SINGLE ZONE OLIGOMERIZATION OF LOWER OLEFINS TO DISTILLATE UNDER LOW SEVERITY IN A FLUID BED WITH TAILORED ACTIVITY

Abstract: A turbulent, dense fluid-bed of a medium pore zeolite metallosilicate catalyst, the bed having a tailored distribution of weight fractions of particles having activities which fall generally along a predetermined curve of activity distribution, is operated under low severity conditions at low WHSV, under moderate pressure and moderate severity to obtain surprisingly high conversion to C5+ hydrocarbons having an olefin content in excess of about 80 %, about one third by weight of which is distillate (C9+). Only a small amount of cycloolefins and a negligibly small amount of aromatics and paraffins are produced. The distillate is formed with a single pass yield in the range from about 10 - 40 weight percent. The reactor is operated only in the distillate mode. The dense fluid bed has a density, measured at the bottom of the reaction zone, less than 640 kg/m<3> (40 lb/ft<3>), and preferably operates at a temperature in the range from about 250 DEG C (482 DEG F) to about 371 DEG C (700 DEG F), and a pressure in the range from above about 200 kPa (15 psig) but below about 2800 kPa (400 psig). Despite operating under ''easy'' conditions, this unique process provides per pass conversion of olefins, and selectivity to C5+ hydrocarbons, each in excess of 80 %.

Description: SINGLE ZONE OLIGOMERIZATION OF LOVER OLEFINS TO ISTILLATE UNDER LOW SEVERITY IN A FLUID BED WITH TAILORED ACTIVITY This invention relates to a fluid-bed catalytic process for upgrading olefinic light gas feedstock (termed "light gas" for brevity herein) containing lower, particularly C,-Cr, olefins (alkenes) and paraffins (alkanes) . The olefins are converted to distillate in a single-zone fluid-bed reactor, operating at moderate pressure and temperature, ("low severity" conditions). These operating conditions are referred to as a low severity MDD (for "Mobil Olefin to Distillate") process. By "distillate" we refer to CQ+ hydrocarbons boiling in the range from about 130°C to about 300° C (266°F - 572°F). We know of no other process which can convert olefins with at least 80. conversion and about 20. by weight (wt .) single pass yield, to distillate, in a single zone. More particularly, the invention provides a continuous process for oligomerizing light gas containing a C,-C 5 olefin, namely propone, butenes and pentenes, a mixture thereof, or mixtures thereof with a minor proportion by weight of paraffins, preferably only C, -Cr paraffins, in the absence of added hydrogen, to a Cg+ rich hydrocarbon stream, in a highly efficient operating mode because of a fluid catalyst bed with "tailored" activity. By "C g + rich" we refer to the presence of at least 10. by weight of Cg+ hydrocarbons in the product. By "tailored" activity we refer to a bed which consists ^ essentially of weight fractions, in a specified range, of equilibrated low activity catalyst particles, the weight fraction in each range having a specified activity (alpha) in defined ranges below 10, and less than 10 wt _ of the catalyst has an activity greater than 10. The alpha value for catalyst activity is defined by the specific test described in U.S. Patents Nos. 3,827,968 aiϊi 3,960,978 to Givens et al. Developments in fluid-bed catalytic processes using a wide variety of zeolite catalysts have ^purred interest in commerciali zing the conversion of olefinic feedstocks for produci ng C-.+ gasoline, diesel fuel, etc. In addition to the discovery that the intrinsic oligomerization reactions are promoted by ZSM-5 type zeolite catalysts, several discove ries relati ng to implementing the reactions in an apt reactor envi ronment, have contributed to the commercial success of current industrial processes . These are environmentally acceptable processes or utilizing feedstocks that contain lower olefins, especially C,-Cr alkenes, though some ethene (ethylene) , and some olefins and paraffins heavier than Cr may also be present. Of particular interest is that the ZSM-5 type catalyst used under our "easy" process conditions, and also as used under the severe process conditions of U .S . Patent 4 ,746 ,762, does not appear to suffer from a sensiti vity (poisoning) to basic nitrogen-containing organic coπpounds such as alk lamines (e. g. diethylamine) , or, to oxygenated coπpounds such as ketones, a proclivity which is characteristic of the catalyst under the process conditions of prior art olefin oligomeri zation processes, particularly the fixed bed processes operated at ve ry high pressure. Such processes requi re the addition of hydrogen as a pre vent itive antidote. It will be recognized that alkylamines are used in treating light gas streams, and ketones are typically present in Fische r Tropsc h-de ri v d light ends streams, both of whi ch streams are particularly well-suited for upgradi iξ by oligomerization. Though our process is not adve rsely affected by the presence of hydrogen, there is no readily discernible economic incentive or using it in our single stage reactor, and we prefer not to do so. Prior art moderate pressure processes using a zeolite catalyst to oligomeri ze lower olefins under comparable teπperature conditions produced excellent conve rsions to di stillate range olefins in a ixed bed microreactor but neglected to state what the alpha value of their catalyst was ; nor did they suggest that the alpha value might be of over- riding signi icance (see 'Conversion of C 2~ C 10 01e£ ins to Higher Olefins Over Synthetic Zeolite ZSM-5" by E.W. Garwood presented at the Symposium on Advances in Zeolite Chemistry before the Division of Petroleum Chemistry, Ir ., American Chemical Society, Las Vegas Meeting March 28 - April 2, 1962). When the low severity process is carried out in a fluid bed using a relatively "high alpha" catalyst (above 100), the result is formation of a mixture of aromatics, naphthenes and paraffins, and a minor proportion by wt of olefins. Even when our process is carried out in a fluid bed with a relatively uniform low alpha in the range from 1 to 5, that is, none of the catalyst has an alpha greater than 5, neither the high conversion to Cr+ olefins, nor the high selectivity to C Q + is obtained. Conversion of C,-Cr alkenes and alkanes to produce aromatics- rich liquid hydrocarbon products were found by Cattanach (U.S. Patent No. 3,760,024) and Yan et al (U.S. Patent No. 3,845,150) to be effective processes using the ZSM-S type zeolite catalysts. In U.S. Patents 3,960,978 and 4,021,502, Plank, Rosinski and Givens disclose conversion of C^-Cς olefins, alone or in admixture with paraffinic components, into higher hydrocarbons over crystalline zeolites having controlled acidity; so do Dwyer et al in U.S. Patent NO. 3,700,724. Garwood et al have also contributed to the understanding of catalytic olefin upgrading techniques and have contributed improved processes as in U.S. Patents 4,150,062, 4,211,640 and 4,227,992. The '062 patent discloses conversion of olefins to gasoline or