Ind. Eng. Chem. Res. 1988,27, 616-624
616
Possibilities for the Development of Large-Capacity Methanol Synthesis Reactors for Synfuel Production Rodney J. Dry* Department of Chemical Engineering, Monash University, Clayton, Victoria 3168, Australia
Future synfuel production schemes may call for methanol synthesis reactors larger than those in use today. Methanol may be produced in a number of quite different reactor types, and for each type some kind of practical scale-up limit exists. This investigation involves exploring the features of seven kinds of reactors, and a kinetic modeling approach is used where appropriate to determine the maximum methanol production for a unit 6 m in diameter. Reactor types discussed include the Lurgi tubular packed bed system, the IC1 quench system, and the Casale mixed flow system. Other variations considered are the conventional fluidized bed, the high-velocity circulating fluidized bed, a tube-cooled radial flow system, and a high-voidage version of the four-stage quench reactor. Of the types discussed, four appear t o be capable of producing a n excess of 5000 ton of methanol per day in a single, 6-m4.d. unit: the Casale system, the tube-cooled radial flow reador, the circulating fluidized bed, and the high-voidage quench system. The incentive for constructing large-scale synthetic fuel plants based on natural gas or coal may be weak at present, but despite this there is considerable interest in the processes involved. One of the principal routes involves methanol as an intermediate, and this step is followed by a high-selectivity process aimed a t producing either gasoline (Mobil MTG) or a gasoline-diesel combination (Mobil MOGD). A second contender is based on Fischer-Tropsch synthesis, and the operation may be adjusted to produce either wax or oil. However, selectivity is limited by the chain-growth mechanism (Dry, 1981), and substantial quantities of unwanted byproducta appear. Other processes which have yet to mature will emerge as contenders in due course, one potential example being the catalytic oxidative coupling of methane to ethane and ethylene (Ito and Lunsford, 1985; Edwards and Tyler, 1986). Maturation is likely to be slow, however, and for this reason it appears that for some time yet the selection will be limited to the Fischer-Tropsch approach and the methanol-intermediate approach. New methanol synthesis technologies are also likely to develop, but here too implementation will normally be slow, and conventional technology is likely to remain in place for some time. Methanol synthesis from a mixture of hydrogen and oxides of carbon is a relatively expensive operation, and this is due in part to the combination of high synthesis pressure (typically 80-100 bar for the so-called "lowpressure" version) and a high recycle ratio. The economics of an overall high-selectivity process scheme would be sensitive to changes in the design of the methanol plant, and in view of the large capacities which may be called for to exploit the economies of scale, it appears that information on single-train capacity limitations for the reactor involved could prove useful. If, for example, a plant that produces 5000 ton/day (ca. 40000 bbl/day) of gasoline via the MTG process were to be considered, this would require the production of ca. 13000 ton/day of methanol. The Motunui (New Zealand) methanol units produce 2200 ton/day each (Kirkpatrick, 1984),and if this reactor type were used, one would require six trains. It might be more attractive to consider just two methanol loops, and in this case each reactor would be required to produce 6500 ton/day. Reactors of this size have yet to be constructed, and the purpose of this discussion is to examine design *Current address: CSIRO Division of Mineral and Process Engineering, P.O. Box 312, Clayton, Victoria 3168, Australia.
0888-5885/88/2627-0616$01.50/0
Table I. Plant Data capacity fresh feed recycle ratio
b o o r t e d by Earl and Islam (1986) 1320 ton/day (pure methanol basis) 166 200 m:/h 5.98
compositions, mole fraction fresh feed reactor effluent
component
co
co2
HZ HZO MeOH CHI NO reactor configuration
bed
0.165 0.082 0.715 0.001
0.000 0.029 0.008
0.0103 0.0111 0.7397 0.0131 0.0403 0.1466 0.0389
four beds, cold shot cooling (as per Figure 2) feedlcatalyst distribution catalyst vol, m3 fraction of combined feed
1 2 3 4
reactor dimensions catalyst data equivalent diameter bulk density bed porosity designation first bed feed temp cold shot temp separator pressure
14.7 20.2 27.6 34.6
0.423 0.138 0.203 0.236
height 7.9 m, i.d. 4.575 m 0.006 m (cylindrical) 1200 kg m-3 0.28 IC1 51-2 (age ca. 8 months) 225 "C 110 "C 8200 kPa
concepts which might allow this to be achieved. Practical considerations demand that a constraint to placed on the diameter of the reactor vessel. It should be such that the shell could be transported by road, in sections if necessary. This is desirable since field fabrication of a large, high-pressure shell is likely to be more problematic than shop fabrication. For road transport a limiting diameter of 6 m is typical, though it is recognized that this constraint is not absolute. For the purpose of this investigation, it will nevertheless be assumed that the vessel diameter is limited to 6 m, and the discussion is aimed at identifying the maximum methanol production rate which might be achieved for a given reactor type of this diameter. Basis of Comparison Earl and Islam (1986) recently published a fairly complete set of operating parameters for a 1320 ton/day IC1 0 1988 American Chemical Society
Ind. Eng. Chem. Res., Vol. 27, No. 4, 1988 617 methanol plant-this information is assumed to have come from the Petralgas facility in New Zealand. These data are repeated in Table I, and as such provide a convenient starting point for a model-based analysis of the synthesis process. Modeling details are given in the Appendix, and the main features are as follows. 1. For the methanol synthesis reaction CO + 2H2 CH,OH (1)
fvsr feed
*
