EP0130673B1 - Verfahren zur Konvertierung von Olefinen zu höheren Kohlenwasserstoffen - Google Patents

Verfahren zur Konvertierung von Olefinen zu höheren Kohlenwasserstoffen Download PDF

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Publication number
EP0130673B1
EP0130673B1 EP19840302911 EP84302911A EP0130673B1 EP 0130673 B1 EP0130673 B1 EP 0130673B1 EP 19840302911 EP19840302911 EP 19840302911 EP 84302911 A EP84302911 A EP 84302911A EP 0130673 B1 EP0130673 B1 EP 0130673B1
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Prior art keywords
reactor
stream
gasoline
liquid
effluent
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EP19840302911
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English (en)
French (fr)
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EP0130673A3 (en
EP0130673A2 (de
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Hartley Owen
Chung Hweng Hsia
Bernard Stanley Wright
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ExxonMobil Oil Corp
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Mobil Oil Corp
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Priority claimed from US06/508,959 external-priority patent/US4471147A/en
Priority claimed from US06/508,907 external-priority patent/US4450311A/en
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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G50/00—Production of liquid hydrocarbon mixtures from lower carbon number hydrocarbons, e.g. by oligomerisation
    • F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
    • F02—COMBUSTION ENGINES; HOT-GAS OR COMBUSTION-PRODUCT ENGINE PLANTS
    • F02B—INTERNAL-COMBUSTION PISTON ENGINES; COMBUSTION ENGINES IN GENERAL
    • F02B3/00—Engines characterised by air compression and subsequent fuel addition
    • F02B3/06—Engines characterised by air compression and subsequent fuel addition with compression ignition