distillate in the range from 190-315°C and 42-70 atm. The '978 patent discloses that low alpha Z-3M-5 and ZJ3 -11 catalysts not only have reduced activity for cracking n-hexane and other paraffins, but also produce less than 10% by wt aromatics. The runs were made in a fixed bed micro reactor, and, at that time, it was not known that the process was not operable on a larger scale without the addition of hydrogen to control coke deposition and to prevent poisoning of the catalyst by nitrogen- containing organic impurities. The basic knowledge that low activity ZSM-5 and ZSM-11 type catalysts effectively oligomerized lower olefins was used to arrive at improvements in "Conversion of LPG Hydrocarbons to Distillate Fuels or Lubes Using Integration of LPG Dehyd rogenation a d MOGH" in U.S. Patent No. 4.542, 247 to Chang et al which discloses fixed beds in a two-stage catalytic process for converting olefins to gasoline and distillate; and, more recently, in "Catalytic Conversion of Olef ins to Higher Hydrocarbons" in U.S . Patent No. 4 ,456 ,779 to Owen et al. which discloses oligome rizat ion of olefins in three down-f low fixed beds, in series, with intercoolers. Both f ixed-bed processes requi re the addition of hydrogen for the reasons given here inabove. ? The low severity conditions under which our process produces distillate from lower olef ins is particularly unexpected because it was generally accepted that relatively high severity was essential to obtain conve rsions of economic signif icance. Never has there been any suggestion that a tailored weight fract ion dist ribution of eoμilibrated low activity catalyst particles might have a dominating effect on the oligome rization of a lower olefin feed if the process conditions were aptly ctosen. The prior art processes relate to the conversion of lower olef ins, especially propene and butenes, over ZSM-5 and HZS -5 , at moderately elevated temperatures and pressures. The sought-after conve rsion products are liquid fuels, especially the C>-+ aliphatic and aromatic hydrocarbons. It is known that the product dist ribution may be tailored by controlling process conditions, such as temperature, pressure and space velocity. Gasoline C^-C. Q is readily formed at elevated temperature (preferably about 400°C) and pressure from ambient to about 2900 kPa (420 psia) , preferably about 250 to 1450 kPa (36 to 210 psia) . Olefinic gasoline can be produced in good yield and may be recovered as a product or fed to a low severity, high pressure reactor system for fu rther conversion to heavier distillate-range products. Distillate mode operation can be employed to maximize production of C Q + aliphatics by reacting the lower and intermediate olefins at high pressure and moderate teπperature. Operating details for typical "M3GD' ( for Mobil Olef in to Gasoline ξ, Distillate) oligoπErization units are disclosed in U.S. Patents Nos. 4,456,779 and 4,497,968 (Owen et at); 4,433,185 (Tabak); and U.S. Patent No. 4,746,762. We have now found that C,-C^-rich and higher olefins may be selectively upgraded to normally liquid hydrocarbons in the distillate range by catalytic conversion in a turbulent fluidized bed of solid acid ZSM-5 type of zeolite catalyst at least 90 wt % of which has an equilibrated alpha less than 10, and the remainder an alpha above 10, most preferably operating at a pressure below 1480 kPa (200 psig) in the dense phase, in the absence of added hydrogen, in a single pass, or with recycle of undesired oligoιτerized product. Sxxzh operation results in the most important advantage of our MOD process, namely the use of a relatively low cost, low pressure reactor in which close temperature control is a olded by operation of a fluid-bed in the turbulent regime (referred to as a "turbulent bed"). An essentially uniform conversion temperature may be maintained (often with closer than +5°C tolerance) to ensure that the distribution of activity of the catalyst weight fractions lies generally along a predetermined curve. Except for a small zone adjacent the bottom gas inlet, the midpoint measurement of conditions in the bed is representative of the entire bed, due to the thorough mixing achieved. Moreover, in turbulent beds, fluidization is better at a higher fluidizing gas velocity, and with a higher level of the finer sizes of catalyst (see R.M. Braca and A. A. Fried, in Fluidization, D.F. Othmer, Ed. (Reinhold, New York, 1956), pp. 117-138; W.W. Kraft, W. Ulrich, W. O'Connor, ibid., pp. 184-211). This requires a significant amount of fines, from about 10 to 25 % by weight (% by wt) having a particle size less than 32 . Since it is difficult to control the distribution and activity of catalyst fines it would seem at cross- purposes deliberately to require operation with a predetermined weight fraction of fines. U.S. Patent Nos. 4,417,086 and 4,417,087 to Miller teach a two-zone reactor e rating in the transport mode where the relative sφerf icial gas velocity is greater than the te rminal velocity in free fall. The re is no suggestion that the activity of the bed be tailored by φecifying weight fractions of catalyst particles, all having relatively low equilibrated activity. Though the operation of a fluid-bed is illustrated (exaπple 2 in each of the '086 and r 087 patents) note that no operating pressure is stated in the fo rmer, and that φe rating pressure in the latter is 170 kPa (10 psig, 24.7 psia) . The general disclosure that the processes may be φe rated at a pressure in the range from subatmo spheric to several hund red atmospheres , but preferably 10 bar or less, and most preferably 0 to 6 bar, (see middle of col 6 in '086 , and, near top of col 5 in ' 087) is not so ingenuous as to be meant to apply equally to the fixed bed (example 1 of '086 and '087 , each illustrates 34.5 bar, 500 psi) and the 170 kPa fluid-bed. It has been di scove red that a turbulent dense fluid-bed of a medium pore zeolite metallosilicate catalyst having a low ecμilib rated average activity (alpha) in the range from 1 to 10 as a result of a particular dist ribution of alphas in preselected weight fractions of the bed which φe rates at low WHSV (it bei ng understood that WHSV signif ies kilograms of feed per kilograms of zeolite per hour) ,- under "easy" conditions of pressu re and temperatu re, will neve rtheless oligomerize a "light gas" feed containing C^-C, olefins (predominantly C,-C_ olefins, the remainder being -_~-- lower alkanes ) , to C g + dist illate range hydrocarbons, a minor proportion of C Q - aliphatic and na thenic gasoline components, and essentially no aromatics. By "essentially no aromatics" we refer to a negligible amount, less than about 5% , typically less than 2% by wt, of the product. In our MOD proces s, light gas containing a substantial, preferably a major portion, typically more than 75 % of combined prφene and butenes, is a particularly well-suited feed to the fluid-bed reactor. The presence of paraf fins is not deleterious, but under the conditions of our process, they are left essentially unreacted