the kinetic expression of Schermuly and Luft (1977) as used by Earl and Islam (1986) was employed: rm = YC#CO(YH$'H,) - y m p m /K1 (2) [ A + B ~ c # ~+oCYH,PH~ + &mPm + ETCO~CO,I~ where A,B,C, D,and E are functions of temperature and are given in the Appendix, K1 is the reaction equilibrium constant, y i is the fugacity coefficient for component i, and rm is the rate at which methanol is produced per unit mass of catalyst. 2. For the water gas shift reaction CO + H20 C02 + H2 (3) the rate expression given by Rase (1977) for a commercial copper-zinc shift catalyst, modified to include fugacity correction terms for high-pressure operation, was used:
*
-rCO = F(YCOf)COYH,#H,O
- YC02PC0,rH$'H$2) /p2 (4)
where F is a function of temperature and is given in the Appendix. K 2 is the equilibrium constant, and P is the total pressure. The above expressions were written into differential material and energy balances for the four-stage IC1 quench system on the assumption of plug flow in the gas phase, and the resulting differential equations were solved by a fourth-order Runge-Kutta procedure. When the model was run initially, the methanol production rate and effluent CO and C 0 2concentrations did not agree with those obtained from the plant. It was clear that some adjustment of the rate expressions was called for, and the approach taken was to premultiply eq 2 and 4 by constants which would bring the model results into line with the plant data. It was found that the shift rate expression needed to be increased by a factor of 7.2 to bring the carbon dioxide concentration in the effluent gas down to the level observed, while the methanol synthesis rate equation required a factor of 2.3 to bring the carbon monoxide concentration and the methanol production rate into line. This approach is defensible in view of the fact that catalytic activities are expected to vary with catalyst formulation, and there is no reason to expect the activity of the IC1 catalyst in the Petralgas reactor to coincide with the activities of the two catalysts on which the rate expressions were developed. Material balance closure with plant data was good (by definition) with these activity corrections in place, but observed and predicted temperatures did not agree to the same extent. This is discussed further in the Appendix. In any event, the overall aim of this exercise was to arrive at kinetic expressions which could be used for synthesis in any reactor configuration. It was assumed that the above expressions in their "tuned" form could be employed as such, and these expressions thus form the basis on which the comparative study of different reactor types is carried out.
Reactor Types If a plant were to be built a t short notice, the choice would probably be between the Lurgi tubular packed bed
VI
U
Figure 1. Lurgi system.
system and the IC1 quench system. These reactor types have been well established commercially, and it would seem appropriate to discuss their merits and drawbacks first. 1. Lurgi Tubular Packed Bed Reactor. The Lurgi design involves placing granular catalyst pellets in (typically) 0.05-m tubes and mounting these vertically in a shell-and-tube configuration. Boiling water is circulated on the outside of the tubes, and reaction heat is removed by heat transfer across the tube wall. According to Supp and Quinkler (1985),the maximum temperature difference between the tube centerline and the boiling water is 10-12 "C, and this implies that energy recovery is near optimum from a thermodynamic point of view. This system is shown schematically in Figure 1. Certain features of the Lurgi system make it attractive from an operating point of view. The high heat-transfer area to catalyst bed volume ratio (ca. 80 m2/m3)and the resulting low temperature drop across the tube wall imply that temperature and hence selectivity control is easy. The catalyst itself is treated gently from a thermal point of view, and this argues well for its ability to avoid rapid deactivation. This system is relatively insensitive to changes in feed gas temperature and features direct bed temperature control via shell pressure. The main drawback of the Lurgi system is its intractability as far as scale-up beyond a certain limit is concerned. The problem stems from the mechanical design of the tube sheet, and its integrity becomes questionable at diameters greater than about 6 m. Maximum unit production capacity is around 900-1250 ton/day of methanol (Zardi, 1982), and this is well short of the single-train capacities envisaged for synfuel plants. This suggests that the Lurgi system, whilst being very attractive in some ways, is essentially limited to low-capacity production plants. (As an aside it is interesting to note that this scale-up limitation also applies to the new Mitsubishi reactor system as described by Makihara et al. (1987).) 2. IC1 Quench Reactor. The IC1 quench system relies on cold shot cooling to hold the bed temperature in the range in which high methanol synthesis rates and selectivities may be obtained (ca. 225-275 "C). This system is shown in Figure 2 along with a typical axial temperature profile, and it is not difficult to see that the control characteristics of this system are decidely less friendly than those of the Lurgi system. Different parts of the catalyst beds are required to operate at different temperatures, and these temperatures depend critically on the feed temperatures to each bed. Small deviations in feed temperature tend to be amplified in the system, and it is likely that catalyst life could be reduced by the almost unavoidable thermal swings which would occur during plant upsets.
618 Ind. Eng. Chem. Res., Vol. 27, No. 4, 1988 Table 11. IC1 Quench Simulation Results capacity 2300 ton/day (pure methanol) 282 000 mn3/h fresh feed 6.0 recycle ratio comDositions, mole fraction component fresh feed reactor effluent recycle
co
0.165 0.082 0.715 0.001 0.000 0.029 0.008
COZ
H2 HZO
0.0102 0.0114 0.7406 0.0127 0.0400 0.1461 0.0390
MeOH CHI N, reactor configuration teryerature
bed told
shc
Figure 2. IC1 quench system.
