Ind. Eng. Chem. Process Des. Dev., Vol. 18, No. 3,
1979
527
Res. Dev., 13, 75 (1974).
Field, Calif., Contract Nas 2-4184, Jan 23, 1968. Sweeley, C. C., Bentley, R., Makita, M., Wells, W. W., J . Am. Chem. SOC., 85, 2497 (1963). Trigerman, Sh., Biron, E., Weiss, A. H., React. Kinet. Catal. Lett.,8 (3), 269 (1977). Venuto, P. B., Landis, P. S., Adv. Catal., 18, 340 (1968). Walker, W. E., Bryant, D.R., Brown, E. S., Jr., U.S. Patent 3 952 039 (Apr 20, 1976). Weiss, A. H.,Lapierre, R. B., Shapira, J., J . Catal., 18, 332 (1970). Wisniak, J., Hershkowitz, M., Leibowitz, R., Stein, S., Ind. Eng. Chern. Prod.
Received for review July 20, 1978 Accepted January 8, 1979
This paper was presented both at the Fifth Soviet-American ~~~~~i~~ on catalysis,~ ~USSR, k M~~ ~ 18,, 1978, at the 71st National AIChE Meeting, Session 44, Miami Beach, Fla., Nov 14, 1978.
Process Aging Studies in the Conversion of Methanol to Gasoline in a Fixed Bed Reactor Serge1 Yurchak,' Sterling E. Volts, and John P. Warner Mobil Research and Development Corporation, Paulsboro, New Jersey 08066
Catalyst aging studies of the methanol-to-gasoline process were conducted in an adiabatic fixed bed unit. An aging test of over 200 days on stream was achieved during which 8000 Ib of methanol/lb of conversion catalyst was processed. Catalyst activii was still satisfactory at the end of the aging test. Some changes in product selectivities were observed during individual cycles and from cycle to cycle. The properties of the methanol-derivedgasoline are generally comparable to those of commercially marketed gasolines. Combination of this process with the commercially available coal-to-methanol technology provides an alternate route for the conversion of coal to high octane gasoline.
Introduction Coal is expected to become a more important source of gaseous and liquid fuels during the next several decades. Coal gasification technology is used commercially, and improved gasifiers are under development. Several coal liquefaction processes have been developed, and large pilot plants and demonstration plants are being built and operated. Methanol can be obtained from coal (via synthesis gas) with commercially available technology. Studies have shown that methanol can be used directly as an automotive fuel or blended with petroleum-derivedgasoline. However, there are some serious problems associated with these uses. A novel process is being developed for the conversion of methanol to high octane gasoline (Meisel et al., 1976, 1977; Wise and Silvestri, 1976; Daviduk et al., 1976; Yurchak et al., 1977; Chang et al., 1978; Liederman et al., 1978). Combination of this process with the commercial technology for the production of methanol from coal provides another route to obtain gasoline from coal. This methanol-to-gasoline process uses a new type of zeolite catalyst. The conversion of methanol to hydrocarbons and water is virtually stoichiometric. Only small quantities of CO, COz, coke, and H2are formed as byproducts. The yield of gasoline is typically greater than 75 wt % of the total hydrocarbons, and additional gasoline can be obtained by alkylating propene and butenes with isobutane. The ultimate gasoline yield is about 90 wt %. Small amounts of LPG and high Btu fuel gas are the other hydrocarbon products. In contrast to the Fischer-Tropsch process, no significant amounts of oxygenated products are formed, except water. The conversion of methanol to gasoline is highly exothermic, and the heat of reaction is 650-750 Btullb of methanol depending on the particular product distribution. 0019-7882/79/1118-0527$01.00/0
In the fixed bed process, two reactors are used. Methanol is partially dehydrated to an equilibrium mixture of methanol, dimethyl ether, and water in the first reactor. The methanol and dimethyl ether are converted to hydrocarbons and water over a zeolite catalyst in the second reactor. Light gases are recycled to the second reactor to reduce the adiabatic temperature rise to 100-200 O F under typical operating conditions. About 20 and 80% of the total heat of reaction is released in the first and second reactors, respectively. Some further details of the fixed bed process were presented in the earlier publications. As part of the process studies, long-term aging tests were made to establish catalyst stability and regenerability. The results of an aging test in the fixed bed unit are described in this paper. Some effects of coke formation and steam on catalyst performance and changes in product distribution during individual cycles and from cycle to cycle are presented. Experimental Section The experimental studies of the conversion of methanol to gasoline were conducted in an adiabatic fixed bed unit which is shown schematically in Figure 1. It consisted of two fixed bed reactors in series. The first reactor had an inside diameter of 5/s in. and contained an axial thermowell along its entire length. It had a capacity of about 60 cm3 of catalyst. The products from the first reactor were mixed with recycle gas, and the mixture was passed over the conversion catalyst in the second reactor, where the formation of hydrocarbons occurred. The conversion reactor had an inside diameter of 1.30 in. and contained a thermowell along its axis for the entire length. A catalyst bed length of 9.25 in. was generally used, giving an L I D ratio of 7.2. To minimize entrance effects and to establish the inlet temperature, the catalyst was preceded by a bed of quartz 0
1979 American
Chemical Society
528
Ind. Eng. Chem. Process Des. Dev., Vol. 18, No. 3, 1979
Table I. Operating Conditions for Long-Term Aging Test in Fixed Bed Unit
Recycle Cas Methanol
1
Preheater
Preheatar
charge
1,3 ' I D Reactor
518" I O React 60 cc 01 Catai
0 cc 01 Catalyst
High Pressure Separator Dehydration Reactor
Conversion Reactor
Liauid
Figure 1. Schematic of fixed bed pilot plant.