Definitions

  • This invention relates to a process for converting olefins into higher hydrocarbons, for example gasoline boiling-range and distillate boiling-range fuels.
  • distillate-mode reactor systems designed completely to convert a large ethylenic component of feedstock would require a much larger size than comparable reactor systems for converting other lower olefins.
  • Recycle of a major amount of ethylene from the reactor effluent would result in significant increases in equipment size.
  • propene and butene are converted efficiently, 75 to 95% or more in a single pass, under catalytic conditions of high pressure and moderate temperature used in distillate mode operations.
  • the present invention is based on the observation that an olefins-to-distillate process using a C 2 ⁇ C 4 olefinic feedstock can be combined with a feedstock prefractionating step; in this manner, the C 3 + olefinic components can be catalytically converted into a distillate boiling-range product and ethylene can be recovered economically for use in polymer manufacture or in other industrial processes.
  • the prefractionation step (a) of the process is described in greater detail below.
  • the olefinic feedstock consists essentially of C 2 -C 6 aliphatic hydrocarbons containing a major fraction of monoalkenes in the essential absence of dienes and other deleterious materials.
  • the process may employ various volatile lower olefins as feedstock, with oligomerization of C 3 + alpha-olefins being preferred for either gasoline or distillate production.
  • the olefinic feedstock contains from 50 to 75 mole % C 3 -C 5 alkenes.
  • the members of the class of crystalline zeolites that may be used in the process of the invention are characterized by a pore dimension greater than about 5 Angstroms (0.5 nm); i.e., they are capable of sorbing paraffins having a single methyl branch as well as normal paraffins, and they have silica to alumina mole ratios of at least 12. Although such crystalline zeolites with a silica to alumina mole ratio of at least about 12 are useful, it is preferred to use zeolites having higher ratios of at least about 30. In some zeolites, the upper limit of silica to alumina mole ratio is virtually unbounded; i.e., values of 30,000 and greater.
  • the members of this class of zeolites are exemplified by ZSM-5, ZSM-5/ZSM-11 intermediate, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48 and other similar materials. These zeolites are described in, for example, U.S. Patents 3,702,886; Re. No. 29,948; 4,061,724; 4,229,424; 3,709,979; 3,832,449; 4,076,842; 4,016,245; and 4,046,859; and EP-A-13640.
  • the catalyst and separate additive composition may be prepared in various ways. They may be separately prepared in the form of particles such as pellets or extrudates, for example, and simply mixed in the required proportions.
  • the particle size of the individual component particles may be quite small, for example from about 10 to about 150 pm, when intended for use in fluid bed operation, or they may be as large as 1-10 mm for fixed bed operation.
  • the components may be mixed as powders and formed into pellets or extrudate, each pellet containing both components in substantially the required proportions. It is desirable to incorporate the zeolite component of the separate additive composition in a matrix. Such matrix is useful as a binder and imparts greater resistance to the catalyst to the severe temperature, pressure and velocity conditions encountered in many cracking processes.
  • Matrix materials include both synthetic and natural substances; for example, clays, silica and/or metal oxides.
  • the latter may be either naturally occurring or in the form of gelatinous precipitates, sols or gels including mixtures of silica and metal oxides.
  • zeolite materials have been incorporated into naturally occurring clays; for example, bentonite and kaolin.
  • a-alumina monohydrate in a proportion of about 2:1
  • olefinic feedstock is supplied to the plant through a fluid conduit 1 under steady stream conditions.
  • the olefins are separated in a prefractionator 2 to recover an ethylene-rich stream 2E and a liquid hydrocarbon stream 2L containing C 3 + feedstock components, as described in detail below.
  • This C 3 + feedstream is pressurized by a pump 12 and then sequentially heated by passing through indirect heat exchange units 14 and 16 and furnace 20 to achieve the temperature for catalytic conversion in reactor system 30, including reactor vessels 31 A, 31 B, 31 C, etc.
  • the reactor system section shown consists of 3 downflow fixed bed, series reactors on line with exchanger cooling between rectors.
  • the reactor configuration allows for any reactor to be in any position, A, B or C.
  • the reactor in position A has the most aged catalyst and the reactor in position C has freshly regenerated catalyst.
