CN104276911A - Process for oligomerizing dilute ethylene - Google Patents

Process for oligomerizing dilute ethylene Download PDF

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Publication number
CN104276911A
CN104276911A CN201410454385.9A CN201410454385A CN104276911A CN 104276911 A CN104276911 A CN 104276911A CN 201410454385 A CN201410454385 A CN 201410454385A CN 104276911 A CN104276911 A CN 104276911A
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weight
stream
catalyzer
catalyst
cracking catalyst
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C·P·尼古拉斯
A·巴塔查里亚
D·E·麦克维克
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Honeywell UOP LLC
Universal Oil Products Co
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Universal Oil Products Co
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Priority claimed from US12/416,032 external-priority patent/US8021620B2/en
Priority claimed from US12/416,029 external-priority patent/US8575410B2/en
Priority claimed from US12/416,026 external-priority patent/US8748681B2/en
Application filed by Universal Oil Products Co filed Critical Universal Oil Products Co
Publication of CN104276911A publication Critical patent/CN104276911A/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
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    • C07C2/00—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms
    • C07C2/02—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms by addition between unsaturated hydrocarbons
    • C07C2/04—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms by addition between unsaturated hydrocarbons by oligomerisation of well-defined unsaturated hydrocarbons without ring formation
    • C07C2/06—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms by addition between unsaturated hydrocarbons by oligomerisation of well-defined unsaturated hydrocarbons without ring formation of alkenes, i.e. acyclic hydrocarbons having only one carbon-to-carbon double bond
    • C07C2/08—Catalytic processes
    • C07C2/10—Catalytic processes with metal oxides
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    • B01J23/76—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of the iron group metals or copper combined with metals, oxides or hydroxides provided for in groups B01J23/02 - B01J23/36
    • B01J23/84—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of the iron group metals or copper combined with metals, oxides or hydroxides provided for in groups B01J23/02 - B01J23/36 with arsenic, antimony, bismuth, vanadium, niobium, tantalum, polonium, chromium, molybdenum, tungsten, manganese, technetium or rhenium
    • B01J23/85—Chromium, molybdenum or tungsten
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    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J38/00—Regeneration or reactivation of catalysts, in general
    • B01J38/04—Gas or vapour treating; Treating by using liquids vaporisable upon contacting spent catalyst
    • B01J38/12—Treating with free oxygen-containing gas
    • B01J38/30—Treating with free oxygen-containing gas in gaseous suspension, e.g. fluidised bed
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
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    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
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    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C11/00—Aliphatic unsaturated hydrocarbons
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    • C07C11/04—Ethylene
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    • C07C2/00—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms
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    • C07C2/04—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms by addition between unsaturated hydrocarbons by oligomerisation of well-defined unsaturated hydrocarbons without ring formation
    • C07C2/06—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms by addition between unsaturated hydrocarbons by oligomerisation of well-defined unsaturated hydrocarbons without ring formation of alkenes, i.e. acyclic hydrocarbons having only one carbon-to-carbon double bond
    • C07C2/08—Catalytic processes
    • C07C2/12—Catalytic processes with crystalline alumino-silicates or with catalysts comprising molecular sieves
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    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2/00—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms
    • C07C2/02—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms by addition between unsaturated hydrocarbons
    • C07C2/04—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms by addition between unsaturated hydrocarbons by oligomerisation of well-defined unsaturated hydrocarbons without ring formation
    • C07C2/06—Preparation of hydrocarbons from hydrocarbons containing a smaller number of carbon atoms by addition between unsaturated hydrocarbons by oligomerisation of well-defined unsaturated hydrocarbons without ring formation of alkenes, i.e. acyclic hydrocarbons having only one carbon-to-carbon double bond
    • C07C2/08—Catalytic processes
    • C07C2/24—Catalytic processes with metals
