Pyrolytic and Catalytic Decomposition of Hydrocarbons
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VLADlMlR HAENSEL and MELVIN J . STERBA
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UNIVERSAL OIL PRODUCTS COMPANY, RIVERSIDE, ILL.
This review, in common with those published previously (28-30), summarizes publications which have appeared in the literature during the year ending in M a y 1951. Experimental studies reported have improved understanding of the mechanisms of thermal decomposition reactions-particularly the use of deuterium as a tracer in establishing the identity of that portion of a reactant surviving afler passing through a reaction zone. The production of acetylene b y various thermal decomposition methods has been given industrial attention, apd the role of thermal cracking in the possible production of a marketable oil from Canadian bituminous sands has been investigated experimentally. Important contributions describe the use of tracers in Further elucidating the mechanism of catalytic cracking and the participation of the catalyst in the reaction. In catalytic reforming, attention has been directed toward the production of aromatics, particularly benzene, toluene, and the xylenes from appropriate petroleum fractions.
N ACCORDANCE with the procedure established in previous years, the review of pyrolytic and catalytic decomposition of hydrocarbons covering the last 12 months includes thermal cracking, catalytic cracking, and catalytic reforming.
THERMAL CRACKING It is usual, in hydrocarbon decomposition experiments, to reckon the extent of decomposition b y noting the amount of original unreacted hydrocarbon in the recovered product stream. However, in the decomposition of butane, for instance, the recovered butane in the product stream may not be entirely unreacted original material but may include some reconstructed or regenerated butane. Hurd and Azorlosa (56)considered this situation by studying the thermal decomposition of butane and propylene containing deuterium. I n the decomposition of butane-d the deuterium content of the butane in the recovered product was 2.3 to 5.3% less than in the butane feed. As this change in deuterium concentration was outside the limits of experimental error it was concluded that some of the butane in the product stream was synthetic material and not exclusively the original unchanged butane charged. I n the decomposition of propene-&, however, the deuterium content of the propene recovered in the products was the same a~ that of the feed. These results were explained satisfactorily by the free radical mechanism. Bywater and Steacie reported on the mercury photosensitized decomposition of ethane (8),propane (7), and the trvo butanes (6). I n ethane they found that the ethane radical became unstable at around 400' C. Below this temperature butane and hydrogen were the only significant reaction products. Below 200" C. hydrogen and hexanes were the only reaction products from propane decomposition, and above 300" C. appreciable amounts of methane and ethylene were formed. However, in the absence of ultraviolet readiation the thermal decomposition of propane was found to be negligible even a t 450' C. and 20 minutes reaction time. The photosensitized reactions of both n-butane and isobutane gave a variety of decomposition products of butyl radicals. At temperatures above 250' C. nbutane decomposition resulted in the formation of methane, ethane, ethylene, and propylene. From experimental results the mechanism of and the activation energy for each of the various radical decomposition steps were derived. I n the thermal decomposition of isobutylene at 850" to 900" C., 50 to 22 mm. pressure, and less than '/z second contact time,
Rice and Wall ( 6 4 ) observed that tlie major reaction products were methane, allene, and methylacetylene. When 2 t o 3 kilograms of isobutylene were charged through quartz or chrome steel reactors a sufficiently large quantity of products was made to isolate a small amount of high boiling - aromatic residue which consisted chiefly of benzene, toluene, If these processing conditions were not and naphthalene. followed, large quantities of tar, carbon, and heavy oils were formed. Deanesly and Watkins ( 1 7 ) reported on their experimental work of ethylene production by autothermic cracking, a name they have given to a process in which the endothermic heat of cracking is supplied in situ by the introduction of air or osj.gen with the feed, thus eliminating the necessity of transferring heat through tube walls. This technique avoids the use of alloy steel reactor tubes which are necessary at the high temperatures required for cracking ethane and propane. With the proper use of heat exchange between product and feeds to the reactor, the thermochemical efficiency of the process is such tliat thc dilution of cracked products with carbon oxides and nitrogen is nt a minimum. Pilot plant results for the autothermic cracking of ethane and propane at conversions per pass of 40 to nearly 100c;, have been presented and compared with previously publisht~d tubular cracking results. These high conversions, of ethane for instance, are in contrast to those of 50 to 60yoobtainable under practical conditions in tubular reactors. Daniels (16) has given a brief summary of the Schoch electric. discharge process for producing acetylene from methane or heavier hydrocarbons. In this process the hydrocarbon feed is exposed in a glow discharge for a short time, and the resulting reaction mixhure is resolved into light gases, acetylene, and heavier hydrocarbons which can be recycled for further conversion. A reaction chamber consists of three stationary but adjustable electrodes positioned about a single rotating electrode which resembles the impeller of a centrifugal blower. Arranged in this way, the rotary electrode moves the feed stock into the spark discharge area. The operating reaction pressure is slightly above atmospheric, and because somewhat over half of the discharge energy is liberated as sensible heat, the chamber temperature is limited to about 550' F. by the recirculation of gas through a cooler, after preliminary removal of carbon with a cyclone. An electrical power requirement of about 5.5 kn..-Iir. per pound of acetylene is required. Ilappel and Marsel ( 3 2 ) have made an extensive comparison of acetylene production costs by various methods. Although practically all acetylene is being made currently by the calcium carbide procesq at a cost of about 1 1 cents per pound, the author' have estimated production costs as low as 6 cents per pound for the Wulff regenerative thermal cracking process and intermcdi:itc, costs for the Behoch, Sachse, and air oxidation proccssrs, whew
