Integrated Production of Gasoline From Light Olefins in a Fluid Cracking Process Plant

This invention relates to a technique for integrating an olefins upgrading process for the catalytic conversion of olefinic feedstocks to liquid hydrocarbons boiling in the gasoline and fuel oil range with the processing and separation of light cracking gases.

Hydrocarbon mixtures containing significant quantities of light olefins are frequently encountered in petrochemical plants and petroleum refineries. Because of the ease with which olefins react, these streams serve as feedstocks in a variety of hydrocarbon conversion processes. Many olefinic conversion processes require that the olefinic feed be provided in a highly purified condition. However, processes which may utilize the olefinic feedstocks without the need for further separation and purification are highly desirable.

Although the main purpose of fluidized catalytic cracking (FCC) is to convert gas oils to compounds of lower molecular weight in the gasoline and middle distillate boiling ranges, significant quantities of C.sub.1 -C.sub.4 hydrocarbons are also produced. These light hydrocarbon gases are rich in olefins which heretofore have made them prime candidates for conversion to gasoline blending stocks by means of polymerization and/or alkylation. Fractionation of the effluent from the fluid catalytic cracking reactor has been employed to effect an initial separation of this stream. The gaseous overhead from the main fractionator is collected and processed in the FCC gas plant. Here the gases are compressed, contacted with a naphtha stream, scrubbed, where necessary, with an amine solution to remove sulfur and then fractionated to provide, for example, light olefins and isobutane for alkylation, light olefins for polymerization, n-butane for gasoline blending and propane for LPG. Light gases are recovered for use as fuel.

Since alkylation units were more costly to build and operate than polymerization units, olefin polymerization was initially favored as the route for providing blending stocks. Increased gasoline demand and rising octane requirements soon favored the use of alkylation because it provided gasoline blending stocks at a higher yield and with a higher octane rating than the comparable polymerized product. However, catalytic alkylation can present some safety and disposal problems. In addition, feedstock purification is required to prevent catalyst contamination and excess catalyst comsumption. Further, sometimes there is insufficient isobutane available in a refinery to permit all the olefins from the FCC to be catalytically alkylated.

Conversion of olefins to gasoline and/or distillate products is disclosed in U.S. Pat. Nos. 3,960,978 and 4,021,502 (Givens, Plank and Rosinski) wherein gaseous olefins in the range of ethylene to pentene, either alone or in admixture with paraffins are converted into an olefinic gasoline blending stock by contacting the olefins with a catalyst bed made up of ZSM-5 or related zeolite. In U.S. Pat. Nos. 4,150,062 and 4,227,992 Garwood et al disclose the operating conditions for the Mobil Olefin to Gasoline/Distillate (MOGD) process for selective conversion of C.sub.3 + olefins.

The phenomena of shape-selective polymerization are discussed by Garwood in ACS Symposium Series No. 218, Intrazeolite Chemistry, "Conversion of C.sub.2 -C.sub.10 to Higher Olefins over Synthetic Zeolite ZSM-5", 1983 American Chemical Society.

In the process for catalytic conversion of olefins to heavier hydrocarbons by catalytic oligomerization using an acid crystalline metallosilicate zeolite, such as ZSM-5 or related shape-selective catalyst, process conditions can be varied to favor the formation of either gasoline or distillate range products. In the gasoline operating mode, or MOG reactor system, oligomerized at elevated temperature and moderate pressure. Under these conditions ethylene conversion rate is greatly increased and lower olefin oligomerization is nearly complete to produce an olefinic gasoline comprising hexene, heptene, octene and other C.sub.6 + hydrocarbons in good yield.