and are the refore kept to a minimum in the feedst eam. The conversion of eth/lene, if present, is only about 50% at 315°C (600°F) and is therefore also kept to a minimum. The WD reactor is φe rated at relatively low pressure in the range from about 200 kPa to about 2860 kPa (15 psig - 400 psig) , and moderate temperature in the rarge from 250°C to about 371°C (482°F - 700°F) to produce a relatively large fraction (about 20% by wt, or more) of distillate which is at least 50% by wt olefins, in a single pass. A minor amount of Cr+, mainly Cg-, gasoline range hydrocarbons may also be formed. When such Cr-C 8 gasoline range olefins are formed, they are typically not a desired product in our process, and are therefore separated from the product. The sφa rated olefins may be used as reactants to produce any of numerous products in known processes, or the olefins may be blended into gasoline, or, recycled in our process to yield desired distillate, or, reactive tertiary olefins are converted into ethers, or, lower and intermediate olefins may be reacted at high pressure and moderate teπperature to form C 1Q + aliphatic s which are valuable products in an oil refinery. More specifically, to produce desirable oligomerized product effectively, it has been discovered that when the dense phase fluid bed, referred to hereinabove, is operated at siperatmo spheric pressure, preferably above about 270 kPa (25 psig) but below 1480 kPa_*(200 psig) at a WHSV in the range from about 0.05 hr to about 5 hr with a catalyst fines content in the range from about 10% to 20%, the product contains more than 60% by wt, typically more than 80% by wt of olefins, the remainder being mostly cyclo-olefins, unreacted etlylene, paraffins and cyclφa raff ins, and less than 2% by wt aromatics. This invention provides a process for oligomerizing an ole in-containing feedstream consisting essentially of "light gas" containing predominantly C,=+ olefins, to distillate (Cg+ hydrocarbons), coπprising, operating a dense phase fluid-bed in the turbulent regime at moderate pressure and moderate teπperature conditions while flowing the light gas feed throigh tie bed which consists essentially of a finely di vided medium po re zeolite metallosilicate catalyst having a const raint index in the range from 1 to 12 and consisting essentially of specif ied weight fractions of equilibrated low activity catalyst particles , each weight fraction having an alpha in a preselected range from 1 to below 10 , these we ight fractions totalling at least 90% by wt, tie remaining weight fractions totalling less than 10% by wt of the catalyst havi ng an alpha greate r than 10% ; and, to provide a low seve rity M3D reactor which in combination with a regenerator, is pa rticularly adapted to carry out the foregoing process. More speci ically, this invention provi des a process for the continuous convers ion of a lower olefin feedstock gas to distillate, in which process the feedstock is contacted with a fluidized bed of zeolite catalyst under conversion condit ions, the process coπprising , equilibrating the fluid bed of catalyst which consists essentially of finely di vided medium po re zeolite metallosilicate particles having a const raint index in the range from 1 to 12 , the bed having an average acti vity (alpha) in the range from 2 to 10 and consisting essentially of preselected weight fractions of low acti vity catalyst pa rticles totalling at least 90% by wt havi ng an alpha no greater than 10 , the ream ining " less than 10% by wt of the catalyst particles having an alpha greater than 10% ; passing the feedstock gas φwardly through the fluidized bed in a vertical reactor column having a turbulent react ion zone, while maintaining a superficial velocity in the range from 0 .3 to 2 /sec so that catalyst particles in which the silica : alumina molar ratio is in the range from 20: 1 to 200: 1 are held in a turbulent regime; maintaining an ave rage density, measured at the bottom of the fluidized bed in the raηge from 264 to 608 kg/m , at a pressure in the range from 200 kPa to 2800 kPa and a teπ er ture in the rarge from 250°C to 371°C; withdrawing a portion of coked catalyst from the reaction zone, oxidatively regenerating the withdrawn catalyst and returning regenerated catalyst to the fluidized bed at a rate sufficient to maintain a reaction severity index, expressed as the propane : prop ene weight ratio in the hydrocarbon product, in the range from 0.2:1 to 5:1, whereby at least 10% by wt, and preferably more than 20% by wt of the feedstock is converted to a substantially olefinic C g =+ distillate. This invention provides an economical process for directly converting lower olefins into distillate in a single pass through a turblent fluid bed, with the amotnt of distillate produced ranging from 20% to 50% by wt of the C- product, the remaining product beirg C 8 - gasoline rarge hydrocarborns. The remaining C 8 - gasoline rarge olefins, typically not desired in our process, include tertiary olefins. The undesi ed Cg- olefins are separated for further use. They may be (i) blended into gasoline, (ii) recylcled to tfe MDD reactor to produce additional distillate, or (iii) reactive tertiary olefins may be converted into ethers, and (iv) lower and intermediate olefins may be reacted at high pressure and moderate teπperature to form C,Q+ aliphatics. In the drawings, Figure 1, is a curve graphically illustrating a typical distribution of activity (alpha) as a function of the relative weight fraction of catalyst particles in the fluid bed of the reaction vessel. Figure 2 is a phase diagram showing a plot of dew-point and bubble-point curves for a representatie distillate produced in our process. Figure 3 is a flowsheet schematically illustrating the use of a relatively low pressure fluid bed reactor in association with a regenerator in which coked φ catalyst is oxidatively regenerated and returned to the reactor, either periodically, or continuously. The process of this invention is specifically directed to producing an essentially olefinic distillate from a predominantly olefinic light gas feedstock under moderate severity and relatively low pressure conditions. The process is carried out in a fluid-bed of catalyst under process condit ions at which the fluidi zi ng medium is neve r at, or below its dewpoint. The ref ore, if operation at a pressu re at the top of the rarge is sought, the teπperature at which the fluid-bed operates is raised so that the mixture of hyd rocarbons in the fluid-bed is above its dew-point. In the preferred embodiment, the process of this invention uses an olefinic light gas feedstock (commonly referred to as olefinic LPG) having the following coπposition : C 3= 25 .5 %_ by wt C 3 7 .6 C 4= 43.7 C 4 25.14 In φeration with the preferred feed, a typical φerating pressure is about 445 kPa ( 50 psig) , and temperatu re is about 316 °C ( 600°F ) , though tie reactor may be operated at as low a sψeratmo spheric pressure as 200 kPa (15 psig) . Referring to Figure 1 the re is shown a representati