chips (2 in. minimum length). Since the reactor was surrounded by 4 in. of insulation, the operation was essentially adiabatic. Axial temperature profiles in both reactors were monitored every 15 min by computer-controlled, motor-driven thermocouple probes. The products from the conversion reactor were cooled and flashed in a high-pressure separator. Aqueous and organic liquid phases were withdrawn continuously from the separator and depressured to atmospheric pressure. The liquid phases were collected in a product receiver, and the heavy gas product was mixed with unit purge gas. The gaseous material from the separator was heated above the dew point, demisted, and sent to heated compressors for recirculation to the conversion reactor. All lines handling recycle gas were heated to about 175 O F . The recycle rate was measured by timing the displacement of a known volume of heated ethylene glycol at unit pressure. Only the catalyst in the conversion reactor required regeneration. After the unit was purged, the dehydration reactor was sealed off during regeneration of the conversion catalyst. Nitrogen was circulated over the conversion catalyst, and the temperature of the catalyst bed was lined out a t 650 OF at a pressure of 300 psig. Air was admitted, and during the major burning period, the outlet oxygen concentration was kept at less than 1 vol %. The inlet temperature was continually raised (25 OF/h maximum) to 900 O F . After the burning ceased, the outlet oxygen concentration was allowed to rise to about 20 vol %. After an hour a t 900 OF, the air was shut off. The reactor was then cooled to the desired temperature, and the unit was purged of residual oxygen. Gas chromatography was utilized to analyze charge and product streams. Six streams in the fixed bed process were analyzed to obtain detailed yield compositions and mass balances. They included charge stock, effluent from the dehydration reactor, recycle gas, combined off gases, hydrocarbon product, and water product. Further details of the analytical equipment and procedures have been given elsewhere (Stockinger, 1977). The chromatographic data and most other process information were transmitted directly to an APL computerized data base. Material balances, product yields and compositions, process parameters, and other items could be readily studied on conveniently located APL terminals. Pure methanol was purchased from several commercial sources. The charge stock used in the aging studies was a blend of methanol (83 wt 70)and water (17 wt %); this composition simulated the product from the high-pressure separator of a commercial methanol plant. Various dehydration catalysts can be used in the first reactor, and the performance of a typical catalyst has been described previously (Chang et al., 1978). The catalyst used in the conversion reactor is a member of a new class of shape selective zeolites-the ZSM-5 class. The prop-
dehydration catalyst (reactor 1) conversion catalyst (reactor 2 ) inlet temperature (reactor 1) outlet temperature (reactor 1) inlet temperature (reactor 2 ) outlet temperature (reactor 2 ) high pressure separator temperature pressure WHSV (MeOH/catalyst, reactor 2 ) recycle ratio (mol of recycle/mol of charge)
synthetic crude methanol FVH 292-3
EB 415-2-2 600 " F -770 " F 650-680 F -760°F 125 " F 300 psig 1.6 9
erties of these zeolite catalysts have been summarized elsewhere (Meisel et al., 1977; Chang and Silvestri, 1977; Chang et al., 1978).
Results and Discussions A. Process Variable Effects. Exploratory studies in fixed bed microreactors established the effects of major process variables on conversion, product selectivity, and catalyst stability and regenerability. Single catalyst beds were used-both dehydration and conversion to hydrocarbons occurred over the same catalyst. Early process studies established the desirability of using two fixed bed reactors (separate dehydration and conversion catalysts), including some recycle to the second reactor. The initial work in the fixed bed bench-scale unit was, therefore, devoted to defining the effects of selected process variables on the performance of the dehydration and conversion catalysts. The results of the process variable study were presented elsewhere (Chang et al., 1978). The dehydration catalyst can be operated for long periods without any significant loss in activity. The concentrations of methanol, dimethyl ether, and water in the effluent do not deviate from thermodynamic equilibrium values. The effects of space velocity, pressure, temperature, and recycle ratio on the performance of the conversion catalyst were determined. Ranges of conditions which gave cycle lengths of several weeks were defined. The results from the process variable study provided a basis to select operating conditions for the aging tests. Short aging runs defined the sensitivity of the conversion catalyst stability to critical parameters such as temperature. The specific operating conditions for the successful long-term aging test which is described in the next section were chosen from the combined results of these two studies. B. Long-Term Aging Test in Fixed Bed Reactor. The process operating conditions for the long-term aging test are listed in Table I. As will be discussed later, they led to a successful aging test of the conversion catalyst for more than nine cycles, representing 208 days on stream and processing more than 8000 lb of methanol/lb of catalyst. Catalyst cycle length averaged more than 20 days. 1. Dehydration Catalyst. The performance of the dehydration catalyst throughout the 208-day aging test is shown in Figure 2, where the outlet concentrations of methanol, dimethyl ether, and water are plotted as a function of time on stream. The variation in the reactor outlet temperature is also included. As is evident, the catalyst has converted the synthetic crude methanol charge to a product with an essentially constant composition. Some of the variation in composition is due to changes in outlet temperature. The equilibrium composition at 770 OF of an original charge containing 83 w t % methanol and
Ind. Eng. Chem. Process Des. Dev., Vol. 18, No. 3, 1979
529
Table 11. Cycle Lengths of Long-Term Aging Test in Fixed Bed Unit cycle length to MeOH breakthrougha
a
run no.