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    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/10—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with stationary catalyst bed
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    • 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
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
    • 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
    • C10G57/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one cracking process or refining process and at least one other conversion process
    • C10G57/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one cracking process or refining process and at least one other conversion process with polymerisation
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J21/00—Catalysts comprising the elements, oxides, or hydroxides of magnesium, boron, aluminium, carbon, silicon, titanium, zirconium, or hafnium
    • B01J21/12—Silica and alumina
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    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2521/00—Catalysts comprising the elements, oxides or hydroxides of magnesium, boron, aluminium, carbon, silicon, titanium, zirconium or hafnium
    • C07C2521/12—Silica and alumina
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    • C07C2523/16—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of arsenic, antimony, bismuth, vanadium, niobium, tantalum, polonium, chromium, molybdenum, tungsten, manganese, technetium or rhenium
    • C07C2523/24—Chromium, molybdenum or tungsten
    • C07C2523/30—Tungsten
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    • C07C2523/70—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of the iron group metals or copper
    • C07C2523/74—Iron group metals
    • C07C2523/755—Nickel
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2529/00—Catalysts comprising molecular sieves
    • C07C2529/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites, pillared clays
    • C07C2529/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • C07C2529/70—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of types characterised by their specific structure not provided for in groups C07C2529/08 - C07C2529/65
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    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/201—Impurities
    • C10G2300/207—Acid gases, e.g. H2S, COS, SO2, HCN
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    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
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    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/70—Catalyst aspects
    • C10G2300/708—Coking aspect, coke content and composition of deposits
    • 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
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/20—C2-C4 olefins
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/50—Improvements relating to the production of bulk chemicals
    • Y02P20/584—Recycling of catalysts

Abstract

The process and apparatus converts ethylene in a dilute ethylene stream that may be derived from an FCC product to heavier hydrocarbons. The catalyst may be an amorphous silica-alumina base with a Group VIII and/or VIB metal. The catalyst is resistant to feed impurities such as hydrogen sulfide, carbon oxides, hydrogen and ammonia. At least 40 wt-% of the ethylene in the dilute ethylene stream can be converted to heavier hydrocarbons.

Description

The oligomerization process of rare ethene
The name that the application is the applying date is on March 11st, 2010, application number is 201080023632.X, international application no is PCT/US2010/026889 is called the divisional application of the Chinese invention patent application of " oligomerization process of rare ethene ".
background of invention
The field of the invention is the apparatus and method by being become heavier hydrocarbon in hydrocarbon stream by the ethylene conversion diluted.These heavier hydrocarbon can be used as motor spirit.
Dry gas is the common name that the tail gas (off-gas) of fluid catalytic cracking unit flows, and comprises the gas of all boiling points lower than ethane.Compress this tail gas stream to remove C as much as possible 3and C 4gas.Sulphur is adsorbed in a large number from this tail gas stream in the washer adopting amine absorber.Institute's residual air stream is called FCC dry gas.Typical dry gas stream contains 5 ~ 50 % by weight ethene, 10 ~ 20 % by weight ethane, 5 ~ 20 % by weight hydrogen, 5 ~ 20 % by weight nitrogen, each carbon monoxide of 0.1 ~ 5.0 % by weight and carbonic acid gas, and being less than hydrogen sulfide and the ammonia of 0.01 % by weight, surplus is methane.
At present, FCC dry gas stream delivers to burner as fuel.Every day processes 7, and burning FCC unit every day of 949 kilolitres (50,000 barrel) is as 181,000kg (200 tons) dry gas of fuel and 36,000kg (40 tons) ethene.Because the price difference between fuel gas and motor spirit product or pure ethene is large, make great efforts to reclaim this ethene and seem favourable economically.But dry gas stream contains the impurity that oligomerisation catalyst can be made poisoning, and dry gas stream is so rare to such an extent as to unreasonable economically by gas recovery system recovered ethylene.