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using hypersorption as a means of separating acetylene from the reactor effluent. Some of these estimated production costs were shown to be sensitive to variations in the cost of electrical power, and all are based on making 20,000,000 pounds of acetylene per year. One chemical manufacturer plans to install acetylene production facilities at Texas City, Tex., using partial oxidation of methane (natural gas) with oxygen and purification of acetylene from the product effluent with solvent absorption (14,16). The possibility of producing a marketable oil from the bituminous sand deposits in Alberta, Can., has been outlined by Blair ( 3 ) who suggests a sequence of processing steps which include sand mining, separation of the bitumen, thermal coking of the bitumen, desulfurization of the coker distillate, and pipeline delivery of the oil to refining or marketing areas. Bituminous sands contain varying amounts of hydrocarbon u p to 18%; a typical deposit would contain perhaps 12%. The bitumen, having a specific gravity of about 1.02, has been successfully recovered from the sand using water as the separating medium. As a next step, a distillate oil can be produced from the separated bitumen by conventional thermal coking methods. However Gishler ($4)has proposed that the separation and coking steps can be performed in one operation by the use of a fluidized solids technique, charging the bituminous sand to the process. He describes a laboratory glass apparatus in which he has studied the simultaneous distillation and coking of the bitumen at temperatures of 425' to 700' C., using nitrogen as the fluidization medium. The highest yields (75% based on the bitumen in the sand feed) of distillate oil were obtained at 460' to 500' C,; the remaining 25y0 of the bitumen decomposed into nearly equal portions of light hydrocarbon gases and coke which appeared as a thin layer on the grains of sand. Later, Peterson and Gishler (55, 54) used a pilot plant for continuing this study on a larger scale, feeding from 60 to 120 pounds of bituminous sand per hour. This equipment resembled a fluid catalytic cracking unit with its reartor and regenerator vessels situated at the same elevation. Sand particles, from which the bitumen had been distilled, passed from the vessel called the still into the adjacent vessel called the burner where the film of colcc on the sand was burned off with air. The heat of combustion of the coke and additional fuel gas raised the temperature of the sand stream sufficiently so that when circulated back to the still, heat was provided for preheating the raw bituminous sand feed and also for the distillation and decomposition of the bitumen. At 930' to 980' F. still temperatures, up to 86% yields of a distillate with a specific gravity of 0.95 were obtained. From an experimental study of the thermal cracking of 8 catalytic cycle oil a t 750 and 1500 pounds per square inch, Little and Merryfield ( 4 2 ) concluded that the highest octane ratings of the cracked gasoline were obtainable in a once-through operation a t the lower pressure. They found that aniline point reduction of the cycle oil charge in its passage through the reaction zone could be related to conversion and product yields. For a given reduction in aniline point, higher gasoline yields of lower octane ratings were obtained by recycling than in once-through operation. The authors point out that successful cracking of cycle oil at high conversions per pass requires a radiant-type heater having high heat transfer rates and short oil residence times. A description of equipment involved in a recently constructed delayed-coking unit was presented by Fuchs ($0) and Foster (19). This unit, designed to process 9300 barrels per day of 13' A.P.I. reduced crude, was expected to produce 423 tdns of coke per day to be removed from the chambers by hydraulic decoking and subsequently moved to a coke pile with a drag line. Jewel1 and Connor (37) have made an economic comparison of delayed coking with vacuum flashing, propane de-asphalting, and viscosity breaking, used as methods of preparing catalytic cracking feed stocks from reduced crudes. When each of the
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four feed preparation methods were placed successively in an otherwise fixed refining sequence, the delayed coking process was found to have superior earning power when a good coke market or a poor fuel oil market exists. It was pointed out, also, that when the supply of fuel oil exceeds its demand, coking has provided the only complete relief from reducing crude processing or building excessive storage capacity. C H E M I C A L CONCEPTS OF CATALYTIC CRACKNG
