, A. S.; Alcalay, H. A. Structure-Property Relationships for Liquid Transport in Modified Polypropylene Membranes. J . Appl. Polym. Sci. 1969, 13, 577-598. Rautenbach, R.; Albrecht, R. Separation of Organic Binary Mixtures by Pervaporation. J . Membrane Sci. 1980, 7, 203-223. Rogers, C. E.; Fels, M.; Li, N. N. Separation by Permeation through Polymeric Membranes. In Recent Developments in Separation Science; Li, N. N., Ed.; CRC: Cleveland, OH, 1976; Vol. 11. Sikonia, J. G.; McCandless, F. P. Separation of Isomeric Xylenes by Permeation through Modified Plastic Films. J . Membrane Sci. 1978, 4, 229-241. Suzuki, F.; Onozato, K. Pervaporation of CH30H-H20 Mixture by Poly(methy1 L-glutamate) Membrane and Synergetic Effect of their Mixture on Diffusion Rate. J . Appl. Polym. Sci. 1983, 28, 1949-1956.
Received for review August 4, 1988 Accepted December 28, 1988
Registry No. Methanol, 67-56-1; ethylene glycol, 107-21-1.
PROCESS ENGINEERING AND DESIGN Synthesis of Methanol in a Reactor System with Interstage Product Removal K. Roe1 Westerterp,* Michal Kuczynski,+and Charles H. M. Kamphuis Chemical Reaction Engineering Laboratories, Department of Chemical Engineering, University of Twente,
P.O. Box 217, 7500 A E Enschede, T h e Netherlands
The synthesis of methanol has been carried out in a high-pressure miniplant consisting of two packed tubular reactors in series with a high-temperature interstage product removal. In this way, high per-pass conversions can be achieved, even so high that recycle of nonconverted reactants is not necessary anymore. Methanol was absorbed a t reaction temperatures in a countercurrently operated packed bed absorber with tetraethylene glycol dimethyl ether as the solvent, which proved to be efficient and selective. Solvent vapors have no influence on the activity or durability of the copper catalyst used. A closed absorption-desorption loop has been used for the solvent system. Significant energy savings and raw material savings are expected for large-scale applications of this system. The formation of methanol from carbon monoxide and hydrogen CO + 2Hz = CH30H -AH(298 K) = 91 kJ/mol (1) is a strongly exothermic equilibrium reaction. For reasonable reaction rates, the modern low-pressure copper catalysts require operating pressures and temperatures of
* To whom correspondence should be addressed. 'Present address: DSM Research B.V., P.O. Box 18,6160 M D Geleen, T h e Netherlands.
at least 5 MPa and 485 K, respectively. At such process conditions, the attainable conversion is strongly limited by the thermodynamic equilibrium. To maintain a high driving force for the reaction, industrial converters are operated at per-pass conversions far below the equilibrium values. Therefore, recycle ratios as high as 5-10 for the synthesis gas are common in industrial practice. Such operating techniques cause high-energy consumptions in the product separation section and in the reactant recycle system (Westerterp and Kuczynski, 1986). The recycle of the reactants affects the energy consumption unfavorably 0 1989 American Chemical Society
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because (1) the flow through the reactor is increased, increasing the pressure drops over the reactor and the condensors; (2) as the equilibrium is approached, reaction rates are reduced, so that longer catalyst beds are necessary, increasing again the pressure drop and therefore the energy demand; and (3) cooling and heating the nonconverted reactants after each pass of the reactor also increases the energy consumption. Besides the energy consumption, also the investments in the reactor, recompression section, and condensing section are increased due to the fact that a complete conversion cannot be attained. In this view, when looking for possibilities to improve the efficiency of equilibrium processes, first, methods to increase the conversions and decrease the recycle ratios or even to suppress the recycling completely should be taken into consideration. In our previous project (Westerterp and Kuczynski, 1986), we elaborated on a new reactor principle which enables a complete conversion of the reactants in one pass, despite the thermodynamical limitations. The novel reactor reported on is based on the Gas-Solid-Solid Trickle Flow Reactor (GSSTFR) principle: the reaction product is removed from the reaction zone by means of a countercurrent stream of a selective adsorbent which keeps the methanol concentration in the gas phase very low, so that the driving force for the forward reaction remains high. In a miniplant, 100% conversion was easily achieved. A technical-economical evaluation of the GSSTFR system for the methanol synthesis showed great savings of circulation energy, cooling water, and steam, and also the amount of catalyst required can be reduced considerably. As reported previously (Westerterp and Kuczynski, 1986, 1987a,b; Kuczynski et al., 1987a), the high-pressure GSSTFR has been proven t o be completely successful on a miniplant scale. The possible solid-handling problems can be avoided by replacing the solid adsorbent by a selective, liquid absorbent: pumping and flashing of liquid streams are wellmastered operations, in contrast to the respective solidhandling counterparts. However, in a catalytic reactor, where a liquid absorbent flows over the catalyst packing, inevitably the reaction will be hampered by additional mass-transfer resistances over the liquid film on the catalyst surface, especially inside the pores. This will result in an ineffective utilization of the catalyst and accordingly result in an excessive reactor volume. Chem Systems Inc. (1981) proposed to execute the methanol synthesis with a three-phase process, in which an inert liquid is used to fluidize the catalyst and remove the heat of reaction. Overall heat-transfer coefficients in their gas-liquid-solid system may increase perhaps by a factor of 2 , and therefore the required cooling area can decrease considerably in their proposed system. On the other hand, very serious drawbacks of the Chem Systems’ process can be envisaged: (1)The maximum possible solids holdup in the catalyst slurry will be 25-3070 percent at the very most. It is clear that the required reactor volume will be much larger than in the conventional processes, because of the large volume of inert liquid in the reactor. ( 2 ) The additional mass-transfer resistance over the liquid phase will necessarily reduce reaction rates, so that an increased amount of catalyst will be required. In view of these points, the Chem Systems’ process exhibits essential drawbacks for large-scale application; see Graaf (1988). We have developed a system where the advantages resulting from a high per-pass conversion can be achieved
td absorbers b.