cycle
MeOH proc., lb ofMeOH/ Ib of cat. 2
212-48 21 2-49 212-50 212-51 212-52 212-53 212-54 212-55 212-56
1 2 3 4 5 6 7 8 9
501 982 643 695 607 963 810 918 920
total cycle length
cumulative time on stream
time, days
MeOH proc., lb of MeOH/ lb of cat. 2
time, days
MeOH proc., lb ofMeOH/ lb of cat. 2
13.1 24.9 16.9 18.0 15.8 24.9 21.1 23.9 24.2
595 1105 795 808 7 08 1046 895 1048 1006
15.5 28.0 20.8 21.0 18.5 27.0 23.3 27.3 26.2
595 1700 2495 3303 4011 5057 5952 7 000 8006
time, days 15.5 43.5 64.3 85.3 103.8 130.8 154.1 181.4 207.6
Methanol breakthrough taken as first material balance during which MeOH was detected in product water phase
-
K
Outlet lemoerature
775'h 750
no
000
Bo o o o o D e
~~
I
I
.00900
0000
I
",B I
i
A A
A
0
-
0 1
.
02
03
04
05
06
07
0 8
09
10
Fractional Bed Length
10
I 50
I 1W
I
I
150
2w
250
Time On Stream I D a y r I
Figure 2. Performance of dehydration catalyst during long-term aging test.
17 wt % water is 19.6 wt % methanol, 45.6 wt 70dimethyl ether, and 34.8 wt '70 water. This catalyst was not regenerated during the life test. 2. Conversion Catalyst. The long-term aging test of the conversion catalyst was arbitrarily terminated after nine successful cycles representing 208 days on stream. A total of 8006 lb of methanol/lb of catalyst was processed. The lengths (to methanol breakthrough) of cycles 6 through 9 appeared to stabilize at greater than 21 days, resulting in the processing of more than 800 lb of methanol/lb of catalyst per cycle. This catalyst was also used in several cycles after the long-term aging test. Examination of the detailed cycle length data given in Table I1 shows some rather interesting behavior. The shortest cycle was the first. Methanol breakthrough occurred after 13 days on stream. The length (to methanol breakthrough) of the second cycle was almost twice that of the first. This increase has been observed in previous tests in both fixed and fluid bed reactors, although the magnitude seems to depend on process conditions. The lengths of the third through fifth cycles declined with respect to the second. The sixth cycle, however, was unexpectedly long (about 9 days longer than the fifth). Experimental difficulties were encountered during the sixth cycle. The total pressure declined from 300 to 225 psig. Thus, the hydrocarbon partial pressure was about 20% lower in the sixth cycle than in the fifth. Since coke formation is related to hydrocarbon partial pressure, the rate of coke formation was probably lower in the sixth cycle. The lengths of the seventh, eighth, and ninth cycles were similar to the sixth cycle. No operational difficulties
Figure 3. Band aging of methanol-to-gasolinecatalyst in first cycle of long-term aging test.
were encountered during these cycles, yet the reactor pressure during each cycle followed that of the sixth one very closely. The reactor pressure decline can thus be attributed to decreasing yields of light gas. The manner in which product yields and selectivities change from cycle to cycle and within a cycle depends upon catalyst activity. In a fixed bed, adiabatic reactor, the conversion of methanol to hydrocarbons is complete until the catalyst deactivates (by coke formation) to an activity such that only partial conversion of methanol is achieved. At this point, unconverted methanol (Le., 0.1 to 0.5 wt %) appears in the product water phase, and the cycle is terminated. A noteworthy feature of this process is that in each cycle the catalyst ages in a band due to coking. At the aging test process conditions, only 35-40'70 of the catalyst in the reactor is used to complete the conversion of methanol to hydrocarbons. This quantity was estimated from axial temperature profiles. As the catalyst ages during a cycle, the reaction front moves toward the outlet of the reactor. In an adiabatic reactor, this is evident from the movement of the temperature profile through the reactor. An example of this movement is shown in Figure 3 for the first cycle. For clarity, the normalized rather than the actual temperature rise is plotted. Methanol breakthrough occurred at 314 h on stream. The main reaction zone is considered to be that portion of the reactor where temperature gradients exist. The significance of this band aging is that the time, temperature, and water pressure history of each segment of the catalyst bed is different. Since the conversion catalyst is deactivated by steam and steaming depends on these three variables, an activity gradient is established
Ind. Eng. Chem. Process Des. Dev., Vol. 18, No. 3, 1979
530
Table 111. Operating Conditions and Products for Selected Cycles in Long-Term Aging Test cycle 1 conditions time on stream in cycle, h cumulative charge, wt of MeOH/yt conv. cat. dehydration reactor inlet temp., F dehydration reactor outlet temp., F conversion reactor inlet temp., F conversion reactor outlet temp., F high pressure separator temp., O F WHSV, MeOH over conv. cat. recycle ratio, mol/mol of charge pressure, psig products, wt % of charge hydrocarbons H*0
H2
co CO,
methanol dimethyl ether material balance, wt % 9 RVP gasolinea research octane ( R t 0) leaded octane (R t 3 ) specific gravity molecular weight yield, wt % of hydrocarbons a
73 117 601 779 648 774 128 1.6 9.2 300
cycle 6
314 501 604 776 652 777 129 1.6 8.9 300
94 4162 583 773 649 766 125 1.6 9.7 300
cycle 9
333 4545 586 771 652 774 135 1.6 8.8 24 5
77 7123 576 771 650 778 126 1.6 9.2 283
35.5 64.1 0.01 0.04 0.4 0.0 0.0 98.4
35.3 64.2 0.01 0.03 0.3 0.2 0.0 97.8
36.2 63.8 0.0 0.01 0.0 0.0 0.0 98.2
35.2 64.7 0.0 0.01 0.05 0.0 0.0 98.6
95.1 100.8 0.756 95 73.3
93.3 99.4 0.734 95 86.2
93.4 99.9 0.736 95 82.6
93.2 99.3 0.728 95 92.0
315 7504 587 774 651 779 124 1.6 8.3 250
36.9 63.0 0.0 0.01 0.08 0.0 0.0 100.2
35.9 64.0 0.0 0.0 0.07 0.0 0.0 99.0
93.2 99.9 0.731 94 85.5
93.0 99.1 0.723 94 89.3
Gasoline includes alkylate; properties are calculated from compositions.