ve curve of a fluid be of equilibrated catalyst referred to as "3 -alpha catalyst" because it is the preferred ave rage alpha for the fluid-bed less than 10% by wt of the fluid-bed has an equilibrated alpha greater than about 5 . The particular preferred ZSM-5 type catalyst chosen for the fluid bed is less critical than the low ave rage alpha, which is turn is less critical than the gene ral distribution of alphas for relati ve weight fractions of catalyst in the bed. In other words, a preferred bed has an ave rage alpha in the raηge from 2 to 10 , but it is essential that preselected weight fractions totalling about 90 % have an alpha no greater than 10 , the reaπdning having an alpha greater than 10. A pref erred f luid-bed for which the curve is shown in Figure 1 has a distribut ion of the ollowing alphas for the weight fractions specified: Rel. wt. frac. alpha 0.50 1 0.25 2 0.10 3 0.05 4 remainder 5 The precise weight fraction having the foregoing approximate alphas is not narrowly critical provided the weight fractions having specified alphas lie in the following ranges: Rarge of rel. wt. frac. alpha 0.20 to 0.60 1 0.15 to 0.35 1 to 3 0.05 to 0.20 3 to 5 0.03 to 0.20 5 to 10 remainder 5 Referring to Figure 2, the phase diagram shovzn is for a most desirable distillate identified by its coπponents listed here-below, over a rarge of teπperature and pressure rarging from about 65.5°C (150°F) and about 1240 kPa (180 psia), to the critical point, which is at about 282°C (540T) and 5994 kPa (870 psia). C 2= 0.03 % by wt. C 3= 0.09 C 3 7.6 C 4= 1.19 C 4 23.14 C 5 "C 9 28.96 C 10 + 38.99 Oligomerization of the feed can be effected in a single zone, that is, a single fluid bed φerating in the dense phase because the heat duty of the reaction exotherm is transferred to cooling fluid lowing through coils in the reactor. The use of prior art multiple fixed beds in each of which only about a 28°C (82°F) exotherm may be acco odated, is avoided. The operating conditions for the fluid bed reactor to produce the foregoing product is as follows: Teπperature (inlet) 300°C (572°F) WffiV 0.3 hr "1 Pressure (inlet) 480 kPa (55 psig) The e is no hydrogen introduced to the reactor, and no li id recycle. It will be understood that the inlet teπperature is typically lower than the bed teπperature because the reaction exotherm raises the teπperature, and heat transfer to the cooling coils in the bed is controlled to maintain the desired bed teπperature. The turbulent fluidization regime in our process is characteristic of a conventional dense phase turbulent fluid bed operating at a pressure below 2857 kPa (400 psig) , and readily distinguishable from non-turbulent dense beds and transport beds. The bed density of our dense phase bed does not exceed 640 kg/m (40 lb/ft ), generally being in the raηge from about 264 kg/m (16.5 lb/ft 3 ) to about 448 kg/m 3 (28. lb/ft 3 ). The bed density is correlatable with operatirg pressure of the bed at a specified siperf icial velocity, as follows: Pressure Bed dens., § 0.3 m/ sec Bed dens., § 0.61 m/sec k a (psig) kg/m 3 (ft 3 /sec) kg/m 3 (ft 3 / sec) 273 ( 25) 608 (38) 400 (25) 1480 (200) 448 (28) 295 (18.5) 2857 (400) 400 (25) 264 (16.5) The preferred particle density is preferably in the raηge from about 1.2 - 2.5 g/ml. A typical dense fluid bed has a minimum fluidization velocity of 0.014 m/sec (0.047 ft/ sec) and φerates at a superficial velocity in the range from about 0.3 - 2.0 m/sec (1 - 6.5 ft/ sec). A destillate with a coπ arable distribution of olefins may be prodiced from the same light gas feed using a low alpha catalyst in multiple fixed bed reactors, but only if they are φerated under higher severity and higher conditions of teπperature, pressure, and residence time as those in the ranges specified for our dense bed. Three fixed bed reactors in series are used as described in the prior art, with prog re ssi vely diminishing inlet teφeratures of 254.4°C (490°F); 248.9°C (480°F); aπi 243.3°C (470T). The first reactor inlet pressure is 6787 kPa (985 psia) and a H 2 circulation rate of 45 v/v (250 SCF BBL) and diluent rate = 47730 kg/hr (105,000 lb/hr). The hydrogen is essential to keep coke deposition on the catalyst to an accφtable level and also as a preveπtitive measure against poisoning of the catalyst by nitrogen- containing organic coπpounds. The analysis for product from the multiple fixed beds is as follows: C 2= 0.14 % by wt 8.4 C 3= C 3 2.3 C 4= 21.14 C 4 5.14 C 5~ C 9 28.28 When the operation of the three fixed beds is changed to match generally the pressure and teπperature conditions of cur low severity fluid bed process, the following conditions prevail in each of the fixed beds: Pressure Teπpe ratu re WHSV, hr -1 kPa (psig) . °C (°F) First 380 (55) 316 (600) 0.5 Second 345 (50) 310 (590) 0.5 Third 310 (45) 304 (580) 0.5 The analysis of product from the multiple fixed beds operated with recycle to maximize C ς + 'make" under the foregoing low severity conditions is as follows: C 2= 0.0 by wt. C 3= 7.1 C 3 1.0 C 4= 26.7 C 4 1.2 C 5 -C 9 = 62.5 C 10 =+ 1.5 Referring now to Figure 3, a light -gas feed rich in C-r-C. olefins passes, through conduit 10, into' the reaction vessel 20, with the main flow being directed through the bottom inlet for distribution through grid plate 22 into the fluidization zone 24. Here the gaseous feed contacts the turbulent bed of finely divided catalyst particles. Reaction vessel 20 is shovn provided with heat exchange tubes 26, which may be arranged as several Sφarate heat exchaηge tube bundles so that teπperature control can be separately exercised over different portions of the fluid catalyst bed. The bottoms of the tubes are spaced above feed distributor grid 22 sufficiently to be free of jet action by the charged feed through the small diameter holes in the grid. Alternatively, a substantial portion of the reaction heat can be removed by using cold feed. Baffles may be added to control radial and axial mixing. Although dφicted without baffles, the vertical reaction zone can contain φen end tubes above the grid for maintaining hydraulic constraints, as disclosed in U.S. Patent No. 4,251,484 (Daviduk and Haddad). Heat released from the reaction may also be controlled by adjusting feed teirperature in a known manner. Catalyst outlet means 28 is provided for withdrawing catalyst from above bed 24 and passed for catalyst regeneration in vessel 30 via control valve 29. The partially deactivated catalyst is preferably stripped with an inert gas (steam or nitrogen) in a stripper (not shown), and oxidatively regenerated by controlled contact with air or other regeneration gas at an elevated teπperature in a fluidized regenera ion zone to remove carbonaceous dφosits and restore acid activity. The catalyst particles are entrained in a lift gas and tran orted via riser tube 32 to a tφ portion of vessel 30. 