The control scheme therefore carries a heavy burden of responsibility and requires special attention. Catalyst removal and replacement is straightforward due to the IC1 "lozenge" design for cold shot introduction (Rogerson, 1985). However, the impact of this on overall plant operation is probably not great in view of the 3-5 years between catalyst replacements. The system would probably be amenable to on-line catalyst withdrawal and replacement, though the incentive for this is weak for the same reason. The IC1 system is simpler than the Lurgi system from a constructional point of view, though it does entail two rather thanjust a single product-feed heat exchanger. "he reactor vessel itself is a simple synthesis-pressure shell (with lozenge intern&), and ita limitation in terms of throughput appears to be related to bed preasure drop and hence recycle compressor duty. Gas velocities increase from top to bottom as more gas (cold shot) is introduced, and in the bottom bed of the Petralgas reactor, this velocity is ca 0.42 m/s. If this velocity is maintained in the bottom bed of a 6-m-diameter reactor, the result is a production rate of ca. 2300 ton/day of methanol-full simulation results for this system are given in Table 11. This methanol capacity coincides roughly with that of one Motunui unit but is short of the 3000 ton/day suggested by Zardi (1982) as a limiting capacity. It will be assumed that the 6-mdiameter limit is responsible for this reduction from 3000 to 2300 ton/day and that the latter figure applies under the conditions of this investigation. The IC1 quench reactor, while being far from ideal in some ways, is a popular system in wide use today and is likely to remain so for this very reason. The capacity limitation from a synfuels point of view is quite severe, however, and alternative concepts are likely to be more attractive when larger units are called for. 3. Casale Mixed Flow Reactor. Smith et al. (1984) suggest that a radial-flow packed bed system with external heat transfer leads to suitable designs for reactors producing up to 5000 ton/day of methanol. The term "mixed flow" arises from the method used to seal the top of the catalyst basket-this concept evolved in the ammonia industry and is illustrated in Figure 3. In terms of control, the Casale system appears to be similar to the IC1 quench system, at least in broad terms. A relatively sophisticated and carefidly maintained control system would be required, and catalyst thermal history may be expected to be similar to that in the equivalent IC1 system. Catalyst loading and unloading is more compli-
0.0108 0.0119 0.7791 0.0005 0.0030 0.1537 0.0410
four beds, cold shot cooling feed/catalyst distribution catalyst vol, m3 fraction of combined feed
1 2 3 4
25.0 34.4 47.0 58.9
reactor dimensions catalyst data temp bed 1 bed 2 bed 3 bed 4 cold shot temp reactor pressure max gas velocity
0.423 0.138 0.203 0.236
height 7.9 m, i d . 6.0 m as per Table I; 198 ton inlet 225 "C, outlet 273 "C inlet 233 "C, outlet 282 "C inlet 238 "C, outlet 271 "C inlet 233 "C, outlet 252 "C 110 "C 8400 kPa 0.42 m/s
gas feed
fresP 'eed
w
Kipurg cwde
araduc!
Figure 3. Casale mixed flow system.
cated for the Casale system, since removal and replacement of the baskets is required, and this demands that a vessel end be removed. This is not likely to be a frequent event, however, and need not be regarded as a serious drawback. The principal advantage of the Casale system is the potential ease with which scale-up may be performed. The main constraint is related to gas velocity, with velocities outside the range 0.1-0.8 m/s being inadvisable. The lower limit is set to avoid a transition to laminar flow, since this
Ind. Eng. Chem. Res., Vol. 27, No. 4, 1988 619 Table 111. Simulation Results for Casale System capacity 5000, 10 000 ton/day 435 OOO, 870 000 m:/h fresh feed 450,900 ton catalyst 5.1 (X3), 10.2 m (X3) bed height recycle ratio 6.0, 6.0 three beds, all equal height reactor configuration 0.d. 5.9 m, i.d. 1.85 m 0.19-0.20 m/s (outer radius) gas velocities 0.53-0.65 m/s (inner radius) gas compositions, mole fraction (common to both capacities) component fresh feed reactor effluent recycle
co
0.165 0.082 0.715 0.001 0.000 0.029 0.008
COP HP HPO MeOH CH4
N2 bed temp bed 1 bed 2 bed 3
0.0039 0.0050 0.7438 0.0131 0.0404 0.1529 0.0409
0.00410.0053 0.7829 0.0005 0.0030 0.1611 0.0430
feed 220 "C, outlet 257 " C feed 235 "C, outlet 279 " C feed 245 "C. outlet 258 " C
transition, if allowed, would have a deleterious effect on bulk gas to catalyst surface mass transfer. The upper limit is related to pressure drop and a desire to keep the duty of the recycle compressor within reasonable bounds. This limit is significantly higher than the upper velocity limit for the IC1 system (ca. 0.5 m/s)-this is tolerable because the path length for gas flow is shorter in the radial-flow situation. A full simulation was carried out for a 6-m Casale system, and the simulation conditions and results are given in Table 111. For modeling purposes, it was assumed that the gas was in radial plug flow. The gas velocity constraints were easily satisfied, and with the particular catalyst distribution used, the carbon efficiency, defined as the ratio of carbon which ends up in the product to carbon fed (as CO and COJ, was higher at 99.3% than the 98.3% calculated for the Petralgas system. It is recognized that optimization of feed temperatures and catalyst distribution is possible, but it was decided that this would not be considered here-only the general trends are required. The methanol capacity of this system appears to depend primarily on how tall a reactor can be tolerated. A reactor ca. 16 m high can produce 5000 ton/day, whereas a ca. 32-m unit could synthesize 10000 ton/day of methanol. The taller unit would naturally pose some difficult catalyst replacement problems, and ca. 16 m is probably regarded as a limit by Smith et al. (1984) and Zardi (1982)-both suggest a maximum capacity of 5000 ton/day. The antithesis of this argument could be equally valid, however, and in principle a t least there is no reason why a system could not be developed which would allow catalyst to flow from an upper basket to a lower one when the reactor is brought off-line. Catalyst flow characteristics indicate that the system would be amenable to this, and the result would be a catalyst replacement cycle which would avoid removal of the baskets. Realization of such a system would extend the Casale concept to allow 10000 ton/day in a single train, and it would seem that the risks involved would be relatively minor. Overall, the Casale system offers an enormous jump in single-train capacity over the IC1 system. The risks involved appear negligible for the 5OOO ton/day system, while some basic solids handling development work could provide the key to moving this limit out to 10000 tonJday.