in the catalyst bed. Catalyst located near the reactor inlet will be more active than catalyst near the outlet. In addition to this permanent activity gradient, a temporary activity gradient due to coking will also develop in a given cycle. As the main reaction zone moves towards the reactor outlet, the conversion is occurring over less active catalyst, and product selectivities should change in a manner to reflect this movement backward along the reaction path. 3. Product Selectivity Changes. Operating conditions, product yields, and some gasoline properties are detailed in Table I11 for portions of the first, sixth, and ninth cycles. The on-stream times are for a particular cycle only, whereas the amounts of methanol processed are for the aging test as a whole. The conversion of methanol to hydrocarbons is complete and virtually stoichiometric. Only small quantities of CO, COz, and Hzare formed. Material balances are very good. The 9 RVP (Reid vapor pressure) gasoline has an unleaded research octane number of 93-95 (calculated from the raw liquid octane number and the hydrocarbon composition of the finished gasoline). After the first cycle the octane number varied only slightly. Detailed hydrocarbon product distributions for these three cycles are presented in Table IV. Only small quantities of methane, ethane, ethene, and propene are formed. The low yields of ethene and propene are due to the recycle mode of operation. Isobutane yields are high. This is important since virtually all of the propene and butenes can be upgraded to gasoline range material by alkylation. Within a cycle and from cycle to cycle, the following trends are apparent: propane, decreasing; C6+ nonaromatic, increasing; olefins, increasing; aromatics, decreasing; C,+ gasoline, increasing; 9RVP gasoline, increasing. These trends are consistent with the discussion of catalyst activity gradients in the conversion reactor given earlier and the following reaction path: (methanol, dimethyl ether)
-
-
olefins aromatics and paraffins
Since insignificant amounts of H2 are produced, the aromatics are believed to be formed from olefins via hydrogen transfer reactions, which also produce paraffins.
50
0
5w
IOW
0
5oc1
1000
0
5W
lc00
0
XU
lco3
ISM
Methanoi Processed Per Cycle ILb MethdnOl Lb Catalyst1
Figure 4. Gasoline yields in long-term aging test.
Despite the declining yields of high octane aromatics, the clear research octane number of the 9 RVP gasoline remains constant, The production of high octane olefins and isoparaffins is enhanced as the catalyst ages. This offsets the expected decline in octane number due to lower yields of aromatics. More detailed data showing the changes in C,+ and 9 RVP gasoline yields in a cycle and from cycle to cycle are presented in Figure 4 (for cycles 1, 3, 6, and 9). The abscissa is the methanol processed within a given cycle. The trend to higher start-of-cycle gasoline yields is clear. End-of-cycle 9 RVP gasoline yields are in excess of 90 wt 70 of hydrocarbons. Changes in aromatics and olefins yields are shown in Figure 5 for these cycles. The correspondence between aromatics and olefins is evident (i.e., decreasing startof-cycle aromatics and increasing start-of-cycle olefins yields). These yields also pattern the changes observed within a given cycle and can be interpreted in terms of declining catalyst activity. The detailed aromatics distribution is quite revealing (Table IV). Only small quantities of benzene are produced. Methyl-substituted benzenes are the principal aromatics. Those containing 8 and 9 carbon atoms are favored. Aromatics with more than 12 carbon atoms are not produced in significant quantities. Close inspection of the
Ind. Eng. Chem. Process Des. Dev., Vol. 18, No. 3, 1979 531 Table IV. Hydrocarbon Product Distribution in Long-Term Aging Test cvcle 6