  • the cooled reactor effluent is fractionated first in a debutanizer 40 to provide lower aliphatic liquid recycle and then in splitter unit 50 which not only separates the debutanizer bottoms into gasoline and distillate products but also provides liquid gasoline recycle.
  • the gasoline recycle is necessary not only to produce the proper distillate quality but also to limit the exothermic rise in temperature across each reactor to less than 30°C. Change in recycle flow rate is intended primarily to compensate for gross changes in the feed non-olefin flow rate. As a result of preheat, the liquid recycle is substantially vaporized by the time that it reaches the reactor inlet. The following is a description of the process flow in detail.
  • Sorbed C 3 + olefin combined with olefinic gasoline is pumped up to system pressure by pump 12 and is combined with gasoline recycle after that stream has been pumped up to system pressure by a pump 58.
  • the combined stream (C3 feed plus gasoline recycle) after preheat is routed to inlet 30F of reactor 31A of system 30.
  • the combined stream (designated as the reactor feedstream) is first preheated against the splitter tower 50 effluent in exchanger 14 (reactor feed/splitter tower bottoms exchanger) and then against the effluent from the reactor in position C, in exchanger 16 (reactor feed/reactor effluent exchanger).
  • the reactor feed is heated to the required inlet temperature for the reactor in position A.
  • the effluents from the reactors in the first two positions, A and B are cooled to the temperature required at the inlet of the reactors in the last two positions, B and C, by partially reboiling the debutanizer 40. Temperature control is accomplished by allowing part of the reactor effluents to bypass the reboiler 42. Under temperature control of the bottom stage of the sorption fractionator 2, energy for reboiling is provided by at least part of the effluent from the reactor 31 in position C.
  • the reactor effluent reboils deethanizer bottoms 61 and is then routed to the debutanizer 40 which is operated at a pressure which completely condenses the debutanizer tower overhead 40V by cooling in condenser 44.
  • the liquid from debutanizer overhead accumulator 46 provides the tower reflux 47, and feed to the deethanizer 60, which, after being pumped to deethanizer pressure by pump 49 is sent to the deethanizer 60.
  • the deethanizer accumulator overhead 65 is routed to the fuel gas system.
  • the accumulator liquid 64 provides the tower reflux.
  • the bottoms stream 63 (LPG product) may be sent to an unsaturated gas plant or otherwise recovered.
  • the bottoms stream 41 from the debutanizer 40 is sent directly to the splitter, 50 which splits the C S + material into C r -165 * C gasoline (overhead liquid product and recycle) and 165°C + distillate (bottoms product).
  • the splitter tower overhead stream 52 is totally condensed in the splitter tower overhead condenser 54.
  • the liquid from the overhead accumulator 56 provides the tower reflux 50L, the gasoline product 50P and the specified gasoline recycle 50R under flow control, pressurized by pump 58 for recycle.
  • the gasoline product After being cooled in the gasoline product cooler 59, the gasoline product is sent to the gasoline pool.
  • the splitter bottoms fraction is pumped to the required pressure by pump 58 and then preheats the reactor feed stream in exchanger 14.
  • the distillate product 50D is cooled to ambient temperature before being hydrotreated to improve its cetane number.
  • a kettle reboiler 42 containing 2 U-tube exchangers 43 in which the reactor 31 effluents are circulated is a desirable feature of the system. Liquid from the bottom stage of debutanizer 40 is circulated in the shell side.
  • the thermal integration techniques employed in the system depicted in Fig. 2 provide flexible process conditions for startup and steady state operation of MOGD feedstock and effluent fractionation subsystems.
  • the reaction section effluent After preheating the reactor feed, the reaction section effluent reboils prefractionation liquid bottoms and the deethanizer before mixing with the sponge absorber bottoms and entering the debutanizer.
  • Prefractionated olefinic feedstock is fed to the reactor after receiving some preheat from the distillate product stream and, depending on the third reactor effluent temperature, the reactor feedstock may also receive preheat from the reactor effluent before entering the furnace, where it is heated to the temperature required for the reactor in initial position A.
  • the effluents from the first two reactors are cooled to the inlet temperatures for the last two reactors by reboiling the debutanizer and product splitter.
  • Reactor inlet temperature control is achieved by regulating the amount of first reactor effluent sent to the gasoline/distillate splitter reboiler and the amount of intermediate reactor effluent sent to the debutanizer reboiler.