The olefins contained in an FCC gas plant would be an advantageous feed for MOG. U.S. Pat. No. 4,090,949 discloses upgrading olefinic gasoline by conversion thereof in the presence of carbon hydrogen-contributing fragments including olefins and a zeolite catalyst and where the contributing olefins may be obtained from a gas plant. U.S. Pat. Nos. 4,471,147 and 4,504,691 disclose an MOG/D process using an olefinic feedstock derived from FCC effluent. In these two latter patents the first step involves prefractionating the olefinic feedstock to obtain a gaseous stream rich in ethylene and a liquid stream containing C.sub.3 + olefin. While the above patents disclose the general use of olefins obtained from FCC effluent as feedstocks for upgrading conversion, there is not a disclosure of integrating unit operations so as to improve both the oligomerization process and the processing of FCC effluent in a typical FCC gas plant.

U.S. Pat. Nos. 4,012,455 and 4,090,949 (Owen and Venuto) and published European Patent Application No. 0,113,180 (Graven and McGovern) disclose such integration of olefins upgrading with a FCC plant. In the EPA application the olefin feedstock for MOGD comprises the discharge stream from the final stage of the wet gas compressor or the overhead from the high pressure receiver which separates the condensed effluent from the final stage wet gas compressor contained in the gas plant. The present invention improves upon such integrated processes by incorporating olefins upgrading advantageously with the FCC gas plant.

This invention relates to an improvement in the process for upgrading light olefinic crackate gas from hydrocarbon cracking, said light crackate gas containing ethene, propene and other C.sub.1 -C.sub.4 lower aliphatics. This invention provides methods and means for: (a) compressing and cooling the light crackate gas to provide a first pressurized ethene-rich vapor stream and a first condensed crackate stream rich in C.sub.3.sup.+ aliphatics; (b) contacting the first ethene-rich vapor stream under pressure with a C.sub.5.sup.+ liquid sorbent stream in an absorber column under sorption conditions to selectively absorb a major amount of C.sub.3.sup.+ components; (c) recovering a second ethene-rich vapor stream from the absorber column; (d) reacting said second ethene-rich vapor stream in once-through contact with a fluidized bed of acid medium pore zeolite catalyst particles under oligomerization conditions to produce an olefinic hydrocarbon effluent stream rich in C.sub.5.sup.+ hydrocarbons; (e)cooling and separating the reaction effluent stream to provide a light offgas stream and a condensed liquid hydrocarbon product stream; (f) fractionating the liquid hydrocarbon product stream in the absorber column concurrently with sorption of the first ethene-rich vapor stream for recovery of liquid hydrocarbon product with an absorber bottoms liquid stream rich in C.sub.3.sup.+ components; (g) further fractionating the absorber bottoms liquid stream to provide a C.sub.3 -C.sub.4 product and a liquid hydrocarbon fraction consisting essentially of C.sub.5.sup.+ hydrocarbons; and (h) recycling at least a portion of the C.sub.5.sup.+ liquid hydrocarbon fraction to the absorber column as the liquid sorbent stream. Advantageously, the system further comprises the steps of fractionating FCC gas oil crackate in an FCC main fractionation column, contacting a light offgas stream from step (e) with a sponge oil in a secondary sponge absorber to recover residual heavier hydrocarbons, and passing sponge oil sorbate liquid from the secondary absorber to the FCC main fractionation column for recovery.

FIG. 1 is a schematic process diagram of a typical FCC gas plant with an integrated olefins uprgrading unit for fuel gas conversion;and

FIG. 2 is a vertical cross-section view of a preferred fluidized bed reactor system according to the present invention;

FIG. 3 is an alternative embodiment of a typical FCC gas plant with an integrated olefin upgrading unit.

The present invention provides a system for upgrading FCC light olefins to liquid hydrocarbons, utilizing a continuous process for producing fuel products by oligomerizing olefinic components to produce olefinic product for use as fuel or the like. It provides a technique for oligomerizing lower alkene-containing light gas feedstock, optionally containing ethene, propene, butenes or lower alkanes, to produce predominantly C.sub.5.sup.+ hydrocarbons, including olefins.