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 34, which returns any entrained solids to the bed. Flue gas is withdrawn via top conduit 36 for disposal; however, a portion of the flue gas may be recirculated via heat exchaηger 38, sφarator 40, and compressor 42 for return to the vessel with fresh oxidation gas via line 44 and as lift gas for the catalyst in riser 32. Regenerated catalyst is passed to the reaction vessel 20 through conduit 46 provided with flow control valve 48. The regenerated catalyst may be lifted to the catalyst bed with pressurized feed gas, through catalyst return riser conduit 50. Since the amount of regenerated catalyst passed to the reactor is relatively small, the teπperature of the regenerated catalyst does not upset the teπperature constrainsts of the reactor φerations in significant amount. A series of se uentially connected cyclone separators 52, 54 are provided with dip legs 52A, 54A to return ary entrained catalyst fines to the lower bed. These sφarators are positioned in an φper portion of the reactor vessel containing dispersed catalyst phase 24. Filters, such as sintered metal plate filters, can be used alone or in conjuction with cyclones. The product effluent, sφarated from catalyst particles in the cyclone separating system, is then withdrawn from the reactor vessel 20 through t gas outlet means 56. The recovered hydrocarbon product coπprising mainly C..+ olefins with less than 20% by wt aromatics, paraffins and naphthenes is thereafter processed as required, to provide the desired distillate. By virtue of the turbulence experienced in the turbulent regime, gas- solid contact in the catalytic reactor is iπproved, providing at least 80% conversion of C- j -Cg alkenes, enhanced selectivity, and teπperature uniformity. Because of the relatively low pressure, bubbles of the gaseous reaction mixture are relatively large, and relatively loηg- lived, which would normally result in such poor contact between the gaseous reactants and the solid catalyst particles that typically, only a minor amount by weight of the olefins would be olig one ri zed. But the effect of the tailored activity of the catalyst particles is such that at least 75% of propylene and heavier olefins and in the most preferred operation, more than 95% of the butenes and about 99% of the prop ene is converted. Under φtimized process conditions, the turbulent bed has a superficial vapor velocity of 0.1 to 1 m sec (0.328 to 3.28 ft/ sec). At higher velocities, entrairment of fine particles may become excessive and beyond 2 m/sec (6.5 ft/ sec) the entire bed may be transported out of the reaction zone. A preferred, typical dense turbulent bed, has an φerating density of 600 kg/m 3 (38 lb/ft 3 ) to 300 kg/m 3 (18.5 lb/ft 3 ), measured at the bottom of the reaction zone, becoming slightly less dense toward the top of the reaction zone, due to pressure drop and particle size diffeentiation. Pressure differential between two vertically φaced points in the reactor volume may be measured to obtain the average bed density at such portion of the reaction zone. I r instance, in a fluidized bed system eπployiηg HZSM-5 particles laving an apparent packed density of 750 kg/m and real density of 2430 kg/m , and average fluidized bed density of about 500 to 600 kg/m 3 is satisfactory. The MDD operator is φerable with shφe selective medium pore catalysts exeπplified by ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, Z9.-38, ZSM-48 and other similar materials. U.S. Patent No. 3,702,886 describes and claims ZSM-5 ; U.S. Reissue Patent No. Re. 29,948 describes and claims a crystalline material with an X-ray diffraction pattern of ZSM-5 ; and, U.S. Patent No. 4,061,724 describes a high silica Z5M-5 referred to as "silicalite". Similarly reference is made to the disclosures relating to ZSM-11, Z__M- 12, ZSM-23, ZSM-35, ZS4-38, and ZSM-48 set forth in U.S. Patents Nos. 3,709,979, 3,832,449, 4,076,842, 4,016,245, 4,046,859, and 4,375,573. In general the aluminosilicate zeolites are most effectively employed in cur MDD reactor. However, zeolites in which some other framework element which is i so electronic to alιmii-um and which is present in partial or total substitution of alu inum can be advantageous. Illustrative of elements which can be substituted for part or all of the framework aluminum are boron, gallium, titanium, and, in general, ary trivalent metal which is heavier than aluminum. Specific exaπples αf such catalysts include ZSM-5 and zeolite Beta containing boron, gallium and/or titanium. In lieu of, or in addition to, being incorporated into the zeolite framework, these and other catalytically active elements can also be deposited ψon the zeolite by any suitable procedure, e.g., by impregnation. The aluminosilicates are preferred catalysts. These can be described as a three -dimensional framework of SiO^ and A10, tetrahedra in which the tetrahedra are cross- linked by the sharing of oxygen atoms whereby the ratio of total aluπinum and silicon atoms to oxygen atoms is 1:2. In their hydrated form, the aluninosilicates may be represented by the formula: M 2 / n 0:Al 2 O 3 :wSi0 2 : H 2 0 wherein M represents at least one cation which balances the elect o vale n e of the tetrahedra, n represents the valence of the cation, w the moles of Si0 2 0 and Y the moles of H 2 . The cations can be any or more of a nunber of metal ions, dφending upon whether the alu inosilicate is synthesized or occurs naturally. Typical cations include sodium, lithium, potassium, silver, magnesium, calcium, zinc, barium, iron, nickel, cobalt and manganese. Although the prφortions of inorganic oxides in the silicates and their spatial arrangements may vary affecting distinct properties in the aluminosilicate, the main ch racteristic of these materials is their ability to undergo dehydration without substantially affecting the SiO 4. and A104. framework. Aluminosilicates falling within the above formula are well known and, as noted, include synthesized aluminosilicates, natural aluminosilicates, and certain caustic treated clays. Among the aluminosilicates are included zeolites, Y, L, S, X, levynite, erionite, faujasite, analcite, paulingite, noselite, phillipsite, datolite, gmelinite, leucite, scφolite, ordenite as well as certain caustic treated clays such as montmorillonite and kaolin families. The preferred aluminosilicates are those haviiypore diameters of greater than about 6 x 10 ma (6 Argstroras) . Aluminosilicates may be treated with a fluid medium or media in a knovn manner to include a wide variety of aluminosilicates both natural and synthetic which have a crystalline, or, combination of crystalline and amorphous structure. These "promoters" may be provided in the catalyst by impregnation or ion excharge. Though the process of the invention is operable with any of the aluminosilicates the preferred catalyst is 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 withBronsted acid active sites by i corporating a tetrahed rally coordinated metal, such as Al, Ga, B or Fe, within the zeolitic framework. These medium pore zeolites are favored for acid catalysis; however, the advar___ges of ZSM-5 type structures may be utilized by eπplσyirg highly siliceous materials