cvcle 1 time on stream in cycle, h cumulative charge, wt of MeOH/wt of conv. cat. hydrocarbon composition, wt % methane ethane ethene propane propene n-butane isobutane butenes total C,n-pentane isopentane pentenes cyclopentane C,+ nonaromatic benzene toluene ethylbenzene p - and m-xylenes o-xylene trimethylbenzenes methylethylbenzenes propylbenzenes 1,2,4,5-tetramethylbenzene 1,2,3,5-tetramethylbenzene 1,2,3,4-tetramethylbenzene other C,, benzenes C,, alkylbenzenes naphthalenes unknowns total C,+ total aromatics
cvcle 9
73 117
314 501
94 4162
333 4545
77 7123
316 7504
2.1 1.3 0.0 11.6 0.1 5.9 8.5 0.5 30.0 2.7 10.1 0.7 0.3 20.0 0.4 4.4 0.9 9.5 2.7 8.4 2.3 0.1 3.6 1.1 0.4 1.5 0.8 0.1 0.0 70.0 36.1
1.8 0.4 0.0 4.6 0.2 2.6 8.5 1.1 19.2 1.3 11.3 2.4 0.2 38.1 0.2 1.7 0.5 6.7 1.7 7.3 2.5 0.1 4.2 0.2 0.1 1.6 0.6 0.1 0.0 80.8 27.5
0.9 0.6 0.0 6.6 0.2 3.7 9.3 1.0 22.3 1.7 11.4 1.9 0.3 31.5 0.2 2.0 0.6 7.4 2.0 8.2 2.8 0.1 4.2 0.4 0.2 1.9 0.8 0.1 0.0 77.7 31.0
0.7 0.2 0.0 3.0 0.2 2.0 6.4 1.1 13.6 1.2 10.1 2.8 0.3 43.6 0.2 1.8 0.6 7.1 1.8 7.3 2.8 0.2 3.6 0.2 0.1 1.9 0.7 0.1 0.0 86.4 28.5
0.9 0.4 0.0 5.3 0.2 3.1 8.9 1.2 20.0 1.6 11.2 2.3 0.3 35.3 0.2 2.1 0.6 7.2 2.0 7.3 2.8 0.2 3.2 0.4 0.1 1.9 0.7 0.2 0.4 80.0 29.4
0.8 0.3 0.0 3.9 0.3 2.3 8.0 1.4 17.0 1.2 10.6 2.9 0.3 41.4 0.2 1.8 0.6 6.2 1.7 6.5 2.6 0.3 3.2 0.1 0.1 1.8 0.7 0.1 0.7 83.0 26.7
m Cycle 3
Cycle 1
Cycle 1
Cycle 9
Cycle 6
1.m
0
m
1wo
0
m
1wo
0
5w
Iwo
0
m
Iwo
0
lKxl
aromatics distribution (particularly cycle 1) shows that, as a cycle progresses, the light aromatics (C6-C8) decrease and the C9and Clo aromatics increase. CI1 aromatics show no change. A convenient way to examine this behavior is in terms of side chain carbon/ring carbon ratio. These ratios for cycles 1,3,6, and 9 are plotted in Figure 6. For cycles 1 and 3 (actually for the first five cycles) the side chain/ring ratio increases within a cycle. This increase means that the aromatics are more heavily alkylated. After the fifth cycle, discernible changes in this ratio were not evident. The low carbon-number aromatics initially formed are probably alkylated with methanol. Thus, aromatics formed in the presence of unconverted methanol should have a higher side chainlring ratio than those formed after methanol conversion is complete. The increasing side chain/ring ratio of the aromatics as a cycle progresses is thus consistent with the conversion being carried out over progressively lower activity catalysts. A comparison of experimental aromatics distribution with equilibrium values is shown in Table V for the second
Mo
IwO
0
SCo
1wO
0
,
r
F
y
Cycle 1
Cycle 6
5w
1wO
IMoIXa
503
0
hlethanol Processed Per cycle l i b Methanol ltb Catalystl
Methanol Processed Per Cycle ILb Methanol Lb Catalystl
Figure 5. Aromatics and olefins in long-term aging test.
-
Cycle 1
Figure 6. Aromatic side chain to ring ratio in long-term aging test. Table V. Comparison of Experimental Aromatics Distribution from Second Cycle with Equilibrium Values time on stream in cycle, h cumulative MeOH processed, wt of MeOH/wt of cat. maximum temperature, F chainlring ratio, wt/wt
11 613
645 1654
745 0.437 equil exptl
779 0.492 equil exptl
2.1 14.5 33.9 37.4 11.4 0.7
1.4 14.1 41.5 27.4 13.9 1.8
0.9 8.6 27.6 42.9 18.4 1.6
1.1 6.0 35.0 35.5 20.5 1.9
33.0 50.4 16.6
67.6 24.0 8.4
32.8 50.3 16.9
97.3 1.1 1.6
8.1 66.3 25.6
2.2 90.9 6.9
8.2 66.4 25.4
0.3 99.0 0.7
aromatics distribution. mol %
c, C8 c 9
c,
0
CI 1 tetramethylbenzenes, mol % durene (1,2,4,5-) isodurene (1,2,3,5-) prehnitine (1,2,3,4-) trimethylbenzenes, mol % 1,2,3-trimethylbenzene 1,2,4-trimethylbenzene 1,3,5-trimethylbenzene
532
Ind. Eng. Chem. Process
Des. Dev., Vol. 18, No. 3, 1979
Table VI. Yields of C,
+
on-stream time in cycle, h cumulative MeOH processed, wt o f MeOH/wt of cat.