  • the amount of first reactor effluent sent to the debutanizer reboiler is temperature controlled by the debutanizer bottom stage temperature. If needed, a portion of the first reactor effluent sent to the product splitter may be routed through the furnace convection section for auxiliary heating.
  • the product fractionation units 40, 50, and 60 may be a tray-type design or packed column.
  • the splitter distillation tower 50 is preferably operated at substantially atmospheric pressure to avoid excessive bottoms temperature, which might be deleterious to the distillate product.
  • the fractionation equipment and operating techniques are substantially similar for each of the major stills 40, 50, 60, with conventional plate design, reflux and reboiler components.
  • the fractionation sequence and heat exchange features of the present system are operatively connected in an efficient MOGD system provide significant economic advantages.
  • MOGD operating modes may be selected to provide maximum distillate product by gasoline recycle and optimal reactor system conditions. Operating examples are given for distillate mode operation, utilizing as the olefinic feedstock a pressurized stream olefinic feedstock (about 1200 kPa) comprising a major weight and mole fraction of C 3 -/C4 .
  • the adiabatic exothermic oligomerization reaction conditions are readily optimized at elevated temperature and/or pressure to increase distillate yield or gasoline yield as desired, using HZSM-5 type catalyst.
  • Particular process parameters such as space velocity and maximum exothermic temperature rise may be optimized for the specific oligomerization catalyst employed, olefinic feedstock and desired product distribution.
  • a typical distillate mode multi-zone reactor system employs inter-zone cooling, whereby the reaction exotherm can be carefully controlled to prevent excessive temperature above the normal moderate range of about 190° to 315°C.
  • the maximum temperature differential (AT) across any one reactor is about 30°C and the space velocity (LHSV based on olefin feed) is about 0.5 to 1.
  • Heat exchangers provide inter-reactor cooling and reduce the effluent to fractionation temperature. It is an important aspect of energy conservation in the MOGD system to utilize at least a portion of the reactor exotherm heat value by exchanging hot reactor effluent from one or more reactors with a fractionator stream to vaporize a liquid hydrocarbon distillation tower stream, such as the debutanizer reboiler. Optional heat exchangers may recover heat from the effluent stream prior to fractionation.
  • Gasoline from the recycle conduit is pressurized by a pump and combined with feedstock, preferably at a mole ratio of about 1-2 moles per mole of olefin in the feedstock. It is preferred to operate in the distillate mode at elevated pressure of about 4200 to 7000 kPa.
  • the reactor system contains multiple downflow adiabatic catalytic zones in each reactor vessel.
  • the liquid hourly space velocity (based on total fresh feedstock) is about 1 LHSV.
  • the inlet pressure to the first reactor is about 4200 kPa, with an olefin partial pressure of at least about 1200 kPa. Based on olefin conversion of 50% for ethene, 95% for propene, 85% for butene-1 and 75% for pentene-1, and exothermic heat of reaction is estimated at 1050 kJ/kg of olefins converted.
  • a maximum AT in each reactor is about 30°C.
  • the molar recycle ratio for gasoline is equimolar based on olefins in the feedstock, and the C 3 -C 4 molar recycle is 0.5:1.
  • the prefractionation system is adapted to separate volatile hydrocarbons comprising a major amount of C 2- C 4 olefins, and typically contains 10 to 50 mole % of ethene and propene each.
  • the feedstock consists essentially of volatile aliphatic components as follows (in mole %): ethene, 24.5%; propene, 46%; propane, 6.5%; 1-butene, 15%; and butanes, 8%, having an average molecular weight of about 42 and more than 85% olefins.
  • the total gasoline sorbent stream to feedstockweight ratio is greater than about 3:1; however, the content of C 3 + olefinic components in the feedstock is a more preferred measure of sorbate to sorbent ratio.
  • the process may be operated with a mole ratio of about 0.2 moles to about 10 moles of gasoline per mole of C 3 + hydrocarbons in the feedstock, with optimum operation utilizing a sorbent:sorbate molar ratio about 1:1 1 to 1.5:1.
  • olefinic feedstock is introduced to the system through a feedstock inlet 1 connected between stages of a fractionating sorption tower 2 in which gaseous olefinic feedstock is contacted with liquid sorbent in a vertical fractionation column operating at least in the upper portion thereof in countercurrent flow. Effectively this unit is a C Z /C 3 + splitter.
  • Design of sorption equipment and unit operations are established chemical engineering techniques, and generally described in Kirk-Othmer "Encyclopedia of Chemical Technology" 3rd Ed. Vol. 1 pp. 53-96 (1978). In conventional refinery terminology, the sorbent stream is sometimes known as lean oil.