A number of interesting contributions l o the knowledgc of catalytic cracking have been made since the writing of the last review on this subject Although apparent agreement has been reached on the carbonium ion mechanism of catalytic cracking, there is no generalized agreement on the way in which the acidity of a cracking catalyst is attained. However, the information presented during the last 12 months does provide further insight on the chemical characterization of cracking catalysts as well as on the reactions promoted by cracking catalysts. Mills and Hindin ( 4 8 ) describe the work on the exchange of oxygen between water and silica gel, alumina gel, silica-alumina cracking catalyst, and activated and nonactivated clays. The rcsults indicate a rapid initial exchange followed by B considerably slower later exchange when the experiments are carried out a t 100" C. using H2018. The over-all reaction is represented as follows: Mz0,'6
+ HzO"
5= MZO,"
+
I'120'6
The authors believe that the initial rapid exchange takes place by the reaction of the oxide present in the hydroxyl form as [SiOH] and HOII, whereas the later slower exchange is brought about by a hydrolysis of [SiOSi] linkages to produce 2jSiOHl. I n contrast to the extensive exchange using silica, alumina, and composites of silica alumina, there is no exchange observed using kaolin or bentonite clays. Sulfuric acid activation of the bentonite clay produccs a material which undergoes oxygen exchange. It is postulated that nonactivated clays exhibit no oxygen exchange partly because of IOU- surface area and partly because of the inherent stabilit,y of the crystallinc materials towards water. The extensive exchange of oxygen in the cese of silica, alumina, and composites of silica alumina is a good indication of thc ability of the surface and, in the case of the cracking catalyst, a labile surface i s necessary in order t,o permit Ibe change of coordination of aluminum in the catalyst froin six to four. This particular change in coordination is a necessary part of the cracking catalyst structure hypot,hesis proposed by Milliken, Mills, and Oblad ( 4 7 ) . The hydrogen exchange between cracking catalyst and butanes was studied by Hindin, Mills, and Oblad (54.). The experiments involved the interaction of deuterated butanes with the silicaalumina catalyst a t 150' C. for 1 to 4 hours. In other cases a deuterated catalyst was used with n-butane and isobutane. %-Butane undergoes minor exchange as compared with a con. siderable exchange-in isobutane. The hydrogen atoms attnched to primary carbon atoms in the isobutane molecule exchange readily, but the tertiary hydrogen does not appear t o uiidergo the exchange reaction. The extent of hydration of the catalyst has a profound effect on the ability of the catalyst to promote: the exchange reaction. When the catalyst is dried at a high temperature, such as 525" C., the exchange ability is nearly lost; either low temperature drying or rehydration restores the activity. The authors postulate three steps in the exchange reaction: ( a ) complex formation between catalyst and hydrocarbon, ( b ) exchange of hydrogen bctwcen catalyst and complex, and (c) hydride ion transfer between isobutane and deuterated complex. Holm and Blue (36) describe the study of the hydrogen-deuterium exchange activity of silica-a1dmina catalysts. With pure alumina a t 300' C. and 20,400 gaseous space velocity the exchange amounts to 88.1% of the theoretical; with pure silica
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U.O.P. Platforming Unit, Kendsll Refining Co., Bradford, Pa.
at the same temperature and 2230 gaseous space velocity the exchange is only 1.1% of the theoretical. As the silica content is increased, starting with pure alumina, there are considerable variations in the exchange activity particularly in 0 to 10% silicon dioxide concentration. At higher silica contents the reduction in activity is more gradual. A t the lower silica contents high calcination temperatures of the composite increased the exchange activity. There are indications that silica at low concentrations acts &s a stabilizer when the composites are heated to about 800' C. Since hydrogenation reactions are postulated to proceed by d h i a t i o n into hydrogen atoms, and the exchange reaction is a good indication of hydrogenation activity, a series of catalysts with variable silica content were tested for the ethylene hydrogenation reaction. At 500" C., 650 gaseous space velocity, and atmospheric pressure, the hydrogenation proceeds more rapidly with catalysts having a lower silica content. The hydrogenation activity of these composites, relative to the usual hydrogenation catalysts is, of course, very low. In another investigation of the same catalyst system, Blue and Engle (6) studied the hydrogen transfer reaction, using Decalin at^ hydrogen donor and butenes as hydrogen acceptor. CisDecalin reacts more extensively than trans-Decalin, and isobutylene undergoes hydrogen transfer more readily than the nbutylenes. The most active silica-alumina catalysts were those containing 60 to 9OV0 silica. This is quite different from the previously quoted concentration ranges required for the highest hydrogen-deuterium exchange. It is postulated that the hydrogen transfer reaction is an ionic reaction, and cis-Decalin provides a much more accessible location of two hydrogens to be removed during the course of hydrogen transfer. Further insight on the mechanism of the catalytic cracking reaction can be gained from the work of Roberts and Good (65) who studied the effect of nuclear substitution on the rate of depropylation of cumene in the presence o f , a silica-zirconiaalumina catalyst. It was found that a t both 450" and 400" C. using 65 and 13 moles of charge per liter of catalyst per hour, respectively, the same relative order of reactivity is maintained. It was established that the rates of depropylation decrease in
the following order: 1,3-dimethyl-4isopropylbenzene,77%; pcymene, 60%; 1,3-dimethyl-5-isopropylbenzene,57%; cumene 39%; pchlorocumene, 25%; trichlorocumenes, 8%. This order of reactivities is in good agreement with the relative activation energies for electrophilic displacement of a propyl group by a hydrogen. These results provide additional proof of the carbonium ion mechanism of catalytic cracking reactions.