The oligomerization that concentrating ethylene flow to liquid product is known technology.But, oligomerization usually comprise use propylene or butylene, especially from liquefied petroleum gas (LPG) (LPG) or raw material dehydrogenation to obtain propylene or the butylene of the alkene of gasoline boiling range.Ethene, because of its much lower activity, is seldom used as the raw material of oligomerization.
Need the rare ethene utilized in refinery logistics.
summary of the invention
Have been found that with the group VIII on amorphous silica-alumina and/or group vib metal catalyst, can by as oligomeric in the ethylene catalyst in FCC dry gas stream for rare ethylene stream be heavier hydrocarbon.This heavier hydrocarbon can be separated and be mixed in gasoline and diesel oil storehouse.Have been found that the zeolite catalyst fast deactivation under impurity is as the existence of carbon oxides, ammonia and hydrogen sulfide being applicable to ethylene oligomerization reaction.These impurity not remarkably influenceds are included in the catalyzer of group VIII on amorphous silica-alumina and/or group vib metal.Therefore, the rare ethene in FCC dry gas stream oligomericly can become the liquid fuel product be easy to unconverted pneumatic separation.This unconverted gas can burn as fuel gas subsequently, and the ethene having more value shifts out as heavier hydrocarbon.
Advantageously, present method and device make it possible to utilize in diluent stream and under the raw material impurity of poisoning of catalyst can be made to exist ethene.
The feature that the present invention is other and advantage become obvious by putting forward claim in detailed description of the present invention, accompanying drawing and literary composition.
accompanying drawing is sketched
Fig. 1 is the schematic diagram of FCC unit and fcc product recovery system.
Fig. 2 is the graphic representation of conversion of ethylene to the production time of embodiment 6 ~ 8.
Fig. 3 is the graphic representation of conversion of ethylene to the production time of embodiment 9.
Fig. 4 is the graphic representation of conversion of ethylene to the production time of embodiment 10.
detailed Description Of The Invention
The present invention can be used for any hydrocarbon stream containing ethene and the preferably dilution part of ethene.Suitable rare ethylene stream typically can comprise 5 ~ 50 % by weight ethene.FCC dry gas stream is suitable rare ethylene stream.Other rare ethylene streams also can be used in the present invention, as coker dry gas stream.Because the present invention is particularly suitable for FCC dry gas, by using about the ethrel from FCC dry gas stream, the present invention is described.
Translate into Fig. 1 now, wherein identical Reference numeral indicates similar component.Fig. 1 represents refinery facilities 6, generally includes FCC elementary section 10 and Product recycling section 90.FCC elementary section 10 comprises reactor 12 and catalyst regenerator 14.Processing parameter generally includes cracking reaction temperature 400 ° ~ 600 DEG C and catalyst regeneration temperature 500 ° ~ 900 DEG C.Cracking and regeneration are all carried out under the absolute pressure lower than 506kPa (72.5psia).
Fig. 1 illustrates typical FCC reactor 12, and wherein, the heavy hydrocarbon feedstocks in divider 16 or crude stream contact with the Regenerated cracking catalyst entered from regenerated catalyst riser 18.This contact can be carried out in narrow riser 20, and this riser extends up to the bottom of reactor vessel 22.The contact of raw material and catalyzer is carried out the fluidizing gas of self-fluidized type pipeline 24.In one embodiment, when the two along riser upwards shift enter reactor vessel 22 time, from catalyzer heat by hydrocarbon feed or oil evaporation, this hydrocarbon feed is cracked into the hydrocarbon product of lighter molecular weight subsequently in the presence of a catalyst.Also carry out inevitable side reaction in riser 20, by sedimentation of coke on a catalyst, reduce the activity of catalyzer.Be separated with the cracking catalyst of coking by cracking lighter hydrocarbons product with cyclonic separator subsequently, this separator can be included in initial separator 26 in reactor vessel 22 and one-level or two stage cyclone device 28.The crackate of gaseous state leaves reactor vessel 22 to pipeline 32 through product exit 31, is transferred to the product recovery section 90 in downstream.Catalyzer that is useless or coking needs regeneration to use further.After the cracking catalyst of coking is separated with gaseous product hydrocarbon, fall into stripping stage 34, this through nozzle injecting steam to remove the hydrocarbon vapour of any remnants.After stripping operation, the catalyzer of coking is delivered to catalyst regenerator 14 by the riser 36 via useless catalyzer.