12
188
454
574
4030
4313
4742
4935
Nonaromatics sixth cycle
nonaromatic yields, w t % of hydrocarbons 6‘ 12.0 15.3 17.1 c, 5.8 9.4 11.9 C8 3.7 7.0 9.8 1.4 2.6 3.5 c, 0.4 0.8 1.3 Go+ total 23.3 35.1 43.6
17.9 11.8 10.6 3.7 1.0 45.0
Table VII. Types of C, Nonaromatics
IW
0
?x
2W
HM
4w
Time On Stream IHOUrsI
Figure 7. Changes in hydrocarbon composition during fifth cycle of long-term aging test.
sixth cvcle on-stream time in cycle, h cumulative MeOH processed, wt of MeOH/wt of cat.
12
188
454
574
4030
4313
4742
4935
yields, wt % of hydrocarbons C,-C, n-paraffins 1.1 C,-C, isoparaffins 14.1 C, -C, olefins 2.6 C,-C, naphthenes 5.1 C l o + PON 0.4 total 23.3
1.2 18.5 6.7 7.8 0.8 35.0
1.1 20.4 13.9 7.0 1.2 43.6
1.2 21.1 14.9 6.8 1.0 45.0
cycle at a very short time on stream and after methanol breakthrough. The equilibrium calculations only included reactions such as disproportionation and transalkylation. As such, the equilibrium composition is independent of hydrocarbon partial pressure, and depends only on temperature and the side chainlring ratio. From the data, it would appear that the experimental aromatics distribution by carbon number is in reasonable agreement with equilibrium. The distribution of C8 and Cg aromatics does show some departure from equilibrium values. While the xylenes are in equilibrium, the trimethyl- and tetramethylbenzenes are not. For the tetramethylbenzenes, the catalyst selectively produces durene, yet the thermodynamically favored isomer is isodurene. Similarly, for the trimethylbenzenes, the catalyst overwhelmingly forms 1,2,4-trimethylbenzene; the other isomers are disproportionately low. The selectivity to durene and 1,2,4-trimethylbenzene is apparently favored by lower catalyst activity. Because of durene’s high freezing point (175 O F ) , its presence in the gasoline can cause difficulties in automotive carburetors. Vehicle dynamometer tests with simulated gasolines containing durene showed that its content should be less than 4-5 wt %. The durene contents of the gasolines from the fixed bed unit appear satisfactory. The durene specification could be relaxed considerably if the gasoline from methanol were blended with petroleumderived gasolines, which contain very little durene. It was mentioned earlier that aromatic compounds containing more than twelve carbon atoms are not produced in significant quantities. This is also true for the nonaromatic compounds, as shown in Table VI. These data show the Clo+ nonaromatics are typically less than 1.5 wt % of hydrocarbons. The process favors the production of low molecular weight materials. As a cycle progresses, increasing amounts of c,+ nonaromatics are produced. In addition, even though (26 nonaromatics predominate, the high carbon-number nonaromatics increase more rapidly.
0
2w
am
6W
800
Tine On Stream IHoursI
Molecular weight of recycle gas during second cycle of long-term aging test.
Changes in composition of the c,+ nonaromatics are shown in Table VI1 for the sixth cycle. The low production of normal paraffins is characteristic of the methanolto-gasoline process. As the cycle progresses, c&9 isoparaffins and olefins continuously increase. c&9 naphthenes, on the other hand, pass through a maximum. Similar behavior was observed for all nine cycles. Start-of-cycle c 6 4 9 isoparaffins were typically in the range of 12-15 wt % of hydrocarbons; for the first cycle they were about 6 wt %. After the first two cycles, start-of-cycle Cs-Cg naphthenes were about 5-6 wt % of hydrocarbons. Initial start-of-cycle naphthenes were slightly under 2 wt % . Maximum c6-c9 naphthenes yield was about 7-8 wt % of hydrocarbons. The gross behavior of the hydrocarbon product in terms of hydrocarbon type is shown in Figure 7 for the fifth cycle as a function of time on stream. The parallel behavior of aromatics and C,-C6 normal paraffins is particularly striking. The data clearly suggest that olefins are precursors to aromatics. The increasing selectivity to isoparaffins, as shown in this figure, is somewhat misleading for this class of compounds, since isobutane and isopentane decreased. Naphthenes are relatively stable, but do pass through a maximum about halfway through the cycle. These data generally support the reaction path given earlier. It must, of course, by remembered that virtually complete methanol conversion was obtained throughout all cycles of the aging test. 4. Recycle Gas. An important process consideration is the recycle gas composition. Since the hydrocarbon product distribution changes throughout a cycle, the recycle gas composition is not constant. Changes in recycle gas composition will cause gas heat capacity changes which, in turn, will affect the magnitude of the temperature rise in the adiabatic conversion reactor. A typical change in the molecular weight of the recycle gas is shown in Figure 8 for the second cycle. The molecular weight decreased from about 28 at the start of cycle to about 21 at the end.
lnd. Eng. Chem. Process Des. Dev., Vol. 18, No. 3, 1979 533 Table VIII. Equilibrium of Five-Carbon Olefins (Second Cycle, 6 4 5 h o n Stream, 779 F) equil exptl comp., comp., mol % compound mol % 1-pentene cis-2-pentene trans-2-pentene 2-methyl-1-butene 3-methyl-1-butene 2-methyl-2-butene Table IX.