t
r
Figure 1. (a) Possible reactor-absorber configurations for catalytic equilibrium reactions consisting of a series of reactors with absorbers in between and (b) a RSIPR system with only one reactor-absorber set: 1, reactor; 2, absorber; 3, recycle compressor; 4, flash vessel; 5, methanol condenser.
by installing packed tubular reactors in series with absorbers in between. Such a system, which we call the Reactor System with Interstage Product Removal (RSIPR), is shown in Figure la. The reactant conversion obtained in a single-pass operation of the RSIPR is first determined by the reaction equilibrium and by the number of the reactor-absorber sets. For example, by use of four properly dimensioned cooled tubular reactor-absorber sections in series with a conversion of 50% per section, an overall conversion of 94% can be attained. The remaining small gas stream can be recycled or used for other purposes. The energy required for the liquid recirculation will be much lower than the energy consumption of the usual gas recycle loop. Technical and economical aspects of such a methanol process have been discussed in another paper (Westerterp et al., 1988). Of course, the lowest possible number of reactor-absorber sets in a RSIPR is one, as defined in Figure Ib. In that case, a recycle compressor is still necessary and energy savings on recompression are absent. Moreover, carbon efficiency does not improve because still a considerable amount of recycle gas has to be purged to keep the level of inert components in the reaction mixture low. The only savings obtained now refer to the methanol recovery: the methanol is only slightly diluted with some synthesis gas, which leads to much higher heat-transfer coefficients-due to the absence of inerts-in the methanol condensers. So to profit from the RSIPR system, at least two sets of a reactor and an absorber are required and preferably four or more, because in the latter case the gas recycle can be deleted. An advantage of the RSIPR proposed here is that product is removed from the reaction mixture without
Ind. Eng. Chem. Res., Vol. 28, No. 6, 1989 765 cooling down and heating up again large streams of gas, because absorption takes place at the reactor inlet temperature. It is clear that a good selectivity of the solvent for methanol is desirable. If the solvent circulation loop is well insulated, the solvent will take up the absorption heat of methanol in the absorber, which then will be released again in the flasher, so in an ideal situation, the solvent loop even can be operated autothermally. Finally in view of the single-pass operation, no inerts. are recycled therefore, reaction rates are not reduced by the increased presence of inerts, as is the case in a system with recycle of nonconverted reactants.