The preferred feedstock contains C.sub.2 -C.sub.4 alkenes (mono-olefin) , wherein the total C.sub.3 -C.sub.4 alkenes are in the range of about 10 to 50 wt %. Non-deleterious components, such as methane and other paraffins and inert gases, may be present. A particularly useful feedstock is a light gas by-product of FCC gas oil cracking units containing typically 10-40 mol % C.sub.2 -C.sub.4 olefins and 5-35 mol % H.sub.2 with varying amounts of C.sub.1 -C.sub.3 paraffins and inert gas , such as N.sub.2. The process may be tolerant of a wide range of lower alkanes, from 0 to 95%. Preferred feedstocks contain more than 50 wt. % C.sub.1 -C.sub.4 lower aliphatic hydrocarbons, and contain sufficient olefins to provide total olefinic partial pressure of at least 50 kPa. Under the reaction severity conditions employed in the present invention lower alkanes especially propane, may be partially converted to C.sub.4.sup.+ products.

Conversion of lower olefins, especially ethene, propene and butenes, over HZSM-5 is effective at moderately elevated temperatures and pressures. The conversion products are sought as liquid fuels, especially the C.sub.5.sup.+ hydrocarbons. Product distribution for liquid hydrocarbons can be varied by controlling process conditions, such as temperature, pressure and space velocity. Gasoline (eg, C.sub.5 -C.sub.9) is readily formed at elevated temperature (e.g., up to about 510.degree. C.) and moderate pressure from ambient to about 5500 kPa, preferably about 250 to 2900 kPa. Under appropriate conditions of catalyst activity, reaction temperature and space velocity, predominantly olefinic gasoline can be produced in good yield and may be recovered as a product. Operating details for typical olefin oligomerization units are disclosed in U.S. Pat. Nos. 4,456,779; 4,497,968 (Owen et al.) and 4,433,185 (Tabak), incorporated herein by reference.

It has been found that C.sub.2 -C.sub.4 rich olefinic light gas can be upgraded to liquid hydrocarbons rich in olefinic gasoline by catalytic conversion in a turbulent fluidized bed of solid acid zeolite catalyst under low severity reaction conditions in a single pass or with recycle of gaseous effluent components. This technique is particularly useful for upgrading LPG and FCC light gas, which usually contains significant amounts of ethene, propene, butenes, C.sub.2 -C.sub.4 paraffins and hydrogen produced in cracking heavy petroleum oils or the like. It is a primary object of the present invention to provide a novel technique for upgrading such lower olefinic feedstock to distillate and gasoline range hydrocarbons in an economic multistage reactor system.

Recent developments in zeolite technology have provided a group of medium pore siliceous materials having similar pore geometry. Most prominent among these intermediate pore size zeolites is ZSM-5, which is usually synthesized with Bronsted acid active sites by incorporating a tetrahedrally coordinated metal, such as Al, Ga, or Fe, within the zeolytic framework. These medium pore zeolites are favored for acid catalysis; however, the advantages of ZSM-5 structures may be utilized by employing highly siliceous materials or cystalline metallosilicate having one or more tetrahedral species having varying degrees of acidity. ZSM-5 crystalline structure is readily recognized by its X-ray diffraction pattern, which is described in U.S. Pat. No. 3,702,866 (Argauer, et al.), incorporated by reference.

The oligomerization catalyst preferred for use in olefins conversion includes the medium pore (i.e., about 5-7 angstroms) shape selective crystalline aluminosilicate zeolites having a silica to alumina ratio of about 20:1 or greater, a constraint index of about 1-12, and acid cracking activity (alpha value) of about 2-200. Representative of the shape selective zeolites are ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, and ZSM-48. ZSM-5 is disclosed in U.S. Pat. No. 3,702,886 and U.S. Pat. No. Re. 29,948. Other suitable zeolites are disclosed in U.S. Pat. Nos. 3,709,979 (ZSM-11); 3,832,449 (ZSM-12); 4,076,979; 4,076,842 (ZSM-23); 4,016,245 (ZSM-35); and 4,375,573 (ZSM-48). The disclosures of these patents are incorporated herein by reference.