  • Sorption tower 2 has multiple contact zones, with the heat of absorption being removed via interstage pumps around cooling circuits 2A, 2B.
  • the liquid gasoline sorbent is introduced to the sorption tower through an upper inlet 2C above the top contact section 20. It is preferred to mix incoming liquid sorbent with outgoing splitter overhead ethylene-rich gas from upper gas outlet 2E and to pass this multi-phase mixture into a phase separator 2F, operatively connected between the primary sorption tower 2 and a secondary sponge absorber 3. Liquid sorbent from separator 2F is then pumped to the upper liquid inlet 2C for countercurrent contact in a plate column or the like with upwardly flowing ethylene-rich vapors.
  • Liquid from the bottom of upper contact zone 20 is pumped to a heat exchanger in loop 2A, cooled and returned to the tower above intermediate contact zone 2G, again cooled in loop 2B, and returned to the tower above contact zone 2H, which is located below the feedstock inlet 1.
  • a heat exchanger in loop 2A cooled and returned to the tower above intermediate contact zone 2G, again cooled in loop 2B, and returned to the tower above contact zone 2H, which is located below the feedstock inlet 1.
  • the lower contact zone 2H provides further fractionation of the olefin-rich liquid. Heat is supplied to the sorption tower by removing liquid from the bottom via reboiler loop 2J, heating this stream in heat exchanger 2K, and returning the reboiled bottom stream to the tower below contact zone 2H.
  • the liquid sorbate-sorbent mixture is withdrawn through bottom outlet 2L and used as feedstock in the olefins oligomerization unit.
  • Ethylene rich vapor from the primary sorption tower is withdrawn via separator 2F through conduit 3A.
  • Distillate lean oil is fed to the top inlet 3B of sponge absorber 3 under process pressure at ambient or moderately warm temperature (for example 40°C) and distributed at the top of a porous packed bed, such as Raschig rings, having sufficient bed height to provide multiple stages.
  • the liquid rate is low; however, the sponge absorber permits sorption of about 25 weight percent of the distillate weight in C 3 + components sorbed from the ethylene-rich stream.
  • This stream is recovered from bottom outlet 3C. It is understood that the sorbate may be recovered from the mixture with the sorbent by fractionation and the sorbent may be recycled or otherwise utilized.
  • High purity ethylene is recovered from the system through gas outlet 3D and sent to storage, further processing or conversion to other products.
  • the sorption towers depicted in the drawing employ a plate column in the primary tower and a packed column in the secondary tower, however, the fractionation equipment may employ vapor-liquid contact means of various designs in each stage including packed beds of Raschig rings, saddles or other porous solids or low pressure drop valve trays (Glitsch grids).
  • the number of theoretical stages will be determined by the feedstream composition, liquid:vapor (UV) ratios, desired recovery and product purity. In the detailed example below, 17 theoretical stages are employed in the primary sorption tower and 8 stages in the sponge absorber, with olefinic feedstock being fed between the 7th and 9th stages of the primary sorption tower.
  • Example 1 illustrates the invention. They are based on the feedstock described above at 40°C and 2100 kPa supplied to stage 9 of the primary sorption tower. Gasoline is supplied at 85°C and 2150 kPa, and distillate lean oil is supplied at 40°C and 2100 kPa. Table I shows the conditions at each stage of the primary sorption tower, and Table II shows the conditions for the sponge absorber units for Example 1 (2 moles gasoline/mole of olefin in feedstock).
  • the C 3 + olefin sorbate and gasoline are fed directly to the oligomerization process, with a portion of the recovered gasoline and distillate being recycled to the sorption fractionation system.
  • Table IV shows the boiling range fraction composition for typical gasoline and distillate sorbents.
  • the sponge absorber may be constructed in a separate unit, as shown, or this operation may be conducted in an integral shell vessel with the main fractionation unit.
  • the rich sponge oil may be recovered from the upper contact zone as a separate stream, or the heavy distillate sorbent may be intermixed downwardly with gasoline sorbent and withdrawn from the bottom of the main fractionation zone.
  • the stream components of the olefinic feedstock and other main streams of the sorption/ prefractionator unit and reactor feedstreams are set forth in Table V, based on parts by weight per 100 parts of feedstock.
  • a typical byproduct of fluid catalytic cracking (FCC) units is an olefinic stream rich in C Z -C 4 olefins, usually in mixture with lower alkanes.
  • Ethylene can be recovered from such streams by conventional fractionation means, such as cryogenic distillation, to recover the C 2 and C 3 + fractions; however, the equipment and processing costs are high.