Experimental Program The reasoning as outlined in the introduction has to be proven experimentally. T o demonstrate its feasibility, it is not necessary to construct a full-scale pilot plant and operate it successfully. For demonstration purposes, we devised a restricted program in which we used a highpressure miniplant consisting of two tubular reactors and a countercurrently operated packed column absorber installed in between. Further, the system is equipped with a continuously operable absorption-desorption loop. This configuration allows us to study the following aspects of particular interest: (1)absorption capacity, selectivity, and durability of the methanol absorbent used; (2) possible effect of solvent vapors on the activity and the durability of the methanol synthesis catalyst; (3) the performance of the packed column methanol absorber at high temperatures and pressures; and (4) optimum operating conditions for the methanol desorption section. In a previous study (Kuczynski et al., 1986), the binary vapor-liquid equilibria of methanol with sulfolane, tetraethylene glycol dimethyl ether (TEGDME), and 18crown-6 have been investigated. Sulfolane and TEGDME were found to absorb methanol vapors equally well; 18crown-6 was found to be a less effective methanol solvent. Since the copper catalysts are extremely sensitive to poisoning by sulfur compounds, we decided not to use sulfolane as the methanol solvent in our experimentseven a very slow decomposition of the solvent could then cause a degradation of the catalyst-and therefore we selected TEGDME. Since in our RSIPR methanol process a good selectivity of the solvent is required, we feel that the Cl6-C= paraffin mixtures as used by Chem Systems (1981) are not suitable absorbents for our system, because of the good solubility of CO and H, in these liquids. It was not our aim to optimize the choice of solvent: we will prove the feasibility of the reactor system with interstage product removal, realizing that there may exist better solvents than TEGDME, which we have used. To determine the possible effect of solvent vapors on the activity and durability of the copper catalyst, we operated simultaneously two identical packed bed reactors in series with the absorber installed in between them. The first reactor will always be fed with a solvent-free synthesis gas, whereas the catalyst in the reactor installed downstream of the absorber will be exposed to the solvent vapors which possibly are entrained in the reaction gas. Using this reactor-absorber-reactor system, we will obtain reliable information on the possible degradation of the catalyst in the second bed by executing conversion tests periodically for each reactor separately and at identical conditions. Of course, reliable information on the durability of both the catalyst and the solvent can only be obtained by operating the system a long time. Our experiments have been carried out over a period of 12 months: such a long operating time allows us to draw
reliable conclusions on the durability of both the catalyst and the solvent used. Due to the lack of literature data on the mass-transfer rates a t high pressures and temperatures, it is also an objective of our study to determine experimentally the HTU values for the methanol absorber in our RSIPR. To this end, it is necessary to measure the flow rates and methanol concentrations in all process streams entering or leaving the absorber. The concentration of methanol a t the inlet of the absorber can easily be varied by adjusting the degree of conversion in the first reactor; this can simply be effected by varying the residence time of the reaction gas or the reaction temperature. We will operate the first reactor at conversions of 35-55%, where optimum specific reaction rates are obtained in isothermal large-scale methanol reactors. The total pressure and the temperature of the absorber are very important operating parameters influencing the absorber performance. In our experimental program, a temperature range for the absorber of 175-230 "C and process pressures of 6-8 MPa have been selected. In the experiments dedicated to the determination of the performance of the absorber, it is not necessary to measure the reactant conversion in the second reactor. In order to limit the number of sample takings and gas analyses, in these experiments we will only measure the conversion in the first reactor and the concentrations a t both ends of the absorber column. However, in order to examine the catalyst durability also in these experiments, the second catalyst bed will always be maintained onstream and checked periodically. In the continuous absorption-desorption liquid loop, it is desirable to conduct the liquid flashing a t conditions where methanol is completely recovered from the saturated liquid, so that a methanol-free solvent will be recycled to the absorber. Moreover, if a coabsorption of synthesis gas in the solvent occurs, it will be necessary to recycle the flashed reactants back to the reactor. In order to study various methods of flashing the absorbent, we equipped our miniplant with a closed liquid loop and installed two expansion stages in series. Both one-stage and two-stage desorptions have been examined. The experimental range of the flashing selected is 445-495 K and 0.1-6 MPa for the temperature and the pressure, respectively.
Experimental Installation Our high-pressure miniplant (Figure 2) consisted of two identical packed tubular reactors, each with an inner diameter of 25 X m and a length of 0.5 m. Each reactor contained 360 g of BASF S3-85 copper catalyst, consisting of cylindrical pellets of 5 x 5 mm. Each reactor was equipped with an electrical heater and two thermocouples: one installed axially in the catalyst bed and the other a t the outer side of the reactor tube wall. The thermocouples installed inside the catalyst bed were used to measure the reaction temperature, whereas the thermocouples at the reactor walls were coupled to Eurotherm 810 electronic temperature regulators. Because of the low volumetric gas and liquid flow rates, a small diameter absorber was required. We decided to use a packed bed column in view of its simplicity and flexibility. No reliable design criteria are available in the literature for small-diameter packed columns operating at such extreme conditions: about 480 K temperature and 6-8-MPa pressure in our case. So in model experiments, a proper geometry for the column enabling a stable countercurrent operation, no flooding, good wetting of the packing, and a good radial distribution of the liquid was determined in a series of tests with a glass apparatus at
766 Ind. Eng. Chem. Res., Vol. 28, No. 6, 1989
Figure 2. Experimental setup. (1)Pressure regulators; (2) pressure gauges; (3) electronic mass flow controllers; (4) mixing pipe; (5) sampling points and gas chromatography analysis system; (6) gas preheater; ( 7 ) first reactor; (8) absorber with heating jacket; (9) demister; (10) second reactor; (11) cold trap vessels; (12) back-pressure controller; (13)wet gas meter; (14) sampling points for refractometric liquid analysis; (15) buffer vessel for the solvent; (16) liquid level indicators; (17) electronic level detector; (18) expansion valves; (19) level adjusting valve; (20) expansion vessels; (21) liquid pump; (22) liquid preheater; (23) thermostatic bath. The location of thermocouples is described in the text. Table I. Volumetric Compositions of the Feed Gases, Manufacturers Specification (Hoek Loos, Amsterdam) comDonent
C3
>99.570
99.90/0