While suitable zeolites having a silica to coordinated metal oxide molar ratio of 20:1 to 200:1 or higher may be used, it is advantageous to employ a standard ZSM-5 having a silica alumina molar ratio of about 25:1 to 70:1, suitably modified. A typical zeolite catalyst component having Bronsted acid sites may consist essentially of aluminosilicate ZSM-5 zeolite with 5 to 95 wt.% silica clay and/or alumina binder.

These siliceous zeolites may be employed in their acid forms ion exchanged or impregnated with one or more suitable metals, such as Ga, Pd, Zn, Ni, Co and/or other metals of Periodic Groups III to VIII. Ni-exchanged or impregnated catalyst is particularly useful in converting ethene under low severity conditions. The zeolite may include other components, generally one or more metals of group IB, IIB, IIIB, VA, VIA or VIIIA of the Periodic Table (IUPAC). Useful hydrogenation-dehydrogenation components include the noble metals of Group VIIIA, especially platinum, but other noble metals, such as palladium, gold, silver, rhenium or rhodium, may also be used. Base metal hydrogenation components may also be used, especially nickel, cobalt, molybdenum, tungsten, copper or zinc. The catalyst materials may include two or more catalytic components, such as a metallic oligomerization component (eg, ionic Ni.sup.+2, and a shape-selective medium pore acidic oligomerization catalyst, such as ZSM-5 zeolite) which components may be present in admixture or combined in a unitary bifunctional solid particle. It is possible to utilize an ethene dimerization metal or oligomerization agent to effectively convert feedstock ethene in a continuous reaction zone. Certain of the ZSM-5 type medium pore shape selective catalysts are sometimes known as pentasils. In addition to the preferred aluminosilicates, the borosilicate, ferrosilicate and "silicalite" materials may be employed.

ZSM-5 type pentasil zeolites are particularly useful in the process because of their regenerability, long life and stability under the extreme conditions of operation. Usually the zeolite crystals have a crystal size from about 0.01 to over 2 microns or more, with 0.02-1 micron being preferred.

A further useful catalyst is a medium pore shape selective crystalline aluminosilicate zeolite as described above containing at least one Group VIII metal, for example Ni-ZSM-5. This catalyst has been shown to convert ethylene at moderate temperatures and is disclosed in a copending (allowed) U.S. Pat. Application Ser. No. 893,522, filed 4 August 1986 by Garwood et al, incorporated herein by reference.

A typical system for integrating MOG into an FCC gas plant is shown in FIG. 1. The present invention contemplates integrating gasoline mode (MOG) olefin upgrading into a FCC gas plant. Referring to FIG. 1, the condensed overhead from the FCC main fractionator 2 flows through indirect heat exchanger 3 via line 4 and into low pressure separator 6 for separation into a gaseous phase and a liquid phase. The gaseous portion from separator 6 flows through line 8 into the suction of first stage wet gas compressor 9 for the initial increase in pressure. The wet gas discharges from first stage compressor 9, is condensed and discharges into intermediate pressure separator 14. Gases from separator 14 are directed through second stage wet gas compressor 10 from which the effluent is discharged through line 16 where again the gases are condensed in heat exchanger 17 and directed into high pressure separator 18. The purpose of the gas plant is to maximize liquid recovery. Thus, any C.sub.3 and C.sub.4 hydrocarbons in the gas plant which are recovered as LPG are more valuable than the C.sub.1 and C.sub.2 fuel gas. Thus, the final high pressure gas from high pressure separator 18 is directed via line 20 to primary absorber 22. In primary absorber 22, C.sub.5 + liquids pass in countercurrent flow to the high pressure gas to absorb heavy hydrocarbons including C.sub.3 and C.sub.4 hydrocarbons from the gas stream. The C.sub.5 + liquids employed include the liquid phase from the FCC low pressure separator 6 via line 21 as well as a portion of the final liquid product from the gas plant shown entering absorber 22 via line 48.