INDUSTRIAL FEATURES A bench scale fluid catalytic cracking unit has been developed which is capable of duplicating the important 'operating conditions existing in full scale commercial fluid units ($66).This miniature pilot plant, in which catalyst circulates continuously, will operate with as little as 1 liter of catalyst and is said to be useful for evaluating laboratory preparations of catalyst or small quantities of feed stock. Product distribution obtained by its use is generally similar to that derived from large scale fluid units. The evolution of fluid catalytic cracking plant design has been traced briefly by Armstrong (1) who describes improvements in spent catalyst stripping, catalyst recovery procedures, catalyst handling techniques, regenerator linings, vessel arrangement, and oil injector design, and comments on the use of centrifugal compression for handling the wet gas product. A recent modification in fluid catalytic cracking plant design has been named Orthoflow ( 5 1 ) because the circulating catalyst stream is conveyed in straight pipes. Very little descriptive matter is to be found in the literature, although reference is made to a central pipe which carries a mixture of raw oil and regenerated catalyst straight upward to the reactor. An excellent review of the entire field of catalytic cracking is presented by Sittig (67-70) in a series of articles describing the history and recent developments relating to the McAfee aluminum chloride, Houdry, Thermofor, fluid, cycloversion, and suspensoid processes. Basic differences between thermal and catalytic cracking have been pointed out and ascribed to the distinct selective function of the catalyst in directing the course of
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the decomposition reaction to produce more desirable materials than is possible by thermal decomposition. That the cracking catalyst has an accelerating effect on the reaction is of little industrial importance according to Sittig. What is believed to be the largest catalytic cracking plant in the world is being built this year according to announcements and brief descriptions of the design (49, 55, 61). This fluid-type unit will process a reactor charge of 60,000 barrels per day, and will have its spent. catalyst stripper within the reactor and above the main grid. As reported by Johnson (38) a fluid catalytic crackmg unit has recently established a new record of 677 days of continuous operation. During this run 3,394,540 barrels of gasoline were produced from 6,807,219 barrels of fresh gas oil feed, and the catalyst consumption amounted to 0.212 pound per barrel of oil processed. In a description of one of the first Houdriflow units to be put into operation, Barton ( 2 ) has developed the specific refinery requirements which were to be met by the catalytic cracking unit. The initial experimental operating period was discussed, and modifications to the original catalyst lift installation were described. Two of the four original 19-inch diameter catalyst lift pipes were replaced by banks of multiple lifts each containing seven pipes of %inch diameters. Attrition of clay catalyst was reported to be only half as much when the multiple lift replaced the single pipe. -4n indication is given that the attrition rate of synthetic bead catalyst was only 70% of that for the clay type catalyst. Thornton (71 ) described another Houdriflow installation which was designed to operate at a catalyst-to-oil ratio of 8 and with a feed preheat temperature of 670" F. The same author has described an interesting application of gamma rays in the measurement of catalyst levels in this and other cracking units (7B). In this instrument, named a Gagetron, the essential components are a source of gamma radiation, a Geiger tube to detect varying intensities of radiation as affected by the amount of intervening catalyst, and electronic circuits to actuate recording devices. One company has announced (4, 79, 80, 83) that it has arranged to niake available small prefabricated modified thermofor caraiytic cracking (T.C.C.) units in packaged form, completely shop fabricated and ready for erection in an oil derrick type of support,ing structure a t the refinery. These units will be desi.gned to contain the reactor and regenerator in a single vertical column; they employ an air lift for transporting catalyst from ground level to the top of the structure. The catalyst lift system was tested (77) in an experimental 16-inch diameter pipe of 200-foot length which was found capable of transporting 250 tons per hour of catalyst, an amount sufficient to provide recommended catalystcto-oil ratios of from 3 to 5 for 15,000barrrl-per-day units. The designers expect wide use of a newly developed chrome bead catalyst (77) which is said to have normal cracking characteristics but reduces after-burning in the re generation zone by promoting higher carbon dioxide-carbon monoxide ratim than are obtainable with conventional synthetic catalysts. The first modified T.C.C. units employing an air lift were put ~ 73). It was stated into operation during the past year ( 4 ~ 360, (7'3)that a somewhat harder bead catalyst is used in a t least one of these units instead of the catalyst developed for the early T.C.C. units. This catalyst contains about 0.003% chromium t o minimize after-burning in catalyst regeneration. The higherthan-expected catalyst attrition experienced during the first 2 weeks of one of the operations was attributed to the use of fresh catalyst of high moisture content (49). Accordingly, a dryer was installed to reduce the moisture in the fresh catalyst make-up stream. The modification of an early T.C.C. unit to permit operation on a partially liquid feed hae been discussed by Weber (81) and Uhl (?,$). This was accomplished by spraying the liquid portion
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of the gas oil feed within a falling cylindrical curtain of hot regenerated catalyst a t the top of the reactor. Performance test data are given for the processing of a Wyoming gas oil to a conversion of 76%in recycle operation a t an average reactor temperature of 875" F. The liquid portion of the feed to the reactor entered a t 698' F.,and the vapors entered a t 890' F. A statistical report has been prepared by Dougherty ( 1 8 ) for the maintenance of a 5500-barrel-per-day T.C.C. unit in a small refinery. This unit is normally operated by three men and during a shutdown over 40 men of various crafts are used for the usual maintenance work involved. During its history the cracking unit has been turned around three times. In these turnarounds the down time varied from 264 to 382 hours, and the man-hours required for normal maintenance work varied from 3362 to 8553. As an extension of experimental work and correlations previously reported for once-through moving bed catalytic cracking, Maerker, Schall, and Dart (44)studied the effects of process variables on product distribution in recycle operation. They arranged their correlations so that for the East Texas gas oil studied it is possible to predict the product distribution