Table IV. Reduced Catalyst Charge Quench Simulation Results capacity 7650 ton/day (pure methanol) 997 000 m /: h fresh feed 6.0 recycle ratio comDositions. mole fraction component fresh feed reactor effluent recycle ~
co
0.165 0.082 0.715 0.001 0.000 0.029 0.008
COP H2 H20 MeOH CH4
N2
0.0419 0.0334 0.7297 0.0110 0.0370 0.1159 0.0311
0.0439 0.0350 0.7639 0.0005 0.0030 0.1213 0.0325
reactor configuration bed
four beds, cold shot cooling feedlcatalyst distribution catalyst vol, m3 fraction of combined feed
1 2 3 4
reactor dimensions catalyst data temp bed 1 bed 2 bed 3 bed 4 cold shot temp loop pressure max gas velocity
82.0 112.8 154.0 193.1
0.423 0.138 0.203 0.236 height 19.5 m, i.d. 6.0 m 6-mm Raschig rings; 276 ton
inlet 232 "C, outlet inlet 233 O C , outlet inlet 233 O C , outlet inlet 231 " C , outlet 118 " C 8400 kPa 1.48 m/s
270 274 267 254
"C "C "C "C
4. Open-Catalyst Quench System. The current IC1 quench design appears to be limited in terms of throughput by axial pressure drop. If this is true, then one way of loosening this constraint would be to replace the 6-mm catalyst cylinders with (say) 6-mm catalyst Raschig rings-this would allow around 3 times more flow for the same pressure drop across the same mass of catalyst. One disadvantage is that bed voidage would rise from 28% (Table I) to ca. 70%, and this implies an increase in bed volume by a factor of 2.4. The reactor would thus have to be taller but could handle significantly more gas at a sharply increased gas-to-catalyst loading ratio. Catalyst activity could be higher due to larger effectiveness factors in this mode, but the potential advantage is ignored. Results obtained here may therefore be regarded as conservative to some extent, though the magnitude of this effect is unlikely to be very great. This operation was simulated along the lines of the IC1 quench system but with a gas-to-catalyst ratio of 2.5 times the value implicit in Table I. The results are given in Table IV, and the main conclusion is that methanol production per unit of gas feed suffers only marginally due to the drop in relative catalyst presence. In particular, methanol production per unit of feed gas is reduced relative to the standard IC1 case by 6-7 % only, and this must be compared with a 60% reduction in catalyst on the same basis. This effect may be explained in terms of a higher carbon oxide content in the recycle gas, and a resulting increase in the chemical driving force for methanol formation in all catalyst beds. A plot of the methanol synthesis rate per unit mass of catalyst vs axial position for the standard IC1 system and for the system discussed here is shown as Figure 4, and it is clear that the catalyst is required to work much harder in the high-voidage case. This mode of operation, whilst allowing the production of ca. 7600 ton/day of methanol in a single, 6-m-i.d. unit, has some potential drawbacks which need to be highlighted. A first point of concern relates to running the
620 Ind. Eng. Chem. Res., Vol. 27, No. 4, 1988 steam BF'd
I
Figure 4. Methanol synthesis rates for the high-voidage and standard IC1 systems.
system with increased partial pressures of the oxides of carbon. I t is possible that methanol selectivity could be detrimentally affected, though this may not be too serious in view of the partial pressures which prevailed in the older high-pressure (typically 300 bar) systems. In any event, a careful evaluation of the strength of this effect would be called for. A second point of concern is that of carbon efficiency. This quantity, defined previously as the ratio of carbon which emerges in the methanol product to carbon feed in the form of CO and C02, is essentially a measure of how much potentially reactive material is lost to the gas purge stream. For operations based on coal or C02-richnatural gas this is not a critical parameter, but for processes based on C02-leannatural gas it is important that carbon losses be minimized. Carbon efficiency for the high-voidage system is 94.3%, and this must be compared with 98.3% for the standard IC1 process and 99.3% for the Casale system. It would appear, therefore, that this type of operation should be considered only for coal-based or C0,-rich natural gas based operations. A final point concerns catalyst life-clearly one would not expect the catalyst to last longer than about 40% of the normal 3-5 years. This is still a healthy 1-2 years, though, and given the IC1 easy-unloading system for rapid turnaround, it appears that an increased frequency of replacement may not upset the overall operation too much. The option of on-line replacement is also open, and this could serve as a backup strategy should the need arise. 5. Tube-Cooled Radial Flow System. According to Zardi (19821, some early design ammonia reactors featured tube-cooled catalyst beds. Tubes containing a cooling medium could in principle be placed in the bed, though the fact that this concept in ammonia synthesis has not emerged as a favored option suggests that it is unattractive for some reason. It would be logical to mount such tubes vertically and to allow them to protrude above the bed in order to accommodate thermal expansion, and if this configuration were combined with an axial flow catalyst bed, it is not difficult to envisage substantial lateral variations in temperature. The annular region in the vicinity of a tube could be overcooled, while the between-tube regions could overheat. Ammonia systems are not prone to temperature-related selectivity problems as are their methanol synthesis counterparts, and it is possible for this reason that tube-cooled methanol reactors have never appeared attractive in axial flow mode. Closer tube spacing and smaller diameter cooling coils would naturally ameliorate the situation, but the cost of such a solution is likely to be high. A somewhat different situation emerges when tube cooling of a radial-flow catalyst bed is considered. System geometry would permit the use of axial-flowtubes as shown in Figure 5, and since the bed aspect ratio is high, the tube count and hence construction cost would not be unrea-
Yl
U
Figure 5. Tube-cooled radial flow system.