Fig. 1 describes the revivifier 14 being called burner.But the revivifier of other types is also suitable.In catalyst regenerator 14, by oxygen flow as air is introduced through air distributor 38, to contact the catalyzer of coking.From the catalyzer burn off coke of coking to provide catalyzer and the flue gas (flue gas) of regeneration.Catalyst regeneration process with the addition of a large amount of heat to catalyzer, provides energy with the endothermic cracking reaction occurred in compensatory reactionBu Changfanying device riser 20.Catalyzer upwards flows along the burner riser 40 being arranged in catalyst regenerator 14 together with air, after regeneration, by discharging and initial gross separation through separator 42.The further recovery of the regenerated catalyst and flue gas that leave separator 42 adopts first and second grades of cyclonic separators 44 in catalyst regenerator 14 respectively, and 46 realize.The catalyzer be separated with flue gas is through dipping tube (dipleg) from cyclonic separator 44, and 46 distribute, and flue gas relatively light in catalyzer leaves cyclonic separator 44 immediately, and 46 and exhanst gas outlet 47 in flue gas pipeline 48 leaves regenerator vessel 14.Regenerated catalyst sends riser 20 back to through regenerated catalyst riser 18.As the result of burning, the flue gas steam left at catalyst regenerator 14 top in pipeline 48 comprises CO, CO 2, N 2and H 2o, with other materials of small amount.The exhanst gas outlet 47 of heat smoke in pipeline 48 leaves revivifier 14 to process further.
Product recovery section 90 and product exit 31 communicating downstream." communicating downstream " refers to that the component that at least part of material of certain component flowing to communicating downstream operationally can be communicated with from this component flows out." connection " refers to that the upper material that allows of operation flows between cited component.In product recovery section 90, the gaseous state fcc product in pipeline 32 leads to the low side of FCC main fractionating tower 92.King-tower 92 and product exit 31 communicating downstream.Can be separated and comprise from some cuts of the fcc product of king-tower taking-up: starching from the heavy oil of bottoms pipeline 93, heavy cycle oil stream in pipeline 94, take from the light cycle oil of outlet 95a in pipeline 95, in pipeline 96, take from the heavy naphtha stream of outlet 96a.Arbitrary or whole pipeline 93 ~ 96 can be cooled and usual at higher position blowback king-tower 92 to cool king-tower.Gasoline and gaseous state lighter hydrocarbons are shifted out from king-tower 92 and makes it condensation before entering king-tower susceptor 99 in overhead line 97.King-tower susceptor 99 and product exit 31 communicating downstream, king-tower 92 is communicated with king-tower susceptor 99 upstream." upstream connection " refers to that at least part of material that the component be communicated with from upstream flows out operationally can flow to the component be communicated with this component.
Current shift out from the susceptor (boot) susceptor 99.Further, the light naphthar stream of condensation shifts out and overhead streams shifts out in pipeline 102 in pipeline 101.Overhead streams in pipeline 102 comprises gaseous state lighter hydrocarbons, and it can comprise rare ethylene stream.Stream in pipeline 101 and 102 can enter the vapor recovery section 120 of product recovery section 90.