The unabsorbed gases pass from the top of absorber 22 through line 24 where the gases may be optionally scrubbed with diethanolamine or other suitable solvent in scrubber means 26 to reduce the acid gases, such as H.sub.2 S and the like, to acceptable levels. The scrubber may be placed prior to absorber 22. The gases pass from line 24 then optionally through scrubber 26 and into line 50, then via heat exchanger 78 and furnace 73 to oligomerization reactor 74. Effluent 76 is cooled in exchanger 78 to partially condense liquid, and then is passed to separator 80. Gas leaving separator 80 via line 82 is compressed in compressor 84. A portion of this gas can be recycled to the feed via line 86. The liquid which has dropped out in separator 80 is pumped via line 88 and gas from compressor 84 which is not recycled are combined and directed via line 90 to sponge absorber 30 where the gases are contacted in countercurrent flow with sponge oil, a liquid sorbent stream which may be stripped heavy naphtha or light fuel oil boiling in the 350.degree.-500.degree. F. range. In sponge absorber 30, the C.sub.3 + gases are absorbed by the sponge oil fed from fractionator 2, then line 31. Subsequently, after absorption the sorbate stream passes from sponge absorber 30 through line 32 into the FCC main fractionating column 2. The unabsorbed C.sub.2 - gases pass from absorber 30 through line 34 and are recovered and may be eventually burned as fuel gas.

Regarding the liquid phase in high pressure separator 18, this liquid passes from the receiver through line 36 to stripper 38 where steam is employed as a reboiler heat source to remove the light gases from this stream. The light gases pass from the top of stripper 38 through line 40 and discharge back into high pressure separator 18 from which the useful light gases are recovered.

The stripped C.sub.3 + liquid passes from stripper 38 through line 39 to debutanizer 42 where a C.sub.4 - fraction is separated and passes from the column as overhead through line 44 where it is recovered as LPG product. Gasoline and/or fuel oil fraction (C.sub.5 +) is removed from debutanizer 42 as the bottoms fraction through line 46. A portion of this fraction is recycled through line 48 to primary absorber 22 as a portion of the absorbing liquid as described previously. The remaining portion of the C.sub.5 + bottoms is recovered as product and can be employed as blending stock for gasoline and/or fuel oil following further fractionation as required.

The process operating technique provides for maintaining in a low severity continuous reaction zone 74, a fluidized bed of zeolite catalyst particles in a turbulent reactor bed, usually at a temperature of about 260.degree. to 510.degree. C. Hot feedstock vapor can be passed upwardly through the fluidized catalyst bed in a single pass at reaction severity conditions sufficient to convert feedstock alkenes predominantly to intermediate range olefins in the C5-C9 range without substantial formation of aromatics. This can be achieved by maintaining turbulent fluidized bed conditions through the reactor bed at a superficial fluid velocity of about 0.3 to 2 meters per second. The reactor effluent contains a major amount of C.sub.5.sup.+ hydrocarbons and a minor amount of C4- hydrocarbons, including pentane and pentene in a weight ratio of about 0.1:1 to 7:1. Substantially all C4-light gas components are removed from the reactor effluent stream to provide an intermediate hydrocarboon stream comprising a major amount of intermediate C.sub.5.sup.+ olefins.

The stream which enters reactor 74 is rich in all of the FCC olefins. A typical composition of this stream is given in Table 1.

Conditions in the MOG reactor can vary within the limits previously described to form liquid hydrocarbon but most preferably will be such so as to maximize production of a gasoline range hydrocarbon liquid.

Particle size distribution can be a significant factor in achieving overall homogeneity in turbulent regime fluidization. It is desired to operate the process with particles that will mix well throughout the bed. Large particles having a particle size greater than 250 microns should be avoided, and it is advantageous to employ a particle size range consisting essentially of 1 to 150 microns. Average particle size is uually about 20 to 100 microns, preferably 40 to 80 microns. Particle distribution may be enhanced by having a mixture of larger and smaller particles within the operative range, and it is particularly desirable to have a significant amount of fines. Close control of distribuiion can be maintained to keep about 10 to 25 wt % of the total catalyst in the reaction zone in the size range less than 32 microns. This class of fluidizable particles is classified as Geldart Group A. Accordingly, the fluidization regime is controlled to assure operation between the transition velocity and transport velocity. Fluidization conditions are substantially different from those found in non-turbulent dense beds or transport beds.