sonable. Such a system would have a low tendency toward the formation of hot spots since flow would be predominantly across the tubes, while the external heat-exchange system and the operating characteristics could resemble those of the Lurgi system. In other words, some potential exists for the creation of a system with Casaletype scale-up features and Lurgi-type operating characteristics. The geometry of such a system would be constrained by the low-velocity limit on radial gas flow. For a 6-m-diameter reactor and a fresh feed flow of 870000 m,3/h (gasflow equivalent to the 10000 ton/day case in table 111), the maximum bed height which may be tolerated for a radial velocity greater than or equal to 0.1 m/s is ca. 22 m. This must be compared with the three 10.2-m beds used for the conventional Casale system, and this implies that the gas-to-catalyst ratio would increase by about 39%. This is unlikely to have serious consequences for methanol production or even carbon efficiency, however, since the radial temperature increase would be more gradual than in the tubeless situation. This would be reflected as a favorable thermal effect on synthesis rate, and it is likely that the carbon oxides concentration in the purge stream would be low as a result. In view of the obvious difficulties surrounding the prediction of the radial temperature profile, it was decided that a full simulation would not be performed, though on the basis of previous results it is still possible to examine in-bed heat transfer. Removal of reaction heat from the bed would require a certain thermal driving force for a given tube arrangement. If it is assumed that only an open-tube structure can be tolerated, e.g., 0.05-m tubes on a 0.15-m triangular pitch giving a minimum free span of 0.10 m, then it is expected that the required temperature drop across the tube wall would be substantial. If an overall heat-transfer coefficient of 350 W m-l K-l (estimated) is used, the required thermal driving force is about 140 "C. This implies that reaction heat would be recovered in the cooling medium a t about 120 "C, and from a thermodynamic point of view, this is clearly of little value. The heat-transfer area to catalyst bed volume ratio is 9 m2/m3,and when this is compared with 80 m2/m3inherent in the Lurgi design, this result is not too surprising. The situation is improved somewhat if 0.05-m tubes on a 0.10-m triangular pitch are allowed. The heat-transfer area to bed volume ratio is ca. 22 m2/m3in this case, and the required thermal driving force for heat transfer is reduced to about 60 "C. With this geometry, energy recovery a t ca. 200 "C would be possible, and even though this is still well short of the equivalent temperature in a Lurgi system, it is not unreasonable when compared with
Ind. Eng. Chem. Res., Vol. 27, No. 4, 1988 621 the energy recovery schemes of the IC1 and Casale systems. The main problem associated with a tube-cooled radial flow design would be catalyst replacement. The tubes would have to be integral with the catalyst basket, and one alternative would be to remove the tube-basket combination as a single unit. This would be awkward to say the least, and perhaps a better option would be to develop an unloading procedure which does not involve removing the baskets. The free-flowing nature of the catalyst argues well for this, and as with the loo00 ton/day Casale unit, some basic solids handling development would be required. Conventional Fluidized Bed. Amongst the advantages usually cited for fluidized bed operation are uniform temperature and good heat transfer. It is assumed that the methanol catalyst could- be produced in a form resembling that of an FCC powder (average size ca. 70 Km) and that the fluidization behavior of this powder would also be similar to that of its widely used FCC counterpart. It is quite possible that the reduction in catalyst pellet size from 6 mm to 70 Km could enhance access to active sites in the pores, but this effect will not be taken into account. It will rather be assumed that catalytic activity on a unit mass of catalyst basis is constant, and as with the highvoidage quench system, any prediction made on this basis must be regarded as conservative to some extent. The regime of operation envisaged is that of turbulent fluidization. In this state the bed does not consist of bubbles and a dense phase but rather of a poorly defined mixture of dense and lean regions which split and coalesce rapidly (Yerushalmiand Avidan, 1985). Gas-solids contact is likely to be good under these conditions, i.e., little or no gas is expected to pass through the bed without "seeing" catalyst. The constraining factor on gas throughput is likely to be the ability of the cyclone diplegs to return entrained solids to the bed, and if the cyclones themselves are to be located inside the main reactor vessel, then it would seem reasonable to limit the dipleg return area to (say) 10% of the total bed cross section. Entrainment rates are notoriously difficult to predict, but for the purpose of this discussion, it will be assumed that the Zenz-Weil (1958) correlation yields reasonable values. If the dipleg cross-sectional area is set at 2.83 m2 for a 6-m vessel and a maximum dipleg (or standpipe) downflow velocity of 0.5 m/s is assumed, it is possible to evaluate the gas velocity in the bed which would give rise to the implied solids carryover rate. The correlation suggests a gas velocity of 0.65 m/s, and this will be used as a basis for examining the potential of the conventional fluidized bed system. A second factor which must be accounted for is heattransfer area. This necessarily involves the estimation of a heat-transfer coefficient, and this is by no means straightforward since it is not known what the voidage or solids residence characteristics on the cooling surfaces will be. A crude estimate is therefore likely to be as useful as any other, and it is assumed that at 0.65 m/s in air the film coefficient would be 250 W mW2K-I. Correction for gas thermal conductivity according to the Vreedenberg (1960) correlation leads to a fluid bed film coefficient of 650 W m-2 K-l, and it will be assumed that the bulk of the resistance to heat transfer resides in this film. The system envisaged is known in Figure 6, and simulation was carried out by assuming four perfectly mixed stages in series. This mixing pattern is akin to dispersed plug flow, and the modeling approach is similar to that used by Mobil (Avidan and Edwards, 1986) for a turbulent bed MTG reactor. Isothermal operation a t 270 "C and a bed voidage of 70% were assumed, and the gas-to-catalyst ratio came to ca. 1.9 times that implied in Table I for the
I BFh
filter or :iquid scrub
I
1
-
Figure 6. Conventional fluidized bed.