Vapor recovery section 120 is illustrated as the system based on absorption, but any vapor recovery system can use, and comprises cold box system.For realizing the abundant separation of lighter-than-air gas component, the gaseous stream in pipeline 102 is compressed in compressor 104.The compressor more than one-level can be adopted, but usually adopt Two-stage Compression.Stream in pipeline 107 and 108 is incorporated in light hydrocarbon stream compressed in pipeline 106, cools and delivers to high pressure receiver 110.Current from susceptor 110 can be sent to king-tower susceptor 99.The gaseous state hydrocarbon stream comprising rare ethylene stream in pipeline 112 is sent to elementary adsorber 114, contacts wherein, to carry out C with the stable gasoline that do not add from king-tower susceptor 99 in pipeline 101 3+ and C 2separation between-hydrocarbon.Elementary adsorber 114 and king-tower susceptor 99 communicating downstream.Liquid C in pipeline 107 3+ flow return line 106 before cooling.The rare ethylene stream for the object of the invention is comprised from the elementary tail gas stream of elementary adsorber 114 in pipeline 116.But reclaim heavy component for further concentrating ethylene stream, not necessarily pipeline 116 can be led to secondary adsorption device 118, at this, the cycling stream of the light cycle oil shifted from pipeline 95 in pipeline 121 absorbs the overwhelming majority residue C in elementary tail gas stream 5+ and some C 3~ C 4material.Secondary adsorption device 118 and elementary adsorber 114 communicating downstream.By in pipeline 119 from bottom secondary adsorption device, be rich in C 3the light cycle oil of+material sends king-tower 92 back to through the backflow (pump-around) of pipeline 95.The overhead materials of secondary adsorption device 118 comprises dry gas, and it is mainly C 2-hydrocarbon, with hydrogen sulfide, ammonia, carbon oxides and hydrogen, this overhead materials shifts out from secondary tail gas stream in pipeline 122, to form rare ethylene stream.
Stripping tower 126 is delivered to from the liquid of high pressure receiver 110 by pipeline 124.Most C 2-shift out at the top of stripping tower 126, and through overhead line 108 return line 106.The bottoms liquid of stripping tower 126 is flowed through pipeline 128 and delivers to debutanizing tower 130.Overhead streams from debutanizing tower in pipeline 132 comprises C 3~ C 4olefin product, and the bottoms stream comprised in pipeline 134 through stable gasoline can process further and deliver to gasoline basin.
Rare ethylene stream of the present invention can comprise FCC dry gas stream, and this dry gas stream comprises the ethene between 5 to 50 % by weight, preferably the ethene of 10 ~ 30 % by weight.Main ingredient in the normally rare ethylene stream of methane, concentration is between 25 to 55 % by weight, and ethane also exists in a large number, usually between 5 to 45 % by weight.The hydrogen and the nitrogen that to be between 1 to 25 % by weight and to be generally 5 ~ 20 % by weight can be there is in rare ethylene stream.The water of each species saturation may reside in rare ethylene stream.If use secondary adsorption device 118, C 3+ existence will be no more than 5 % by weight, with the propylene being usually less than 0.5 % by weight.
In addition to hydrogen, other impurity such as hydrogen sulfide, ammonia, carbon oxides and acetylene also may reside in rare ethylene stream.
Have been found that the many impurity in dry gas ethylene stream all can make oligomerisation catalyst poisoning.Hydrogen and carbon monoxide meeting position, reducing metal are to inactivation.Carbonic acid gas and ammonia can encroach on the acidic site on catalyzer.The metal that hydrogen sulfide can corrode on catalyzer generates metallic sulfide.Acetylene can be polymerized and gum deposit on catalyzer or equipment.
The secondary exhaust stream comprising rare ethylene stream in pipeline 122 is introduced nonessential amine absorbing unit 140, to remove hydrogen sulfide to lower concentration.The poor amine aqueous solution such as comprising monoethanolamine or diethanolamine is introduced adsorber 140 through pipeline 142, contacts with adsorption of hydrogen sulfide with the secondary exhaust stream of flowing; Rich amine adsorption aqueous solution containing hydrogen sulfide shifted out adsorption zone 140 through pipeline 143 and reclaims, also may process further.