or crystalline metallosilicate having one or more tetrahedral species havάηg varying degrees of acidity. The ZSM-5 crystalline structure is readily recognized by its X-ray diffraction pattern, which is described inU.S. Patent No. 3,702,866 (Argauer, et al) . The oligomerization catalysts preferred for use herein _ — include the medium pore (i.e., 5-7 x 10 mm (5-7 _U) shape-selective crystalline aluminosilicates zeolites having a silica-to-alumina ratio of at least 12, a constraint index of about 1 to 12, acid cracking activity of about 10-250, and, in the M)D reactor, it is essential that the coked catalyst have the tailored equilibrated activity referred to hereinabove to achieve the reψired selectivity and per pass conversion. Representative of the ZSU-S type zeolites are ZSM-5 , ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, and ZSM-38. Other suitable zeolites are disclosed in UD. Patents Nos. 3,709,979; 3,832,449; 4,076,979; 3,832,449; 4,076,842; 4,016,245; 4,046,839; 4,414,423; 4,417,086; 4,517,396 and 4,542,251. A typical zeolite catalyst coπponeπt having Bronsted acid sites may consist essentially of aluminosilicates ZS -5 zeolite with 5 to 95 wt % silica and/or aluπina binder. These siliceous zeolites may be eπpl oyed in their acid forms, ion exchanged, or iπpregnated with one or more suitable metals, such as Ga, Pd, Zn, Ni, Co and/or other metals of Periodic Groups III to VIII. The zeolite may include a hydrogen at ion-de hydrogen at ion coπponeπt (sometimes referred to as a hydrogenation coπponeπt) which is generally one or more metals of groφ IB, IB, IIIB, VA, VIA, or VIIIA of the Periodic Table (IUPA.) , especially a ro mat izat ion metals, such as Ga, Pd, etc. Useful hydrcgenation coπponents include the noble metals of Group VIIIA, especially platinium, but other noble metals, such as palladium, gold, silver, rhenium or rhodium, may also be used. Base metal hydrogenation coπponents may also be used, especially nickel, cobalt, molybdenum, tungsten, copper or zinc. The catalyst materials may include two or more catalytic coπponents, such as a metallic oligomerization coπponent (e.g. ionic Ni , and a shape-selective medium pore acidic oligomerization catalyst, such as ZΣ-M-5 zeolite) which coπponents may be present in admixture or combined in a unitary bifunctional solid particle. Certain of the ZΣ.M-5 type medium pore shφe selective catalysts are sometimes known as peπtasils. In addition to the preferred aluminosilicates, the bo ro si lie ate, ferrosilicate and "silicate" materials may be employed. It is advantageous to employ a standard ZS4-5, suitably modified, h v ng a silica: alumina molar rato in the range from 12:1 to 100:1, a constraint index in the raηge from 5 to 12, and with the aforesaid average alpha to convert substantially all the olefins in the feedstock. Z5M-5 type peπtasil zeolites are particularly useful in the process because of their reg ene lability, long life and stability under the extreme conditions of φe rat ion. Usually the zeolite crystals have a crystal size from 0.01 to over 2 ^ rn. or more, with 0.02-1 m being preferred, and an φparent crystal density in the range from 0.6 to 1.9 gm/ml. In order to obtain the desired particle size for fluidization in the turbulent regime, the zeolite catalyst crystals are bound with a suitable inorganic oxide, such as silica, alumina, etc. to provide a zeolite concentration of about 5 to 95 wt %. In the best mode a bed of 25% HZSM-5 catalyst contained within a silica -alumina matrix and having the aforementioned distribution of equilibrated alpha is eφlσyed. Operation of a dense phase turbulent regime with the preferred catalyst is distinguished by its particle size distribution to achieve overall homogeneity, particularly reψiring from about 10% to 20% of "fines". It is advantageous to employ a particle size rarge consisting essentially of 1 to ZOO^ . Average particle size is usually 20 to 150 ^m, preferably 50 to 100 m. The φtimum particle size distribution is obtained with a mixture of larger and smaller particles within the above-specif ied range, haviηg from 15-20% by weight fines. Close control of distribution is maintained with the fines in the size range less than 32 ^m. The average particle density of the catalyst as used may be tailored for φtimum fluid-bed operation by coπpositing it with a matrix coπponent of φpropriate density. Such matrix coπponents which provide particles of progressively increasing overall packed density are silica, alumina, beryllia, magnesia, barium oxide, zirconia, and titania, yielding values of from 2.2 gm/ml for silica, to 5.9 gm/ml for zirconia. In our MDD reactor, the overall packed density of medium pore zeolite particles so coπposited, including the matrix coπponent, can advanta eously vary from 0.6 to 4 gm/ml, more preferably from 2 to 3 gm/ml. Several useful parameters contribute to fluidization in the turbulent regime in accordance with the process of the present invention. When employing a ZS -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 φparent "outside" volume) in the rarge from 0.6-2 g/ml, preferably 0.9-1.6 g/ml. When these solid particles are placed in a fluidized bed whe re the superficial fluid velocity is 0.06-0.5 m s, φe rat ion in the turbulent regime is obtained. Those skilled in the art will φpreciate that at higher pressures in the rarge, a lower gas velocity may be employed to ensure operation in the turbulent fluidization regime. The MDD reactor is designed as a moderate pressure vessel reφired to contain a turbulent fluid-bed φe rating to produce φtiπau yields of the desired oligomers, and iπportant criteria are taken into consideration. The bed of catalyst in the reactor is desirably in the range from 3-20 meters in height, preferably about 8 meters, and, φerates at relati vely low WHSV, most preferably less than 1 hr. "1 . Fine particles are deliberately included in the bed, some of which are due to attrition, and the fines may be entrained in the product gas stream. A typical turbulent bed may have a catalyst carryover rate of about 1.5 times the reaction zone inventory per hour. If the fraction of fines becomes large, a portion of the carryover may be removed from the system and replaced by larger particles. It is desirable to have a fine particle sφarator such as a cyclone and/or filter means, disposed within or outside the reactor shell to recover catalyst carryover and return this fraction continuously to the bottom of the reaction zone for recirculation at a rate of about one catalyst inventory per hour. Optionally, fine particles carried from the reactor vessel entrained with effluent gas can be recovered by a sintered metal filter φe rating at the MDD reactor conditions. Typically, in carrying out the process in the relatively low pressure MDD reactor, the light gas feed is brought into contact with a medium pore zeolite catalyst haviig a silica to alumina ratio in the rarge from 12 to 1000, m re preferably from 50 to 90, maintained as a dense phase in a fluid-bed at about 274 kPa to 