at a given set of operating conditions or to make this prediction at a given temperature and recycle ratio when the yield of one product is known. It is also possible to predict operating conditions required to obtain a specified yield of one of the products. The authors point out that their correlation can be adapted to other charge stocks by establishing reference lines on !heir charts from three to five properly chosen pilot plant tests. Viland (76) has described a laboratory cracking catalyst aging tester which is a miniature fixed-bed cracking unit, automatically operated over a sufficiently long period of time to measure the rate of catalyst activity decline and the changes in product distribution. It is useful for evaluting different catalysts in this respect or for comparing the chronic effects of contaminants in various feed stocks on the selectivity and cracking activity decline of a particular catalyst. The importance of the effect of impurities on cracking catalyst performance has been generally recognized, and the determination of the extents of contamination has been investigated from the standpoint of both the catalyst and the oil charge, Thus, Harmon and Russell (33)describe the use of a spectrographic method of analysis of catalysts for iron, vanadium, nickel, and sodium. 10% of the amount The method is stated to be accurate within =t present, and the analysis can be made in le33 than 4 hours for a group of samples. Carlson and Gunn (9) have reported on a method of determining the extent of metallic contamination of petroleum oils. The method involves the direct spectrographic inspection of carbon electrodes imprepabed with the sample of the oil. The method is reported to be rapid and capable of providing better than semiquantitative accuracy over the established calibration ranges. The metals investigated include iron, nickel, chromium, vanadium, calcium, sodium, silicon, and aluminum. In some instances where especially high boiling gas-oil fractions with high trace metal contents are processed, it becomes feasible to use more fresh catalyst than is necessary to replace normal physical losses in order to maintain proper equilibrium selectivity and activity of the catalyst. This requires that some of the circulating inventory must be withdrawn, and any such direct wit,hdrawals comprise catalysts of all ages in the unit, including the last fresh material added. As cracking catalysts age with use, they become more dense because of physical changes occurring with a loss in surface area. This change in density suggests a method of selectivity removing the oldest and least desirable particles from the circulating inventory. Viland ( 7 6 ) has described equipment used for this purpose, and has given results obtained by the gravity separation of inactive catalyst from a fluid catalytic cracking plant. The apparatus consists of a centrifugal separator which produces a fine and a coarse frac-
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tion, followed by a classification table which rejects dense particles from the coarse catalyst fraction. This separation equipment operates on a slip stream of 50 tons per day and can be regulated to reject any amount of denst. material desired (figures of 2 to 4.7 tons per day were shown). Catalyst not rejected is returned t o the circulating system. The problems concerned with the use of a natural clay, first for decolorizing lubricating oil stocks, and subsequently as a fluid cracking catalyst have been summarized by Weber (78, 83). By modifying the grinding and activation of conventional Filtrol, Grade D, this catalyst was found to have satisfactory color adsorption properties, and could be used for the two refining functions in series. The principle of conducting chemical reactions in fluidized beds of catalyst particles has been applied so extensively during the past decade in the cracking of petroleum hydrocarbons that it has' become accepted generally as a unit operation of chemical engineering. Although industrial application of the principle was well ahead of an understanding of the fundamentals involved, a great deal of attention has been given in recent years to methodical study of the dynamics, mass transfer, and heat transfer in fluidized systems. One of the most recent studies concerned with fluidization dynamics was presented by Miller and Logwinuk (46) who have developed an equation to predict the critical mass velocity of vapor required to initiate fluidization in a bed of solid particles. Wilhelm and Valentine (84) have studied the dynamics of vapor-solid particle systems over a wide range of gas velocities by introducing the particles into an u p ward rising stream of gas. At low velocities the particler dropped downward to form dense or fluidized beds, and beyond a limiting gas velocity the particles were transported upward with the moving gas stream. A comprehensive literature survey on the general subject of fluidization was presented by Leva (40) in a recent chemical engineering unit operation review. I n a discussion of the techniques of contacting fluids and solids, Gilliland (2.9)has presented some results of experiments designed to study the degree of mixing of solids and the extent of mixing in the gas phase in fluidized systems. T o study the mixing in the solid phase, a stream of tracer particles was introduced about halfway up into a fluidized bed t o which a stream of solids was fed a t the bottom. The concentration of tracer particles in the, bed was measured at various points over its entire length, and the results indicate considerable mixing of the tracer particles with the solid below the point of addition of the tracer stream. In a similar fashion, mixing in the gas phase was studied by introducing a tracer gas midway into a fluidized bed and measuring its concentration a t various points below the point of introduction of the tracer. Considerable downmixing was observed, and attributed to the rapid stirring action of the solids under conditions of fluidization. Gamson (22) has presented an extensive mathematical t r e a t ment of heat and mass transfer in fluid-solid systems in an attempt to derive generalized expressions which would apply to such transfer in both fixed-bed and fluidized solid systems. Hall and Jolley (31) have reviewed the characteristics of fluidized solid beds and have discussed their application to catalytic cracking, Fischer-Tropsch synthesis, phthalic anhydride manufacture, gasification of coal, shale oil retorting, and lime burning. I n an experimental study of the elutriation of fines from fluidized solid systems, Leva (39)observed that the rate of removal of fines could be expressed as a kinetic function resembling the first-order reaction law. He found that the elutriation velocity constant could be related to gas velocities in much the same way as reaction velocity constants are related to temperature in chemical reactions. During the past year announcements have been made which indicate rather large expansion in cracking catalyst manufacture. One manufacturer intends to build new facilities in the vicinity