Table V. Conventional Fluidized Bed Simulation Results capacity 3200 ton/day (pure methanol) fresh feed 400000 m;/h recycle ratio 6.0 compositions, mole fraction component fresh feed reactor effluent recycle co 0.165 0.0092 0.0096 0.0161 CO2 0.082 0.0153 0.715 0.7420 0.7805 H2 0.0005 H20 0.001 0.0127 MeOH 0.000 0.0399 0.0030 0.029 0.1419 0.1493 CH, 0.008 0.0389 0.0410 N2 reactor configuration tube-cooled fluidized bed reactor dimensions 16 m high (10-m bed, 6-mfreeboard, 6-m i.d.) 70-pm powder, 144 ton of particle density catalyst data 1700 kg m-3 temp feed 200 "C bed 270 OC (isothermal) pressure 8400 kPa gas velocity 0.65 m/s in-bed cooling 25MW bed voidage 70 %
standard IC1 operation. The results of this simulation are shown in Table V. The methanol capacity of this system represents a substantial 45% increase over the standard IC1 figure for a 6-m vessel, but at the cost of a taller reactor. The carbon efficiency of the fluid bed process is 98.0%, and this is surprisingly close to the IC1 quench system's 98.3%-it appears that the reduced catalyst inventory is counterbalanced by favorable thermal conditions which lead to rapid synthesis. Premature aging of the catalyst is likely but need not be considered a cause for concern since online replacement would be a logical strategy anyway. The level at which reaction heat may be recovered would depend, as before, on how dense a tube bank could be tolerated. If 10 m long, 0.05-m-0.d. tubes on a 0.15-m triangular pitch were selected, the thermal driving force across the tube wall would be about 17 "C. A denser tube bundle containing 0.05-m tubes on a 0.10-m pitch would reduce this to ca. 7.5 "C, and it is clear that the control and energy recovery characteristics of this system would be similar to those of the Lurgi system. There would be scope for increasing the net amount of energy recovered
622 Ind. Eng. Chem. Res., Vol. 27, No. 4, 1988
I
i
ter c i
I
Figure 7. Circulating fluidized bed.
by raising the duty of the external product-feed exchanger, though to what extent this should be performed will not be ventured into here. The main drawback associated with the fluidized bed concept is the fact that the reactor effluent stream would contain fine solids. The traditional way of handling this, at least in fluidized catalytic cracking and in Sasol's Syntho1 process, is to scrub the dusty gas with a suitable liquid-usually an oil of some kind. This approach might work for the methanol system, but it would be imperative that the scrub occur at as high a temperature as possibleisothermal at 270 "C would be ideal. Any drop in temperature across the scrubber would lead to a reduction in heat transferred to the reactor feed stream, and this would in turn lend to a reduction in energy recovery from the bed itself. An alternative approach would be to use a large, modular porous metal filter system with automatic on-line blowback. This approach may be expensive and does not yet appear to be widely accepted, but indications are that this situation could change before long. In overall sense, the conventional fluidized bed offers attractive Lurgi-type energy recovery and operating features. Methanol capacity is limited to 3200 ton/day, and the main drawback is the requirement for a gas dedusting system. Circulating Fluidized Bed. A circulating fluidized bed is essentially a fluidized bed operated at a velocity too high to allow the use of a conventional above-bed cyclone arrangement. The high carryover would rapidly overwhelm such a solids return system, and a completely separate device for returning solids to the base of the bed is required. The entire catalyst inventory would tend to cycle around the bed-return leg loop, and such designs generally rely on pressure recovery in a standpipe to maintain solids circulation (Shingles and Dry, 1986). Relative to the conventional fluidized bed, circulating beds are characterized by a high gas velocity and a low solids inventory. A gas velocity of 2 m/s would be typical for the system envisaged, and provided a reasonably efficient solids return arrangement were present, a voidage of 85% could probably be maintained in the bed. A methanol reactor of this type would be similar in principle to a Sasol circulating bed Synthol unit (Shingles and Dry, 1986; Dry, 1981),though certain design changes would be required to account for the differences between high-temperature Fischer-Tropsch synthesis and methanol production. The most important of these would probably be in the arrangement of the cooling surfaces-methanol synthesis would require substantially more cooling surface
Table VI. Circulating Fluidized Bed Simulation Results capacity 9000 ton/day (pure methanol) fresh feed 1100000 m,3/ h recycle ratio 6.0 compositions, mole fraction component fresh feed reactor effluent recycle co 0.165 0.0145 0.0512 0.082 0.0322 0.0338 co2 0.7757 HZ 0.715 0.7386 0.001 0.0121 0.0005 HZO MeOH 0.000 0.0390 0.0030 0.1348 CH, 0.029 0.1284 N, 0.008 0.0352 0.0370 reactor circulating fluidized bed configuration reactor 35 m high, 6.0-m i.d. dimensions 70-rm powder, 220 ton on reaction (upflow) side catalyst data temp feed 200 oc bed 270 "C (isothermal) pressure 8400 kPa 2.0 m/s gas velocity in-bed cooling 70 MW bed voidage 85 "70
in order to reduce the required thermal driving force for heat transfer. The design envisaged is shown in Figure 7 . Simulation was carried out on the basis of an isothermal bed a t 270 "C, and as before a stages-in-series model was employed to account for more dispersion and backmixing. A reactor bed voidage of 85% and a height of 35 m were assumed, and in view of the high velocity and aspect ratio, it was felt that 10 mixed stages in series would be appropriate as an axial mixing model. The gas-to-catalyst ratio under these conditions is ca.3.3 times that used in the IC1 quench design, and on this basis a low carbon efficiency and a high concentration of carbon oxides in the recycle gas is expected. The simulation results are given in Table VI, and it appears that this design concept would allow the production of 9000 ton/day of methanol in a single, 6-m-i.d. unit. The carbon efficiency is 95.9%, and on this basis it would appear that the circulating bed concept is more readily applicable to coal-based or C02-rich natural gas based synthesis work. The thermal control characteristics of this system are similar in principle to those of the conventional fluidized bed. If an estimated in-bed heat-transfer coefficient of 400 W m-2 K-' and a tube arrangement involving 30 m long, 0.05-m-0.d.tubes on a 0.15-m triangular pitch are assumed, the required temperature driving force is ca. 26 "C. Denser tube arrangements and increased gas feed temperatures are possible as before, though overall it is clear that Lurpi-type control and energy recovery behavior would prevail. As with the conventional fluidized bed, dedusting of the reactor effluent would be necessary. On-line catalyst removal and replenishment would mean that aging could be controlled, and the success of the Sasol Synthol system suggests that the risks associated with this design are not unreasonable.