Rare ethylene stream through amine process in pipeline 144 can be introduced nonessential washing unit 146, carry from amine adsorber 140 residual amine that comes secretly to remove and reduce the concentration of ammonia and carbonic acid gas in rare ethylene stream pipeline 144.By in the water scrubber in water introduction pipe line 145.Usually by slight for the water in pipeline 45 acidifying with strengthen to the trapping of alkali molecules as amine.In pipeline 147, the current of enriched amine and possible ammonia and carbonic acid gas leave washing unit 146, and can process further.
Then can may process to remove one or more impurity, under carbon monoxide, hydrogen sulfide and ammonia to lower concentration in nonessential guard bed 150 through the stream of rare ethene of washing through nonessential amine process by pipeline 148.Guard bed 150 impurity that oligomerisation catalyst can be able to be made poisoning with absorption containing sorbent material, as hydrogen sulfide.Guard bed 150 can containing multiple sorbent material to adsorb more than a kind of impurity.The typical adsorption agent of adsorption of hydrogen sulfide is ADS-12, CO absorption for ADS-106, adsorb ammonia be UOP MOLSIV 3A, all can obtain from UOP, LLC.Sorbent material may be combined in single bed or is arranged in the bed of order.
In pipeline 151 may through amine process, will may may usually have the following impurity concentration of at least one through adsorption treatment with the rare ethylene stream removed in more hydrogen polysulfide, ammonia and carbon monoxide through washing: 0.1 % by weight to maximum carbon monoxide of 5.0 % by weight and/or 0.1 % by weight to maximum carbonic acid gas of 5.0 % by weight, and/or at least 1 weight ppm to the hydrogen sulfide of maximum 500 weight ppm and/or the ammonia of at least 1 to maximum 500 weight ppm, and/or at least 5 to maximum hydrogen of 20 % by weight.Kind and the concentration thereof of existing impurity can change, and depend on the initial composition for the treatment of process and rare ethylene stream.
Rare ethylene stream is delivered to compressor 152 and is forced into reactor pressure by pipeline 151.Compressor 152 and king-tower 92, product recovery section 90 and product exit 31 communicating downstream.Compressed rare ethylene stream can be compressed at least 3,550kPa (500psia), 10,445kPa (1500psia) may be no more than, be suitable for 4, between 930kPa (700psia) Yu 7,687kPa (1100psia).Preferred rare ethylene stream is compressed into the emergent pressure higher than ethene, is 4,992kPa (724kPa), to avoid rapid catalyst deactivation to this emergent pressure of pure ethylene.Compressor 152 can comprise one-level or multistage with interstage cooling.May need well heater that compressed stream is heated to temperature of reaction.Compressed rare ethene delivers to oligomerization reactor 156 in pipeline 154.
Oligomerization reactor 156 respectively with compressor 152 and primary and secondary adsorber 114 and 118 communicating downstream.Oligomerization reactor is preferably containing stationary catalyst bed 158.Rare ethylene feed streams preferably fall liquid operation in catalyst exposure.But upper flow operation is also suitable.Catalyzer is preferably the metal with group VIII and/or group vib in the periodic table of elements of employing Chemical Abstracts Service labelling method in amorphous silica-alumina substrate.On the one hand, catalyzer has by the metal promoted group VIII metal of group vib.On the one hand, catalyzer have be no more than 30, the ratio of silica-alumina preferably more than 20.Usually, silicon oxide and aluminum oxide exist only in substrate, being therefore compared to catalyzer and being identical for substrate of silica-alumina.Metal can flood or ion-exchange in silica-alumina substrate.Grinding is also expected altogether.At least 0.15 can be had for catalyzer of the present invention, be suitably 0.2, be preferably greater than 0.25 low temperature acidity ratio (Low Temperature Acidity Ratio), ammonia temperature programmed control desorption (Ammonia Temperature Programmed Desorption, ammonia TPD) surveyed as mentioned below.In addition, suitable catalyzer has 50 and 400m 2surface-area between/g, as nitrogen BET survey.