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of Salt Lake City, Utah, capable of producing between 30,000 and 50,000 tons per year of natural catalysts from Utah clays (11). Another manufacturer plans to produce primarily microspheroidal synthetic cracking catalyst in a new plant in East Chicago, Ill.; with this plant expansion its total production capacity would be 3000 tons per month (10). A new installation at Weeks, La., will produce an undisclosed amount of catalyst using a recently developed process (If?). During the last war catalytic cracking units processed rather light gas oils to produce aviation gasoline base stock and light hydrocarbons used as alkylation feed stock. In the post-war period, catalytic cracking feed stocks have become distinctly heavier with respect to their boiling range and molecular weight because refiners have used the light gas oils as directly marketable products and because they have also employed the heaviest gas oil boiling just below asphalt in order to obtain the maximum amount of gas oil feed from a given quantity of crude. In general, the ole& content of catalytically cracked products becomes higher as the molecular weight of the feed is increased. I n an evaluation of methods used during the last war for producing military fuels required at present, Read (62, 63) points out that most aviation base stocks obtained by the distillation of catalytically cracked gasolines made today would be of unsuitable quality without further processing. Retreating methods such as were used during the last war would result in low yields because of the high olefin content of the material being retreated. Read concludes that existing catalytic cracking units can be utilized advantageously by continuing the processing of heavy gas oils to produce motor fuel and light hydrocarbon feed stocks for alkylation units, Thus, the light gas oils are left to supply the probable demand for jet and Diesel fuels Read points out that high yields of base stock suitable for aviation gasoline blending can be produced by retreating the olefinic catalytic gasoline cracked from heavy gas oils, using the Platforming process and presenta experimental data bearing on this retreating technique. One rather detailed survey, published in March 1951 (52),indicates a domestic crude refining capacity of just over 7,000,000 barrels per day. Catalytic cracking capacity is approaching that for thermal cracking as shown by the following figures: Type of Cracking Catalytic Thermal
Capacity, Bbl./Dtty 1,887,485 2,334,630
CATALYTIC REFORMING Because of the increasing demand for benzene as a basic chemical raw material, and for other aromatic hydrocarbons to supply the needs of the plastics industry, the synthetic rubber program, and aviation gasoline, the oil industry has been called on to develop quagtity production of aromatics. According to Linz ( 4 1 ) the Petroleum Administration for Defense was assigned the task of supervising this program in the fall of 1950. At that time Weber (82) indicated that a technical committee on aromatics was being formed in the Military Petroleum Advisory Board to determine what steps might be taken by tho refining industry, Shortly aftorward t w o oil companies announced that they had developed catalytic reforming processes, and the already existing U.O.P. Platforming process demonstrated in a trial commercial run that it was capable of producing high yields of aromatics from petroleum fractions. At least two of the eight Hydroforming units in operation a t the end of the last war have been producing benzene and other aromatics in recent years. Thus, the petroleum refining industry has already developed tools, some of which are in use, for mccting increasing aromatic demands by augmenting the benzene output of coal distillation.
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The application of U.O.P. Platforming to the production of aromatics has been discussed by Haensel and Berger (27) who present test results from a commercial unit processing a 202' to 224' F. feed stock from which the gross molal yield of aromatics was 95% of theoretical. From this feed stock, containing 62% naphthenes and 10% toluene, there was a production of 1.0% benzene, 55.801, toluene, 1.2% ethylbenzene, and 6.6% xylenes, all percentages being on a weight basis. Similar data are presented for pilot plant operations in which benzene and CS aromatics were produced from appropriate virgin petroleum fractions. The processing of a 150" to 190" F. cut from a West Coast crude resulted in a 44 weight % yield of benzene. The first U.O.P. Platforming unit to be put into operation completed its second run a t the end of 1950. The catalyst life for this run amounted to 42.3 barrels of fresh feed processed per pound of catalyst. This represents a catalyst cost of about 9 cents per barrel of fresh feed (86). Based largely on previously published material, Fulton (21) has presented a review of the Platforming process with special emphasis on its applicability in the improvement of natural gasoline octane numbers. Another general description of the process, its flow diagram and of the equipment involved has been presented by Wright (86). In April 1951, there were 66,050 barrels per day of Platforming capacity either in operation or under construction in 14 different refineries (56). These units vary in size from 900 to 16,000 barrels per day of design capacity.
about the same as found in naturally occurring CS aromatic fractions in petroleum. These similarities are summarized in Table I.
LITERATURE CITED (1) Armstrong, R. T., Petroleum Refiner, 29, No. 9, 190 (1950). (2) Barton, P. D., Oil Gas J . , 49, N o . 47, 232 (1951). (3) Blair, S. hl., "Report on Alberta Bituminous Sands," Government of Province of Alberta, Can. (December 1950). (4) Bland, W. F., PetroleumPr'ocessing, 7, 711 (1950). (5) Blue, R . W., and Engle, C. J., IKD.ENG.CHEM.,43, 494 (1951). (6) Bywater, S., and Steaeie, E. W. R., J . Chem. P h y s . , 19, 172 (195 1). (7) Ihid., p. 319. (8) Ihid., p. 326. (9) Carlson, ?VI,T., and Gunn, E. L.. A n a l . C h n . , 22, 1118 (1950). 1 Chem. Eng., 57, No. 8, 172 (1950). Ibid., p. 176. Chem. E n g . N e w s , 29,290 (1951). Ihid., p. 761. Ihzd., p. 1799. Ihid., p. 1957. Daniels, L. S., Petrolelurn Refiner, 29, S o . 9, 221 (1950).
Deanesly, R. M., and Watkins, C. H., Chem. Eng. Progress, 47, 134 (1951).
Dougherty, W.F., Petroleum Engr., 22, No. 5, C-33 (1950). Foster, A. L., Ibid., 23, KO.4, C-53 (1951). Fuchs, 0. A , , Petroleum Processing, 5; 1058 (1950). Fulton, W. F., Petroleum Befiner, 29, KO. 12, 109 (1950). Gamson, B. TV., Chem. Eng. Progress, 47, 19 (1951). Gilliland, E. R., Cun. Chsm. Process I n d . , 34, No. 8, 632 (1950).
Gishler, P. E., Can. J . Research, F27, 104 (1949). Glasgow, A. R., Willingham, C. B., and Rossini, F . D., IND. ENG.CHEX.?41, 2292 (1949). Grote, H. W., Hoekstra, J., and Tobiasson, G. T., Ibid., 43, 545 (1951). Haensel, V., and Berger, C. V., Petroleum Processing, 6 , 264
Table I. Composition of CS Aromatic Fractions
o-Xylene m-Xylene p-Xylene Ethylbenzene Temperature,
F.