1480 kPa (25 psig - 200 psig) , at a teπperature below 371°C (7007) at which conditions no li id will form in the bed, while the light gas is f loved through the reactor at a sψerf icial vapor velocity preferably in the range from 0.15 m/sec (0.5 f/sec) to 0.305 m/sec (1 ft/ sec) over a HZ&1-5 catalyst having a constraint index in the range from 5-12, and, an eφilibrated alpha in conformance with the curve shovn. in Figure 1, and φerating with a WHSV (based on olefin eφivalent and total reactor catalyst inventory) in the range from about 0.1 to 1 hr "1 . The operation of the turbulent luid bed in the dense phase produces a remarkably low coking rate, less than 0.3 wt %, and preferably less than 0.05 wt % of the olefins in the feed, which low rate allows one to φerate the bed without regeneration for a long period of time. In some instances periodic regeneration may be a • preferred mode of φe ration. If not continuously regenerated, the coke content will gradually increase to about 15% by wt of the catalyst, at which point the run is desirably terminated. When continuous regene ration is chosen, the regenerator will be φeiated at the lowest pressure at which effective decokiηg can be ickly effected without exceeding a teπperature deleterious to the activity of the regenerated catalyst. Typically, the regeneration pressure is in the same rarge as that for the reactor. Whe t he r withdrawal of coked-up catalyst is periodic or continuous, the desired activity of the regenerated catalyst is maintained such that the eφilibrated activity of the fluid-bed durirg operation of the reactor will be generally in coif oimance with the curve shown in Figure 1. Eφilibrating the activity of the specif ied weight fractions of the fluid- bed catalyst, rather than providing a generally uniform fluid-bed with the desired average alpha, is the distirguishi g characteristic of cur process. The uni e effect of this tailored eφilibrated activity is evidenced by (i) a per pass conversion of lower olefins of at least 80%, typically better than 90%, and a selectivity to Cr+ hydrocarbons typically exceeding 80%; and (ii) formation of less than 5% by wt of aromatics. Because Cg- hydrocarbons in the product are not particularly desirable, operating conditions are chosen so that a relatively very small amount of gasoline rarge hydrocarbons, less than about 20% by wt of distillate to gasoline range hydrocarbons is 5:1, and may be as much as 10:1. Antoher distinguishing characteristic of our process is that lower C 2 - r alkanes in the feed, if present, are substantially unconverted; that is, there is no net depletion of alkanes in the product. This can be accounted for by the foimation of some C 2 -C alkanes with a purely C j -C^ olefinic feed. To maintain the high per pass conversion and selectivity in the dense bed, the catalyst is preferably continously regenerated. Whether regeneration is periodic or continuous, it is done by withdrawing a portion of coked catalyst from the reaction zone, oxidatively regenerating the withdrawn cata st and returning regenerated catalyst to the reaction zone at a rate sufficient to maintain the desired selectivity and per pass conversion, each at least 80%, based on the total olefins in the feedstream. It will be φpreciated that the design, construction and operating conditions of a coπme rcial oligomerization reaction vessel will be dictated by the economics of producing distillate, and the optimum φeratiηg conditions will be dictated by the conversions sought. Such a reactor will h ve a height to diameter ratio of at least 5, a pressure rating of at least 2800 kPa, and a teπperature rating of at least 380°C. The feed to the reactor will preferably be preheated by the internal coils 26 through which liφid feed is puφed under sufficient pressure to provide a gaseous feed to the fluid bed, and at a teπperature not much lower than about 20° C from the desired φe rating teπperature in the fluid bed. The regenerator will typically operate in the range from 3 1°C - 570°C (700°F - 950*F) , and the flue gas from the regeneration zone will be cooled to a sufficiently low teπperature in the range from 35°C - 50°C (95T - 122°F) , so that a portion of it may be recycled to the regeneration zone, and the remainder discharged. In the best mode for the production of distillate raηge hydrocarbons from an essentially monoolef inic light gas feedstock ha ii-g at least 75% by wt of C,-C 6 alkenes, the oligoπe rized product will contain less than about 80% by wt of (Cg-) hydrocarbons, and about 20% by wt, or m re, of (C g +) hydrocarbons. In most instances, a major portion of the C 8 - hydrocarbons are separated from the product stream and recycled to the dense reaction zone. Baπple In this illustrative example, the process is carried out in a 640 1/ stream day (4 BPSD (bbls/ stream day)) pilot plant in which a moderate pressure reactor is used. The rector is provided with a bed of medium pore ffiSM-5 zeolite catalyst l ving a silica to alumina ratio of about 70, and the usual means for controlling fluidization αf the bed in the turbulent regime. The _£ΣM-5 catalyst has a constraint index of about 10. The activity of the bed is tailored so as to fall φproximately along the curve shown in Figure 1. The feedstrea is essentially pure propylene which is brought into contact with the catalyst which is maintained as dense phase in a fluid-bed at a pressure of 445 kPa (50 psig) and, with the inlet teπperature of the feed, beirg 254.4°C (490T), the teπperature of the bed is maintained at 315.5°C (600°F). The light gas is flowed through the reactor at a WI-SV of 0.36 hr , operating in the distillate mode, and the results are for a single pass, that is, without any recycle of Cr-C 8 gasoline rarge hydrocarbons. For coirparison, the same feed is flowed through a series of fixed bed reactors φe rating under low severity conditions defined by substantially the same pressures and temperatures specified hereinabove for them, with the same flow of hydrogen and inert diluents, and recycle of lower olefins to maximize the production of coπponents. The analysis of products made in each of the runs are presented side- by- side in Table I below, as % by wt of the product: TABLE I - % by wt in the product C 2= 0.2 0.0 % by wt C 3 5.62 7.1 C 3 0.38 1.0 C 4= 12.02 26.7 C 4 1.96 1.2 C 5 "Cg= 67.93 62.5 11.75 1.5 C 10 =+ As is evident from the foregoing example, the per pass conversion is better than 80%. Typically, the per pass conversion is over 90%, though in some cases where the ole in content is relatively low, the per pass conversion may be less. Having thus provided a general discussion, and a φecif ic illustration of the best mode of operation of a single zone, dense, turbulent fluid bed φe rating at relatively low pressure and moderate temperature, and described tie oligom rization of a predominantly monomeric olefinic light gas feedstream in such a bed, it is to be understood that no undue restrictions are to be imposed by reason thereof, excφt as provided by the f ollowi g claims.