The oligomerization reaction severity conditions can be controlled to optimize yield of C.sub.5 -C.sub.9 aliphatic hydrocarbons. It is understood that aromatic and light paraffin production is promoted by those zeolite catalysts having a high concentration of Bronsted acid reaction sites. Accordingly, an important criterion is selecting and maintaining catalyst inventory to provide either fresh catalyst having acid activity or by controlling catalyst deactivation and regeneration rates to provide an average alpha value of about 2 to 50, based on total catalyst solids.

Reaction temperatures and contact time are also significant factors in determining the reaction severity, and the process parameters are followed to give a substantially steady state condition wherein the reaction severity index (R.I.) is maintained within the limits which yield a desired weight ratio of alkane to alkene produced in the reaction zone. This index may vary from about 0.1 to 7:1, in the substantial absence of C3+ alkanes; but, it is preferred to operate the steady state fluidized bed unit to hold the R.I. at about 0.2 to 5:1. While reaction severity is advantageously determined by the weight ratio of propane:propene (R.I..sub.3) in the gaseous phase, it may also be measured by the analogous ratios of butanes:butenes, pentanes:pentenes (R.I..sub.5), or the average of total reactor effluent alkanes:alkenes in the C.sub.3 -C.sub.5 range. Accordingly, the product C5 ratio may be a preferred measure of reaction severity conditions, especially with mixed aliphatic feedstock containing C.sub.3 -C.sub.4 alkanes.

This technique is particularly useful for operation with a fluidized catalytic cracking (FCC) unit to increase overall production of liquid product in fuel gas limited petroleum refineries. Light olefins and some of the light paraffins, such as those in FCC light gas, can be converted to valuable C.sub.5.sup.+ hydrocarbon product in a fluid-bed reactor containing a zeolite catalyst. In addition to C.sub.2 -C.sub.4 olefin upgrading, the load to the refinery fuel gas plant is decreased considerably.

The use of fluidized bed catalysis permits the conversion system to be operated at low pressure drop. Another important advantage is the close temperature control that is made possible by turbulent regime operation, wherein the uniformity of conversion temperature can be maintained within close tolerances, often less than 10.degree. C. Except for a small zone adjacent the bottom gas inlet, the midpoint measurement is representative of the entire bed, due to the thorough mixing achieved.

In a typical process, the olefinic feedstock is converted in a catalytic reactor under oligomerization conditions and moderate pressure (ie-400 to 2500 kPa) to produce a predominantly liquid product consisting essentially of C.sub.5.sup.+ hydrocarbons rich in gasoline-range olefins and essentially free of aromatics.

Referring now to FIG. 2, feed gas rich in lower olefins passes under pressure through conduit 210, with the main flow being directed through the bottom inlet of reactor vessel 220 for distribution through grid plate 222 into the fluidization zone 224. Here the feed gas contacts the turbulent bed of finely divided catalyst particles. Reactor vessel 210 is shown provided with heat exchange tubes 226, which may be arranged as several separate heat exchange tube bundles so that temperature control can be separately exercised over different portions of the fluid catalyst bed. The bottoms of the tubes are spaced above feed distributor grid 222 sufficiently to be free of jet action by the charged feed through the small diameter holes in the grid. Alternatively, reaction heat can be partially or completely removed by using cold feed. Baffles may be added to control radial and axial mixing. Although depicted without baffles, the vertical reaction zone can contain open end tubes above the grid for maintaining hydraulic constraints, as disclosed in U.S. Pat. No. 4,251,484 (Daviduk and Haddad). Heat released from the reaction can be controlled by adjusting feed temperature in a known manner.