Reference
Hydroforming
Platforming
Catalytic Cracking
20 43 17
23 40 21 16 100
20 50
2 100 900-1000 (46)
20
ThermoVirgin dynamic Petroleum Equilibrium 20 50
20
-
-
..
sa
..
850
(28)
(66)
(66)
,
($7)
100
(1948).
47 21 9
in -
100
(1951).
J'.I ,, IND.ESG. CHEM.,40, 1660 Haensel, V., and Sterba, &
23
in -
Ibid., 41, 1914 (1949). Ibid., 42, 1739 (1950). Hall, C. C., and Jolley, L. J., Petroleum ( L o d o n ) , 113, 217
100
A recently developed catalytic reforming process was anaounced during the past year ( I S , 57-59). This process employs fixed beds of pelleted catalyst at t,emperatures of 850" to 1000" F., pessures of 300 to 700 pounds per square inch, and a recycle gas stream containing hydrogen. With normal operating condit2ons and gasoline feed stocks the process is said to be nonregenerative, but a t extreme reforming severities or when processing feed stocks containing mat,eriai in the kerosene boiling range periodic catalyst regeneration may be required to remove small quantities of coke. Processing data are presented to illustrate yields obtainable in gasoline reforming for octane number improvement, aromatic production, and when making aviation gasoline blending stock. Another catalytic reforming process was announced in 1950 (50), although no description of its features was given. The production of aromatics using the Hydroforming process has been considered by Marshall (45) who describes operations a t the Baytown Ordnance Works for making toluene during the last war. In the postwar period the plant has been used for octane number improvement of -naphthas, for production of high purity solvents, and for making aviation gasoline blending stocks. It is interesting to note the similarity in distribution of compounds in the Cs aromatic group produced by several catalytic decomposition processes. This distribution among the CS aromatic isomers is in good agreement with the relative amounts calculated for thermodynamic equilibrium a t processing temperatures, and as noted by Rossini (66) this distribution is
202 1
(1950).
Happel, J., and Marsel, C., Chem. I n d . , 68, No. 15, 17 (1951). Harmon, D. D., and Russell, R. G., Anal. Chem., 23, 125 (1951). Hindin, S. G., Mills, G. A., and Oblad, A. G., J . Am. Chem. SOC., 73, 278 (1951). Holm, V. C. F., and Blue, R. W., IND.ENG.CHEM.,43, 501 (1951).
Hurd, C. D., and Aeorlosa, J. L.. J . Am. Chem. Soc., 73, 33 (1951).
Jewel], J. W., Jr., and Conqor, J. P., Petroleum Refiner, 30,No. 2, 75 (1961).
Johnson, R. L., Oil Gas J . , 49, No. 45, 92 (1951). Leva, M., Chem. E n g . Progress, 47, 39 (1951). Leva, M. and Weintraub, M . , IND. ENG.CHEY.,43, 90 (1951). Line, B. F., Oil Gas J., 49, No. 24, 60 (1950). Little, D. M., and Merryfield, 6.E., Ibid., 49, S o . 47, 242 (1951).
MeCaslin, L. S., Jr., Ibid.. 49, No. 38, 168 (1951). Maerker, J. B., Schall, J. W., and Dart, J. C., Chem. Eng. Progress, 47, 95 (1951). Marshall, C. H., Ibid., 46, 313 (1950). Miller, C. 0 . and Logwinuk, A. K., IND.ENG.CHEM.,43, 1220 (1951).
Milliken, T. H., Jr., Mills, G. A,, and Oblad, 4.G., Trans. Farad a y Soc., Symposium on Heteregeneous Catalysis, 1950. Mills, G. A,, and Hindin, S. G., J . Am. Chem. S o c . , 72, 5549 119.5Oi. \ _ _ _ _ ,
(49) (50) (51) (52) (53)
Oil Gas J . , 49, No. 14, 42 (1950). Ibid., S o . 27, 190. Ibid., No. 45, 112 (1951). Ibid., No. 47, 329 (1951). Peterson, W.S., and Gishler, P. E., Can. J . Research, F28, 6 2
(1950). (54) Peterson, W. S., and'Gishler, P. E., Petroleum Engr.,-23,No. 4 C-66 (1951). (55) Petroleum Engr., 23, No. 3, C-39 (1951). (56) Ibid.. No. 4. C-33. i57) Ibid.,C-47. (58) Petroleum Processing, 6, 249 (1951).
2022
INDUSTRIAL AND ENGINEERING CHEMISTRY
(59)Petroleum Refiner,30,N o . 3,86 (1951). (60)Ibid., p. 95. (61) Zbid., p. 144. OiE Gus J.,49, No. 45,68 (1951). (62) Read, D., (63) Read, D., Petroleum Refiner,30,No. 3,130 (1951). (64) Rice, F. 0.. and Wall, L. A., J. Am. Chem. SOC.,72, 3967 (1950). (65) Roberts, R. M., and Good, G. M., I b i d , , 73, 1320 (1951). (66) Rossini, F. O.,“Reilly Lectures,” Vol. 111, South Bend, Ind., The Univ. of Notre Dame, 1949. (67) Sittig, M., Petroleum Refiner,29, N o . 6,91 (1950). (68)Ibid., No.8,p. 99. (69)Ibid., No. 10,p. 130. (70)I b i d . , No. 11, p. 125. (71) Thornton, D. P.,Jr., Petroleum Processing, 5, 601 (1950). (72)I b i d . , p. 941.