Claims: CLAIMS 1. A catalytic process for converting a feedstream contaiπLηg C 2 -C^ lower olefins to distillate in a fluid bed of catalyst, the process coπprisiηg, (a) eφilibrating a fluid bed of catalyst comprising finely divided medium pore zeolite metallosilicate particles having a constraint index in the rarge from 1 to 12, the bed baving an average activity, alpha, in the range from 2 to 10, the bed consisting essentially of preselected weight fractions of low activity catalyst particles addiηg up to 90% by weight haviηg an alpha no greater than 10, the remaining less than 10% by weight having an alpha greater than 10; (b) contacting the olefin feedstream with the catalyst maintained as a dense fluid-bed φerating in a siigle zone turbulent regime at si eiatmo spher c pressure but below 2800 kPa and a teφerature above the devpoiπt of the hydrocarbon mixture in the bed at operating pressure, (c) f lcwiηg the olefin feedstream through the bed at a weight hourly space velocity in the rarge from 0.05% to 5 hr " , (d) maintaining a catalyst fines content of from 10% to 25% by wt, based on the weight of the catalyst in the bed, the fines having a particle size less than 32^*_. m, and, (e) converting at least 80% by wt of the feedstream to olefins in a product which is essen ially free of aromatics, with a yield per pass of 20 percent by weight of the distillate. 2. The process of claim 1 wherein the olefin feedstream consists essentially of C^-C^ monoole ins, and, the averεge alpha of the fluid bed consists of relative weight fractions having the following alphas: Range of rel. wt. frac. alpha 0.20 to 0.60 1 0.15 to 0.35 1 to 3 0.05 to 0.20 3 to 5 0.03 to 0.20 5 to 10 remainder 5 3. The process of claim 2 whe rein the catalyst has silica :alumina ratio greater than 12, and said bed operates at a bed density, measured at the bottom of the reaction zone, no greater than 608 kg/m , a teπperature in the range from 250°C to 371°C, and a pressu re in the raηge from abo ve 200 kPa but below 2857 kPa. 4. The process of claim 3 wherein the catalyst is a siliceous metallosilicate acid zeolite ha ving a ZSM-5- type structure, the ave rage alpha is 3 , less than 10 % by weight of the catalyst has an eφilibrated alpha above 10 , and the si lica :alumina ratio is in the range from 12 : 1 to 70 : 1. * 5. The process of claim 4 wherein said C, -C^ monoolefins comprise a major prφortion by weight of prφene and butenes. 6 . The process of claim 4 wherein the f luid bed is maintained in a reactor φerated at relati vely low pressure in the range from 274 kPa to 1480 kPa, and moderate temperature in the raηge from 204°C to 371°C, and, the distillate has a boiling in the range from 138° C to 349 °C. 7 . The process of claim 4 wherein the essentially monoolefinic feedstream is essentially free of hydrogen. 8 . The process of claim 1 whe rein the feestream contains a minor prφortion by weight of dilutent C 2 ~C, paraffins; the feedstream is passed as a g as upward ly through the f luidized bed in a ve rtical reactor column, while maintaining a sφerf icial velocity in the rarge from about 0.3 m/sec to about 2 m/sec so that catalyst particles in which the silica :alumina molar ratio is in the raηge from about 20 :1 to about 200:1 a re held in a turbulent regime; the average bed density, measured at the bottom of the fluidi zed bed is maintained in the raηge from 250 to 600 kg/m , at a pressure in the range from 200 kPa to 2800 kPa and a teπperature in the raηge from 204°C to 371°C; and a portion of coked catalyst is withdrawn from the reaction zone, oxidati vely regenerating the withdrawn catalyst and returni ng regenerated catalyst to the f luidized bed at a rate sufficient to maintain a reaction se verity index, expressed as the prφane: prφene weight ratio in the hydrocarbon product, in the raηge from about 0.2:1 to 5:1, whereby at least 10% by weight of the feedstream is converted to a substantially olefinic C g =+ distillate. 9. The process of claim 8 wherein the olefinic content of the feedstream consists essentially of at least 75% by wt of propylene and heavier olefins (C j =+), and the turbulent bed has a fluidized height of at least 3 meters. 10. The process of claim 8, including in addition, the steps of separating hydrocarbons having less than 9 carbon atoms (Cg-) from the distillate, and, recycling the (C 8 -) hydrocarbons to the reactor. 11. A fluidized bed catalytic process for converting a light olefinic gas feedstock to a distillate- rich predominantly olefinic stream, in a low activity, 'low-alpha', fluid bed of catalyst, which bed φerates at relatively low pressure and moderate teπperature, coπprising, feeding the gas feedstock coπprising at least 50 ol % propylene and header olefins C,=+, to a catalytic fluid bed reaction zone, maintaining the fluid bed in a turbulent regime wherein pressure is in the rarge above atmospheric but below 2850 kPa, and teπperature in the range from 250°C to 371°C, the catalyst being a particulate zeolite h ving a silica :alumina molar ratio in the rarge from 20:1 to 200:1, an apparent particle density of 0.9 to 1.6 g/ml , a size range of 20 to and average catalyst particle size of 20 to lOO^cf m containing 15 to 25 weight percent of fine particles having a particle size of less than 32t_ m; passing the gas feedstock in the substantial absence of added propane upwardly through the fluid bed in a single pass under turbulent flow conditions which provide reaction severity conditions sufficient to convert at least 70% of olefins in the feedstock; maintaining turbulent conditions at a sφerf icial feedstock velocity of 0.3 to 2 meters per second through the bed having an average fluidized bed density, measured at the bottom of the reaction zone, in the raηge from 200 to 400 kg/ra ; and recovering hydrocarbon product containing a major amount of C,=+ hydrocarbons containing at least 10% by weight of distillate and a ratio of propane: propene in the range from above 0.01:1 but below 0.2:1. 12. In a continuous process for converting a light hydrocarbon feedstock to heavier hydrocarbon products wherein the feedstock is contacted with a fluidized bed of zeolite catalyst under relatively low pressure and moderate teπperature conversion conditions, the improvement which comprises; maintaining the fluidized catalyst bed having an average density measured at the bottom of the reaction zone in the range from 200 to 400 kg/m , at a pressure in the range from 274 kPa to 2857 kPa and a teφerature in the raηge from 250°C to 371°C, in a vertical reactor column having a turbulent reaction zone, by passing the gas feedstock comprising at least 50 mol % prφylene and heavier ole ins C,=+, while maintining a sφerficial velocity in the range from 0.3 m/sec to 2 m/sec; and withdrawing a portion of coked catalyst, in which the silica :alumina molar ratio is in the raηge from 20:1 to 200:1, from the reaction zone, oxidatively regenerating the withdrawn catalyst and returning regenerated catalyst to the reaction zone at a rate sufficient to maintain a reaction severity index expressed as the prφane: prφene weight ratio in the hydrocarbon product at 0.01:1 to 0.2:1 whereby at least 70% of olefins in the feedstock is converted.

Inventors: AVIDAN AMOS ANDRES, JOHNSON DAVID LEE, AVIDAN, AMOS, ANDRES, JOHNSON, DAVID, LEE
Applicants: MOBIL OIL CORP, MOBIL OIL CORPORATION
Application year: 1989
Main classification: C07C2
IPC codes: B01J8, C07C2