Catalyst outlet means 228 is provided for withdrawing catalyst from above bed 224 and passed for catalyst regeneration in vessel 230 via control valve 229. The partially deactivated catalyst is oxididatively regenerated by controlled contact with air or other regeneration gas at elevated temperature in a fluidized regeneration zone to remove carbonaceous deposits and estore acid acitivity. The catalyst particles are entrained in a lift gas and transported via riser tube 232 to a top portion of vessel 230. Air is distributed at the bottom of the bed to effect fluidization, with oxidation byproducts being carried out of the regeneration zone through cyclone separator 234, which returns any entrained solids to the bed. Flue gas is withdrawn via top conduit 236 for disposal; however, a portion of the flue gas may be recirculated via heat exchanger 238, separator 240, and compressor 242 for return to the vessel with fresh oxidation gas via line 244 and as lift gas for the catalyst in riser 232.

Regenerated catalyst is passed to the main reactor 220 through conduit 246 provided with flow control valve 248. The regenerated catalyst may be lifted to the catalyst bed with pressurized feed gas through catalyst return riser conduit 250. Since the amount of regenerated catalyst passed to the reactor is relatively small, the temperature of the regenerated catalyst does not upset the temperature constraints of the reactor operations in significant amount. A series of sequentially connected cyclone separators 252, 254 are provided with diplegs 252A, 254A to return any entrained catalyst fines to the lower bed. These separators are positioned in an upper portion of the reactor vessel comprising dispersed catalyst phase 224. Filters, such as sintered metal plate filters, can be used alone or in conjunction with cyclones.

The product effluent separated from catalyst particles in the cyclone separating system is then withdrawn from the reactor vessel 220 through top gas outlet means 256. The recovered hydrocarbon product comprising C.sub.5.sup.+ olefins and/or aromatics, paraffins and naphthenes is thereafter processed as required to provide a desired gasoline or higher boiling product.

Under optimized process conditions the turbulent bed has a superficial vapor velocity of about 0.3 to 2 meters per second (m/sec). At higher velocities entrainment of fine particles may become excessive and beyond about 3 m/sec the entire bed may be transported out of the reaction zone. At lower velocities, the formation of large bubbles or gas voids can be detrimental to conversion. Even fine particles cannot be maintained effectively in a turbulent bed below about 0.1 m/sec.

A convenient measure of turbulent fluidization is the bed density. A typical turbulent bed has an operating density of about 100 to 500 kg/m3, preferrably about 300 to 500 kg/m.sup.3, measured at the bottom of the reaction zone, becoming less dense toward the top of the reaction zone, due to pressure drop and particle size differentiation. The weight hourly space velocity and uniform contact provides a close control of contact time between vapor and solid phases, typically about 3 to 15 seconds.

Several useful parameters contribute to fluidization in the turbulent regime in accordance with the process of the present invention. When employing a ZSM-5 type zeolite catalyst in fine powder form such a catalyst should comprise the zeolite suitably bound or impregnated on a suitable support with a solid density (weight of a representative individual particle divided by its apparent "outside" volume) in the range from 0.6-2 g/cc, preferably 0.9-1.6 g/cc. The catalyst particles can be in a wide range of particle sizes up to about 250 microns, with an average particle size between about 20 and 100 microns, preferably in the range of 10-150 microns and with the average particle size between 40 and 80 microns. When these solid particles are placed

in a fluidized bed where the superficial luid velocity is 0.3-2, operation in the turbulent regime is obtained. The velocity specified here is for an operation at a total reactor pressure of about 400 to 2500 kPa. Those skilled in the art will appreciate that at higher pressures, a lower gas velocity may be employed to ensure operation in the turbulent fluidization regime. The reactor can assume any technically feasible configuration, but several important criteria should be considered. The bed of catalyst in the reactor can be at least about 5-20 meters in height, preferably about 9 meters.

The following example tabulates typical FCC light gas oligomerization reactor feed and effluent compositions and shows process conditions for a particular case in which the reactor temperature is controlled at 400.degree. C. The reactor may be heat balanced by controlled preheating the feed to about 135.degree. C. The preferred catalyst is H-ZSM-5 (25 wt %) with particle distribution as described above for turbulent bed operation.