Vol. 43, No. 9
(73)Ibid., 6, p. 146 (1951). (74) Uhl, W. C., Petroleum Processing, 5, 950 (1950). (75) Viland, C.K.,Oil Gas J.,49, No. 30,74 (1950). (76) Viland, C. K.,Petroleum Processing, 5, 830 (1950). (77) Weber, G.,Oil Gus J., 49, No. 1, 53 (1950). (78) I b i d . , No.5,p. 54. (79)I b i d . , No.7,p. 158. (80)Ibid., No. 13,p. 50. (81)Ibid.. No. 17.D. 54. (82j Ibid.; No. 24;-p. 60. (83) Ibid., p. 78. (84) Wilhelm, R. H.,and Valentine, S.. IND.ENQ.CHEM.,43, 1199 (1951). (85) World Petroleum, 22, No. 2,35 (1951). (86) Wright, J. F.,Petroleum Refiner, 29, No. 9, 163 (1950). RECEIYE~D June 13, 1961.
Sulfona tion and Sulfation I
EVERETT
E. GILBERT and E. PAUL JONES,
GENERAL CHEMICAL
DIVISION. LAUREL HILL RESEARCH LABORATORY, LONG ISLAND CITY, .1. Y. Manufacture of industrial sulfonates has continued to expand sharply, particularly of dodecyl benzene sulfonate detergents and sulfonated fatty oils. A n interesting technical improvement in detergent alkylate sulfonation is the proposed use of liquefied petroleum gas as a solvent. Papers continue to appear on the use of pyridine-sulfur trioxide as a special direct sulfonating agent of fairly broad applicability. A noteworthy advance in the aliphatic field is a study on direct sulfonation of aliphatic ketones. The preparation of various heterocyclic sulfonates b y an indirect procedure (aqueous chlorination of the thiol) was shown to b e more widely useful than previously known approaches. Publication of reports on German sulfonation processesup to the present time a rich source of detailed data-has begun to diminish markedly. The relative usefulness of the various compounds of sulfur trioxide in sulfonation and sulfation i s indicated in a tabulation of their frequency of use in the examples and processes cited.
I
N T H E continuation of previous sulfonation reviews by Lisk
(193-226), the present study covers information on this unit process published during 1950,with some reference to earlier work. As previously d e k e d by Lisk (223), the formation of sulfonic acids (sulfur-to-carbon bond) and sulfamic acids (sulfur-tonitrogen bond) is the subject of primary consideration. In addition, however, the subject of sulfation of olefins and alcohols has also been reviewed, because similar reagents are used, and the properties and uses of the industrially important long-chain organic sulfonates and sulfates are generally similar. A separate section is devoted to treatment of fats and oils, as this operation, referred to by the trade as “Sulfonation,” may be either sulfonation or sulfation depending on the conditions and reagent used. Although sulfonation has traditionally been thought of as involving the reaction of an organic compound with an inorganic sulfonating agent, a number of important sulfonated materials are made by reaction of a sulfonic acid with an organic compound t o yield a new sulfonic acid with markedly different properties. This type of reaction is considered under the subheading Polymerieation and Condensation under both Aliphatic and Aromatic headings. Aromatic sulfone formation, often a side reaction in sulfonation proper, is also considered briefly where it occurs a s the result of a direct sulfonation procedure. The manufacture of sulfonates has increased steadily. Production of dodecyl benzene sulfonate detergents doubled in one year, the estimated total for 1950 being over 1 billion pounds (516). The output of sulfonated fats and oils tripled during 1939 to 1947 (332). Sales in 1949 of all sulfonated and sulfated sur-
face active agents totaled 295,000,-
000 pounds, valued a t $66,000.000
(377). In the face of this increase in the volume of sulfonate manufacture, there has developed a shortage in sulfur and sulfuric acid, freight rates have increased, and governmental regulations on waste acid disposal have become more stringent. The result has been an intensification of interest in the possible use of stronger and more efficient sulfonating agents, and in improved methods of spent acid recovery and recycle. This trend was noted by Groggins (227) several years ago. Two general reviews on sulfonation were noted. One of theie (2@), discussed use of sulfur trioxide in its various forms. The second (129)briefly reviewed the manufacture and uses of industrial sulfonates.
ALIPHATIC SULFONATES This general heading covers formation of sulfonic acids in which the -SOsH group is attached to an aliphatic carbon atom regardless of the structure of the rest of the molecule. In most cases this carbon atom in the final sulfonate is saturated; in a few cases it is olefinic. Petroleum sulfonates, which undoubtedly contain some of the aliphatic type, are considered under a separate heading, as are sulfonated fatty oils. DIRECT TREATMENT WITH COMPOUNDS CONTAINING SULFUR TRIOXIDE
Saturated Parafljns and Cycloparafllns. Reaction of methane with sulfur trioxide under pressure (820 to 1350 pounds per square inch) a t elevated temperatures (230’ to 300’ C.)and using mole ratios of sulfur trioxide to methane varying from 0.59 t o 10 with and without mercury sulfate catalyst has been patented (198,169, 327) as a process for producing methanol and sulfonated derivatives of methane [methanesulfonic acid, methanedisulfuric acid (methionic acid), and their methyl esters]
.