Preferred oligomerisation catalyst of the present invention is as described below.Preferred oligomerisation catalyst comprises amorphous silica-alumina carrier.One of component of support of the catalyst used in this invention is aluminum oxide.Aluminum oxide can be hydrous oxide or the alumina sol of any different aluminium, as a hydration α-aluminum oxide, three hydration α-aluminum oxides of gibbsite structure, three hydration βAl2O3 and the analogues of bayerite structure of boehmite or pseudo-boehmite structure.Particularly preferred aluminum oxide can obtain from Sasol North America Alumina Ptroduct Group with trade name Catapal.This material is a hydration α-aluminum oxide (pseudobochmite) of extreme high purity, and it generates high-purity gamma aluminum oxide after high-temperature calcination.Another component of support of the catalyst is amorphous silica-alumina.The ratio of silica-alumina be 2.6 appropriate oxidation silicon-aluminum oxide can obtain from CCIC, it is the Subsidiary Company of JGC, Japan.
Another component be used in the preparation of used catalyst of the present invention is tensio-active agent.Tensio-active agent preferably mixes with above-mentioned aluminum oxide and silica-alumina powder.Subsequently that the mixture of gained tensio-active agent, aluminum oxide and silica-alumina is shaping, dry as mentioned below and calcine.Remove the organic composition of tensio-active agent with being calcined through burning highly effective, but only tensio-active agent fully completed its according to function of the present invention after.Any suitable tensio-active agent all can be used according to the invention.Preferred tensio-active agent is the tensio-active agent being selected from the commodity surfactants that Solvay S.A. sells with trade(brand)name " Antarox "." Antarox " tensio-active agent is commonly called modified linear aliphatic polyether, and is the biodegradable washing composition of low bubble and wetting agent.
By being mixed into the silica-alumina of the volume of ratio and aluminum oxide and preparing suitable silica-alumina mixture, to obtain the ratio of the silica-alumina of expectation.In one embodiment, the ratio of the 85 % by weight silica-aluminas amorphous silica-alumina that is 2.6 and 15 % by weight alumina powder will provide suitable carrier.In one embodiment, the proportioning of amorphous silica-alumina and aluminum oxide is not 85 ~ 15 also can be suitable, as long as the ratio of the final silica-alumina of carrier is suitably and is no more than 30 and preferably more than 20.
Any method easily all can be used for the mixture of tensio-active agent and silica-alumina and aluminum oxide to merge.Mixing and mixed surfactant between Formation period preferably at aluminum oxide and silica-alumina.Preferred method is mixed by the blend of the aqueous solution of tensio-active agent with aluminum oxide and silica-alumina before carrier is finally formed.Preferred surfactant is present in this lotion or dough with the amount of 0.01 ~ 10 % by weight, based on the weight of aluminum oxide and silica-alumina.
Single protonic acid in the aqueous solution can be added as nitric acid or formic acid, with the aluminum oxide in peptization binding agent in mixture.Extra water can be added to provide fully wetting, to build the enough dough of denseness to extrude or spraying dry in this mixture.
This lotion or dough can be prepared with the shape of shaped granule, preferred method is the mouth mould doughy mix of aluminum oxide, silica-alumina, tensio-active agent and water being extruded through the perforate with intended shape and size, and the extrudate subsequently extrudate being broken into desired length is also dry.Calcining further can be adopted to gain in strength to extrudate.Usually, calcine in flow of dried air, carry out at the temperature of 260 DEG C (500 °F) to 815 DEG C (1500 °F).