Process integration can be adapted to employ certain features of an unsaturated gas plant (USGP), especially multistage compression, phase separation, distillation absorption and the operatively connected unit operations essential to recovery of light cracking products or similar aliphatic hydrocarbon streams. In one embodiment, an integrated fluidized bed reactor is maintained in steady state operation at appropriate feed rate, temperature, pressure and catalyst activity to effect the desired oligomerization of lower olefinic components in the feedstock to gasoline range hydrocarbons.

The alternative embodiment depicted in FIG. 3 provides operating techniques and processing equipment for integrating the light FCC crackate recovery with olefins upgrading in a fluidized bed system. Interstage fractionation may be adapted to utilize conventional petroleum refinery cracking plant equipment in a novel process for upgrading light olefinic crackate gas from hydrocarbon cracking. The light crackate gas containing ethene propene and other C.sub.1 -C.sub.4 lower aliphatics is passed from the FCC main column 310 via cooler 312 and overhead accumulator 314 to means 316 for compressing and cooling the light crackate gas to provide a first pressurized ethene-rich vapor stream 318 and a first condensed crackate stream 320 rich in C.sub.3.sup.+ aliphatics. Absorber tower 330 provides means contacting the first ethene-rich vapor stream under pressure with a C.sub.5.sup.+ liquid sorbent stream 346 in the absorber column under sorption conditions to selectively absorb a major amount of C.sub.3.sup.+ components introduced via gas stream 318 and liquid stream 320, thus recovering a second ethene-rich vapor stream 334 from the absorber de-ethanaizer column. The C3+ liquid bottoms stream 336 may be further fractionated in a debutanizer tower 340 to provide a C5+ liquid gasoline product 342 and LPG product 344. Optionally, the pressurized FCC light gas stream may be contacted with amine in absorber tower 338 to remove any H.sub.2 S. The ethylenic gas is then upgraded in reactor means 350 by reacting the second ethene-rich vapor stream in once-through contact with a fluidized bed of acid medium pore zeolite catalyst particles under oligomerization conditions to produce an olefinic hydrocarbon effluent stream rich in C.sub.5.sup.+ hydrocarbons. Preferably, this is a fluid bed reactor as depicted in FIG. 2 and described above

As part of the reactor effluent recovery system, means are provided for cooling and separating the reaction effluent stream to provide a light offgas stream and a condensed liquid hydrocarbon product stream. Advantageously, this is achieved by cooler means 354 and phase separator means 356. Recovery of a wild gasoline liquid stream 332 containing normally liquid components and volatile C3-C4 components permits recycle of this stream to provide for fractionating the liquid hydrocarbon product stream in the absorber column concurrently with sorption of the first ethene-rich vapor stream for recovery of liquid hydrocarbon product with the absorber bottoms liquid stream 336 rich in C.sub.3.sup.+ components.

By further fractionating the absorber bottoms liquid stream to provide a C.sub.3 -C.sub.4 product and a liquid hydrocarbon fraction consisting essentially of C.sub.5.sup.+ hydrocarbons, and recycling at least a portion of the C.sub.5.sup.+ liquid hydrocarbon fraction via conduit 346 to the upper stages of absorber column 330 as the liquid sorbent stream absorber efficiency is enhanced.

The process is particularly useful for fractionating FCC gas oil crackate in an FCC main fractionation column in combination with sponge absorber 360. This is achieved by contacting light offgas stream 358 from accumulator 356 with a sponge oil in the secondary sponge absorber 360 to recover residual heavier hydrocarbons. This can be further integrated by passing sponge oil sorbate liquid from the secondary absorber to the FCC main fractionation column 310 for recovery. The above described integration technique is particularly useful where the condensed liquid hydrocarbon stream 332 contains volatile components and passes into the absorber column at an upper portion thereof to provide additional sorbent liquid. The off gas from the sponge absorber 360 can than be optionally passed to a secondary amine scrubber 370 for further desulfurization and remove any H.sub.2 S which might be containined in the sponge oil sorbent stream.

While the invention has been shown by describing preferred embodiments of the process, there is no intent to limit the inventive concept, except as set forth in the following claims.