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  • Oil, Petroleum & Natural Gas (AREA)
  • Engineering & Computer Science (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • General Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)

Claims (6)

1. Verfahren zur Konvertierung eines olefinischen Ausgangsmaterials, das Ethylen und C3+-Olefine umfaßt, in ein schwereres flüssiges Kohlenwasserstoffprodukt, welches die Schritte umfaßt:
a) Vorfraktionierung des olefinischen Ausgangsmaterials, um einen ethylenreichen gasförmigen Strom und einen flüssigen Strom zu erhalten, der C3'-Olefine enthält,
b) Verdampfung des flüssigen Stromes vom Schritt a) und Kontakt des Dampfes mit einem Oligomerisierungskatalysator in zumindest einer exothermen Reaktionszone, um einen schwereren Kohlenwasserstoffabflußstrom zu schaffen, der Destillat, Benzin und leichtere Kohlenwasserstoffe umfaßt,
c) separate Abkühlung und Fraktionierung des Abflußstomes vom Schritt b), um Destillat, Benzin und leichtere Kohlenwasserstoffe zu gewinnen, und fach reactor i exchanged v
d) Rezirkulierung zumindest eines Teils des im Schritt c) gewonnenen Benzins als flüssiger Sorbtionsmittelstrom zum Vorfraktionierschritt a).
2. Verfahren nach Anspruch 1, welches den zusätzlichen Schritt umfaßt:
e) Leiten des heißen Abflusses aus der exothermen Reaktionszone im Schritt b) in indirektem Wärmeaustausch in einer Reboiler-Kreisleitung mit zumindest einem Teil des C3 +-Olefine enthaltenden flüssigen Stromes im Vorfraktionierschritt a).
3. Verfahren nach Anspruch 2, worin die C3 +-Olefine im Schritt b) in einer Reihe adiabatischer Festbettreaktoren bei erhöhtem Druck und bei einer Temperatur von 190 bis 315°C bei einer maximalen Temperaturerhöhung von etwa 30°C in jedem Reaktor reagieren, wobei der Abfluß von jedem Reaktor vor dem Eintritt in den nächsten Reaktor abgekühlt wird und der Abfluß von zumindest einem Reaktor in Wärmeaustausch mit dem flüssigen Vorfraktionierstrom gebracht wird, um die im Schritt e) sorbierten Kohlenwasserstoffe zu verdampfen.
4. Verfahren nach einem der Ansprüche 1 bis 3, worin der heiße Reaktorabfluß vom Schritt b) Leichtgas-, olefinische Cs'-Benzin-, und Kohlenwasserstoffkomponenten im Destillatsiedebereich enthält und fraktioniert wird, um diese Komponenten im Schritt c) zu trennen.
5. Verfahren nach einem der Ansprüche 2 bis 4, worin der heiße Abflußstrom vom Schritt b) nach dem wärmeaustausch mit dem flüssigen Vorfraktionierstrom verwendet wird, um einen Leichtgas-Entethaner aufzukochen.
6. Verfahren nach einem der Ansprüche 1 bis 5, worin der Oligomerisierungskatalysator einen sauren Zeolith vom Typ ZSM-5 umfaßt.
EP19840302911 1983-06-29 1984-05-01 Verfahren zur Konvertierung von Olefinen zu höheren Kohlenwasserstoffen Expired EP0130673B1 (de)

Applications Claiming Priority (4)

Application Number Priority Date Filing Date Title
US06/508,959 US4471147A (en) 1983-06-29 1983-06-29 Olefin fractionation and catalytic conversion system
US06/508,907 US4450311A (en) 1983-06-29 1983-06-29 Heat exchange technique for olefin fractionation and catalytic conversion system
US508907 1983-06-29
US508959 1983-06-29

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EP0130673A2 EP0130673A2 (de) 1985-01-09
EP0130673A3 EP0130673A3 (en) 1986-08-13
EP0130673B1 true EP0130673B1 (de) 1989-08-02

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EP (1) EP0130673B1 (de)
AU (1) AU574070B2 (de)
CA (1) CA1223279A (de)
DE (1) DE3479224D1 (de)
NZ (1) NZ208636A (de)

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* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4720600A (en) * 1983-06-29 1988-01-19 Mobil Oil Corporation Production of middle distillate range hydrocarbons by light olefin upgrading
US4547612A (en) * 1984-09-25 1985-10-15 Mobil Oil Corporation Production of lubricant and/or heavy distillate range hydrocarbons by light olefin upgrading
US4658073A (en) * 1985-09-23 1987-04-14 Mobil Oil Corporation Control system for multistage chemical upgrading

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Publication number Priority date Publication date Assignee Title
US2403879A (en) * 1944-03-10 1946-07-09 Phillips Petroleum Co Process of manufacture of aviation gasoline blending stocks
US4046830A (en) * 1975-09-18 1977-09-06 Mobil Oil Corporation Method for upgrading Fischer-Tropsch synthesis products
US4227992A (en) * 1979-05-24 1980-10-14 Mobil Oil Corporation Process for separating ethylene from light olefin mixtures while producing both gasoline and fuel oil
US4211640A (en) * 1979-05-24 1980-07-08 Mobil Oil Corporation Process for the treatment of olefinic gasoline
AU2455484A (en) * 1983-02-11 1984-08-16 Broken Hill Proprietary Company Limited, The Diesel fuel production from olefins

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AU2803384A (en) 1985-01-03
EP0130673A3 (en) 1986-08-13
CA1223279A (en) 1987-06-23
DE3479224D1 (en) 1989-09-07
EP0130673A2 (de) 1985-01-09
AU574070B2 (en) 1988-06-30
NZ208636A (en) 1987-11-27

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