Conclusions For the production of more than about 5000 ton/day of methanol in a single-train system, there are at least four design options: 1. The first design is the Casale mixed flow system, either with catalyst unloading via basket removal or with
Ind. Eng. Chem. Res., Vol. 27, No. 4, 1988 623 a more sophisticated solids handling system yet to be developed. Carbon efficiency is high, and the risks associated with this design appear to be low. 2. The second design is an ICI-like system with a sharply increased bed voidage and gas-to-catalyst loading ratio. Potential selectivity problems would need to be investigated and a low carbon efficiency suggests that this system should not be used for COz-lean natural gas-based operations, but the design is otherwise straightforward and able to draw on the experience of the IC1 system. 3. The third design is a single-bed, tube-cooled version of the Casale reactor. Such a system would offer attractive control and energy recovery features, but catalyst replacement represents a significant problem. The future of this concept may depend on whether or not a suitable solids handling system could be developed. 4. The fourth design is a high-velocity circulating fluidized bed unit somewhat similar to the Sasol Synthol reactor. A system for cleaning the dusty reactor effluent would be required and carbon efficiency would be moderate to low, but catalyst replacement would be straightforward and could probably be achieved on-line. The energy recovery and control characteristics of the system would be attractive.
The fugacity coefficient ratio term in eq A2 may be correlated against temperature for a narrow pressure range, and for the conditions encountered in this study the following expression was used:
Nomenclature A-F = given functions of temperature F' = molar flow f O = standard-state fugacity, kPa K,, Kz = reaction equilibrium constants, eq 1 and 2 K, = methanol reaction equilibrium constant, eq A2 K, = water gas shift equilibrium constant, eq A7 P,= partial pressure of component i, kPa P = total pressure, kPa ri = reaction rate, mmol of i/s/kg of catalyst R = gas constant T = temperature W = catalyst weight Yi= mole fraction of component i
and the expression given by Wade et al. (1978) is K , = exp[13.148 - 5939.5/T - 1.077 In T - 5.44 X 10-4T + 1.125 x 10-'T2 + 49170/T2] (A8)
[-I
= -1.513
+ 0.0128T - 1.84 X 10-5T2 (A4)
where T is in "C. The quantity
could thus be computed for a given temperature and used in eq 2 as required (Earl and Islam, 1986). For the reverse water gas shift reaction COZ + Hz
Appendix: Modeling Details 1. Rate Expressions and Equilibrium Constants. For the methanol synthesis reaction CO + 2Hz e CH3OH (AI) the components involved may be assumed to form an ideal solution, and under these conditions the equilibrium constant may be written:
where y i is the pure-component fugacity coefficient for component i and Yiis the mole fraction of component i. P is the total pressure, and f O represents standard-state fugacity. Wade et al. (1978) give the following expression for this quantity:
K,/(f
o ) z = 9.74 X exp[21.225 + 9143.6/T 7.492 In T + 4.076 X 10-3T - 7.161 X 10-*T2] (A3)
where K,/(fo)2 is in kPa-2 and T is the temperature in K.
+ Hz0
(A6)
the equilibrium constant may be written as
where T is in K. The fugacity coefficient ratio, for the narrow pressure range of interest, was correlated against temperature as follows: c
-
[EJ
= 0.62 - 0.00731T + 2.5
X
10-5T2
(A9)
where T is in OC. The equilibrium constant K z in eq 4 could thus be computed for a given temperature: r
Greek Symbols y i = fugacity coefficient for component i Pb = bulk density of catalyst bed, kg m-3 Subscripts m = methanol i = component i = methanol, CO, COz, H2, HzO, CH,, Nz
CO
Equations 2 and 4 in the main text require a knowledge of the fugacity coefficients of the reacting species. The following values, assumed constant over the range of synthesis conditions of interest, were used 7, = 0.85, yco = 1.00, YH* = 1.00, ycoZ= 0.95, and Y H ~ O= 0.80. Equations 2 and 4 also require, between them, temperature-dependent factors A-F. Factors A-E were used in the form
KO exp [-Eo/R T 1
(All)
and the values of KOand Eo for each were taken from Earl and Islam (1986):
A:
KO= 6.63 X
lo1,
Eo = 128.3 X lo3 J mol-l
B:
K , = 2.28 x 10-3
E , = -39.4 x 103
C:
K , = 2.12 x
E, = -65.0 x 103
D: E:
10-6
KO= 8.14
Eo = 3.90
KO = 2.03 X lo-"
X
lo3
E, = -116.0
X
lo3
and the value of F , taken from the same source, was
F = 5.086
X
exp(12.88 - 1855/T)/pb (A12)
2. Stimulation of the IC1 Quench System. For a differential (horizontal) slice of catalyst of mass dw, the
Ind. Eng. Chem. Res L988,27,624-630
624
following balances may be written: dF',/dw = rm for methanol
(A131
dF'co,/dw = rco2
(A14
d T - m m r m -+ mco2rco2 _
(A15)
dw
c KC,,
where F is the molar flow rate of species i, AHt is the heat of reaction for the formation of species i a t synthesis temperature, and Cp,is the molar heat capacity of species i. These equations may be solved for each bed in turn, with initial conditions a t the top of the first bed:
w=o
(AW
F' = inlet flow
(A17)
T = inlet temperature
(A181
A fourth-order RungeKutta procedure was used as stated previously, though it was found that a simple Euler integration procedure yielded essentially the same results. Recycle convergence was handled manually. An assumed recycle composition was used to initiate the calculation, and once all four beds had been solved, the reactor effluent composition was computed. The removal of methanol and water from this stream was simulated by forcing the mole fractions down to 0.003 and 0.0005, respectively, and the recycle composition was calculated around these constraints. The new recycle composition was compared with that assumed previously, and the entire procedure was repeated as necessary. The simulation results generally agreed closely with the plant data, though in terms of axial interstage temperatures there did appear to be some discrepancy. Earl and Islam (1986) found that higher interstage temperatures resulted from their simulation than were observed on the plant, and they attributed this to catalyst aging in the upper beds. Interstage temperatures obtained in this study agreed reasonably with those computed by Earl and Islam (1986) and were also generally higher than those observed on the plant. It therefore seems unlikely that the dis-
crepancy could be due to some kind of energy balance inconsistency, and it is possible that the observed interstage temperatures are not directly comparable with those computed because of the way in which they were measured. While this is a concern it is not regarded as crucial since the procedure aims a t seeking relative differences between reactor types. The model used is by no means perfect, but the general trends are nevertheless expected to be reasonable. Registry No. CO, 630-08-0; CH,OH, 67-56-1.