3.2 9.7 10.5 22.5 3.9 50.2
3.3 6.3 13.9 20.8 0.4 55.6
max temp, OF
11 212 645 equilibrium
745 778 779 745-800
I
e
0.6
0.6
D
E 0.4
E
s 0.2
0
Formation of Isoparaffins
second cycle time on stream, h
1.0
0.2
0
isoDaraffinlnorma1 paraffin 'ratio butanes
pentanes
2.1 2.9 6.5 -0.7
28 27 17 -2.2
These changes are caused primarily by accumulation of methane and hydrogen in the recycle gas. For example, methane increased from about 40 to 60 mol % in the recycle gas during the second cycle, and hydrogen went from about 10 to 20 mol %. C3 hydrocarbons decreased from about 20 to 5 mol 90. Other hydrocarbons remained fairly constant a t lower concentrations. The total concentration of carbon oxides was unchanged at 10 mol %. 5. Light Product Characteristics. The light hydrocarbon product distributions listed in Table IV are not necessarily representative of the reactor effluent. This is due to the recycle mode of operation and the changing recycle gas composition. True reactor effluent compositions can be estimated from recycle gas compositions and product yields. Such calculations show that C5 olefins are essentially in equilibrium (Table VIII). C5 hydrocarbons are only about 1 mol % of the recycle gas. Although data are not available to allow a similar calculation for other olefins, the above finding suggests that isomer equilibrium among low carbon-number olefins is attained. A characteristic of the process for conversion of methanol to gasoline is that isoparaffins are favored over normal paraffins. Data illustrating this point are shown in Table IX for butanes and pentanes. The experimental iso-/normal pentane ratio is almost an order of magnitude greater than the equilibrium value. 6. Catalyst Aging. Throughout previous discussions, changes in product selectivities within a cycle and from cycle to cycle have been interpreted in terms of changes in catalyst activity. Samples of catalyst from the conversion reactor were not removed periodically, since this would have distorted the activity profile established in the reactor. However, some information regarding catalyst deactivation can be obtained from temperature profiles. Temperature profiles at methanol breakthrough are plotted in Figure 9 for cycles 1 , 3 , 6 , and 9. These profiles clearly show the increased quantity of catalyst needed to complete the conversion. The activity profile which existed in the conversion reactor changed from cycle to cycle. Thus, the amount of catalyst needed to complete the conversion depended on reactor position within a given cycle and also varied from cycle to cycle. Because of the unusual cycle-length behavior in the aging test (see Table 11), there is no clear correspondence between the cycle length and the amount of catalyst for complete conversion. In addition, the time rate of progression of the temperature wave through the reactor varied not only from cycle to cycle, but also within each cycle.
0.6
0.4
0.6
1.0
Fractional Bed Length
Figure 9. Temperature profiles at methanol breakthrough indicate catalyst aging in long-term aging test. Table X. Propane, Aromatics, and Olefins Selectivity Changes as Measures of Catalyst Deactivation position of mid-point of main reaction zone (fractional bed length)
cycle 1 5 9
0.2
0.5
selectivity (wt % of hydrocarbons)
selectivity (wt % of hydrocarbons)
pro- aromapro- aromapane tics olefins pane tics olefins 12 6 4
38 30 26
3 10 16
7 3 2
32 26 22
7 18 24
The data in Figure 9 indicate that at any given time within a particular cycle the location of the main reaction zone will vary from cycle to cycle. A comparison of propane, aromatics, and olefins selectivities at constant reactor position is given in Table X for the first, fifth, and ninth cycles a t conditions of constant reaction zone mid-point. At a fractional bed length of 0.2, which is representative of catalyst at the reactor inlet, the data show a clear trend of catalyst deactivation as determined by decreases in propane and aromatics selectivities and increases in olefins selectivity. Methanol conversion was complete in all cases. Similar trends are also shown by data at a fractional bed length of 0.5, and also as the reaction zone moves down the bed within a cycle. The data in Table X provide some measure of catalyst aging. Some propane is probably formed by cracking and/or hydrogen transfer and would be expected to decrease with aging. The decrease in aromatics and increase in olefins with aging is consistent with the reaction mechanism. The aromatics are formed from olefin intermediates and as the catalyst ages the product distribution reflects a movement back in the reaction path. C. Characterization of Gasoline Produced from Methanol. 1. Product Characterization Tests. Composite hydrocarbon samples from the first six cycles of the long-term aging test in the fixed bed unit were examined for octane qualities, volatility, copper strip corrosion, hydrocarbon-type composition, and freezing point. These composite samples were raw products and were not blended to simulate commercial gasoline which would contain alkylate from processing of the light hydrocarbons from methanol. The ranges of some of the results are summarized in Table XI. These values are consistent with those of commercially marketed gasolines, except for the distillations. In distillations of samples with 3 w t % or more of durene (174 O F freezing point and 386 O F boiling point), the recoveries were abnormally low (i.e., 90 vol %). Durene solidified in the 32 O F condenser tube which caused plugging of the
534
Ind. Eng. Chem. Process Des. Dev., Vol. 18, No. 3, 1979
Table XI. Ranges o f Results of Product Characterization Tests octanes research clear (ASTM D2699) research + 3.17 g of Pb (ASTM D2699) motor clear (ASTM D2700) motor + 3.17 g of Pb (ASTM D2700) distribution clear (ASTM D2886) distribution + 3.17 g of Pb (ASTM D2886) API gravity (ASTM D287) Reid vapor pressure, l b (ASTM D2551) distillation, 'F (ASTM D86) IBP 20% 50% 80% EP hydrocarbon types, vol % (ASTM D1319) aromatics olefins saturates
Table XII. Additive Concentrations
93.9-97.O 99.8-101.6 83.2-86.0 89.9-91.3 92.8-94.6 98.7-100.5 53.1-60.1 11.5-13.9 80-85 144-154 221-253 313-330 404-427 29.4-46.7 5.5-17.8 47.8-54.1
apparatus. Blends of the methanol-derived gasoline and other gasolines with durene concentrations of 2 wt 70 or less gave satisfactory recoveries (>95 vol %). No corrosive tendencies were observed in the copper strip corrosion test (ASTM D130); a rating of 1A was attained by the six composite gasoline samples (3 h at 122 OF). Their freezing points were between -29 and -32 OF. There are normally no commercial gasoline specifications for freezing point, however. Filterability quality tests were conducted on two separate composite samples of gasoline. This test indicates the potential tendency of a fuel to plug small-pored, in-line fuel filters in cars. Both composites had acceptable filterability ratings, even at very cold temperatures. 2. Stability Testing. The existent gum test (ASTM D381) measures the oxidation products (gum) formed in the gasoline prior to or during the comparatively mild conditions of the test procedure. A heptane extraction step removes any nonvolatile additives (as well as very high boiling heptane soluble compounds) present so that they are not included in the weight of gum. The potential gum test is conducted in a manner described by the ASTM D873 procedure; a fuel sample is pressured with oxygen to 100 psig in a bomb and the temperature is raised to 212 O F for a period of 5 h. The amount of gum (also heptane-washed to exclude nonvolatile additives) formed under these severe conditions is an indication of gumforming tendencies during field storage. A potential gum of more than a few mg/100 mL higher than the existent gum is strong evidence of product instability. During the long-term aging test, the potential gum levels of gasoline samples from each material balance were determined. Two different additive concentrations (see Table XII) were used. At the higher additive concentrations, satisfactory stability (assuming 10 mg potential gum maximum/ 100 mL) was attained throughout all nine cycles of the long-term aging test. Low potential gum levels were
metal deactivator" antioxidant"
low
high
1 2.5
5 15
" Units are lb/1000 bbl. observed even after the occurrence of methanol breakthrough. During the first five cycles, potential gum levels were usually satisfactory at the lower additive concentrations until the onset of methanol breakthrough. However, they became high about midway through the sixth cycle. Similar behavior was also observed in the last three cycles. Concluding Remarks Conditions for the successful operation of the fixed bed methanol-to-gasoline process have been defined, and the performances of the dehydration and conversion catalysts have been demonstrated in a long-term aging test of over 200 days. Changes in product selectivities during individual cycles and from cycle to cycle have been delineated. The properties of gasoline product are quite comparable to those of conventional commercial gasolines. Gasoline stability (i.e., potential gum formation) is satisfactory at reasonable additive concentrations. The results of this process study have been used to develop the design basis for a 100 BPD fixed bed pilot plant. Base cases for normal operation, regeneration, and start-up were established. The scale-up to a 100 BPD pilot plant or even larger unit from the small process development unit should be reasonably straightforward. The most significant result of the current study is that the technical feasibility of the fixed bed methanol-togasoline process has been demonstrated. Combination of this process with the commercially proven technology for the conversion of coal to methanol provides an alternative route for the production of high octane gasoline from coal. Literature Cited Chang, C. D., Kuo, J. C. W., Lang, W. H., Jacob, S.M., Wise, J. J. Silvestri, A. J., Ind. Eng. Chem. Process Des. Dev., 17. 255 (1978). Chang, C. D., Silvestri. A. J., J. Catal., 47, 249 (1977). Daviduk, N., Maziuk, J., Wise, J. J., paper presented at Eleventh Intersociety Energy Conversion Engineering Conference, State Line, Nev., Sept. 1976. Liederman. D., Jacob, S. M.. VoRz, S.E., Wise, J. J., Ind. Ena. Chem. Process Des. Dev., 17, 340 (1978). Meisel, S.L., McCullough, J. P., Lechthaler, C. H., Weisz, P. E., CHEMTECH, 6 86 11976\ - . -I I
Meis'el, S. L., McCullough, J. P., Lechthaler, C. H., Weisz, P. B., paper presented at the 174th National Meeting of American Chemical Society, Chicago, Ill., Aug 1977. Stockinger, J. H., J . Chromatogr. Sci., 15, 198 (1978). Wise, J. J., Silvestri, A. J.. paper presented at Third Annual International Conference of Coal Gasification and Liquefaction, Pbburgh, Pa., Aug 1976 Oil Gas J., 141 (Nov 22, 1976) Yurchak, S.,Wise. J. J., Silvestri, A. J., Chang, C. D., paper pesented at Eleventh Middle Atlantic Regional Meeting, American Chemical Society, Newark, Del., April 1977.
Received for review July 31, 1978 Accepted January 8,1979 This work was conducted under the Department of Energy (DOE) which was jointly funded by DOE Contract No. E(49-18)-1773, and Mobil Research and Development Corporation.