Literature Cited Avidan, A.; Edwards, M. Fluidisation; Ostergaard, K., Sorenson, A., Eds.; Eng. Found.: New York, 1986; p 457. Dry, M. E. Catalysis;Anderson, J. R., Boudart, M., Eds.; SpringerVerlag: New York, 1981; Vol. I, Chapter 4. Earl, W. B.; Islam, K. A. Proc. Chemeca 86, Adelaide, 1986, p 129. Edwards, J. H.; Tyler, R. J. Proc. Chemeca 86, Adelaide, 1986, p 188. Ito, T.; Lunsford, J. H. Nature (London) 1985, 314(6013), 721. Kirkpatrick, R. D. Proc. Chemeca 84, Melbourne, 1984, p 233. Makihara, H.; Niwa, K.; Nagai, H.; Morita; K.; Horizoe, H.; Kobayashi, K.; Kuwada, C. Energy Prog. 1987, 7(1),51. Rase, H. F. "Case Studies and Design Data". Chemical Reactor Design for Process Plants; Wiley: New York, 1977; Vol. 2. Rogerson, P. L. Handbook of Synfuels Technology; Meyers, R. A.; Ed.; McGraw-Hill: New York, 1985. Schermuly, 0.;Luft, G. Chem. Zng.-Tech. 1977, 536f 77. Shingles, T.; Dry, R. J. Encyclopedia of Fluid Mechanics; Cheremisinoff, N. P., Ed; Gulf: Houston, 1986; Vol. 4, Chapter 33. Smith, R. E.; Humphreys, G. C.; Griffiths, G. W. Hydrocarbon Process. 1984, 63(5), 95. Supp, E.; Quinkler, R. F. Handbook of Synfuels Technology;Meyers, R. A., Ed.; McGraw-Hill: New York, 1985. Vreedenberg, H. A. Chem. Eng. Sci. 1960, 11, 274. Wade, L. E.; Gengelback, R. B.; Trumbley, J. L.; Hallabauer, W. L. Methanol; Kirk-Othmer Encycl. c f Chem. Technology, 3rd ed.; Wiley: New York, 1978. Yerushalmi, J.; Avidan, A. Fluidization; Davidson, J. F., Clift, R., Harrison, D., Eds.; Academic: London, 1985; Chapter 7. Zardi, U. Hydrocarbon Process. 1982, 61 (8), 129-133. Zenz, F. A.; Weil, N. A. AZChE J. 1958, 4(4), 472-479.
Received for review February 9, 1987 Revised manuscript received July 29, 1987 Accepted November 25, 1987
An Experimental Comparison of Four Control Structures for Two-Point Control of Distillation Kurt V. Waller,* Dan H. Finnerman,+Peter M. Sandelin, Kurt E. Haggblom, and Sten E. Gustafsson Process Control Laboratory, Department of Chemical Engineering, Abo A k a d e m i , 20500 Abo, Finland
An experimental 1Bplate pilot-plant distillation column has been used to study four control strategies for two-point control of distillation. T h e control schemes consist of two single-input/single-output (SISO)loops. T h e schemes are labeled after the manipulators used: (L,V),(D,V), (D/(L+D),V), and (D/(L+D),V/B). T h e best control quality for feed composition disturbances was obtained with the scheme (D/(L+D),V)and by far the worst with (D,V). T h e difference between the three best schemes was small. T h e two-ratio scheme (D/(L+D),V/B)was clearly the most difficult one to implement on the column studied. The control structure chosen for a distillation column may strongly influence the sensitivity of the column t o disturbances such as changes in feed composition or in feed flow rate.
A great deal of the recent discussion in chemical process control has been focused on control system structures, viz., which outputs should be connected to which inputs. 'Present address: Holmen Printing Paper, Hallsta Paper Mill, S-76301 Hallstavik, Sweden.
In distillation control, a number of schemes with different manipulators have been suggested (Rademaker et al., 1975; Waller, 1982,1986). Also a number of suggestions have been made to combine the controlled variables in various ways (Waller, 1986; Waller and Finnerman, 1987). The discussion has largely concerned dual-composition 0 1988 American Chemical Society