Extrude particle and can have any suitable shape of cross section, namely symmetrical or asymmetrical, but great majority usually have symmetric cross-sectional shape, and preferably spherical, cylindrical or multi-leaf-shaped.The cross-sectional diameter of particle can be little of 40 μm, but, be generally 0.635mm (0.25 inch) to 12.7mm (0.5 inch), preferred 0.79mm (1/32 inch) to 6.35mm (0.25 inch), most preferably 0.06mm (1/24 inch) to 4.23mm (1/6 inch).Wherein, it is the similar trifolium of shape of cross section that preferred catalyzer is arranged, such as, as US4, shown in Fig. 8 and Tu A of 028,227.Preferred cloverleaf pattern particle is such: each " leaf " of cross section defined by 270 ° of circular arcs of the circle of diameter between 0.51mm (0.02 inch) and 1.27mm (0.05 inch).Other preferred particles are the quadrifoliate shape of cross sections of those tools, comprise asymmetric shape and symmetric figure, as US4, shown in Figure 10 of 028,227.
In literary composition, the characteristic feature of amorphous silica-alumina carrier used is total pore volume, mean pore size and large must being enough to provide the sufficient room of position activity metal component and the surface-area of area.The total pore volume of this carrier, as conventional mercury gaging hole method surveyed, be generally 0.2 ~ 2.0cc/ gram, preferably 0.25 ~ 1.0cc/ gram, most preferably 0.3 ~ 0.9cc/ gram.Usually, in pore volume amount, the hole that diameter is greater than 100 dusts is less than 0.1cc/ gram, is preferably less than 0.08cc/ gram, more preferably less than 0.05cc/ gram.Surface-area, as B.E.T. method surveyed, be generally higher than 50m 2/ gram, such as, higher than 200m 2/ gram, preferably at least 250m 2/ gram, most preferably 300m 2/ gram to 400m 2/ gram.
For Kaolinite Preparation of Catalyst, by solid support material compound, such as by by least one from one or more precursor one-dip of periodictable group VIII or group vib metal component and the multistep dipping amorphous refractory oxide carrier granule through calcining.Group VIII metal, preferred nickel, should exist with the concentration of 0.5 ~ 15 % by weight; Group vib metal, preferred tungsten, should exist with the concentration of 0 ~ 12 % by weight.Dipping can be undertaken by any method known in the art, such as, by spray impregnation, is wherein sprayed onto on carrier granule by the solution of the metal precursor containing solubilized form.Another kind method is repeatedly dipping step, wherein makes carrier granule repeatedly contact with dipping solution, with or without intermittently drying.Another method comprises in dipping solution carrier being immersed in large volume or carrier is circulated wherein.Another kind method is pore capacities or hole wetness technique, and wherein carrier granule being introduced volume is just enough in the dipping solution in the hole being full of carrier.Sometimes, can change hole wetness technique to adopt compared to being just full of hole, volume is being less than it 10% to being greater than its dipping solution between 10%.
If active metal precursor.In is combined by dipping, then subsequently or the calcining of second time at elevated temperature such as between 399 ° to 760 DEG C (750 ° to 1400 °F) convert metals will be become its respective oxide form.Under some situation, all calcine after each dipping of each active metal.Calcining subsequently generates the catalyzer containing its active metal of oxide form separately.
The preferred oligomerisation catalyst of the present invention have with 0.5 ~ 15 % by weight nickel dipping amorphous silica-alumina substrate, for 3.175mm (0.125 inch) extrudate form and density is 0.45 ~ 0.65g/ml.Also it is expected to metal can by other methods combining on carrier, such as ion-exchange and common grinding.
Be applicable to other catalyzer of the present invention and adopt cogelled silicaalumina carrier, prepared by known oil drop method, the method allows to use carrier spherical greatly.Such as, alumina sol, as alumina source, mixes with the acidified sodium silicate solution as silicon source, then mixes further with suitable jelling agent such as urea, vulkacit H or its mixture.By this mixture still remaining in the hot oil bath of gelling temp lower than being discharged into by nozzle or rotating disk during gelling temp.This mixture enters oil bath with droplet form dispersion, and walks the spherical gel particle of period formation wherein.Alumina sol is prepared preferably by this method, wherein by aluminium bead with processed in a large number or the water of deionization mix, add hydrochloric acid wherein with the amount being enough to melt part aluminum metal, form the colloidal sol expected.Suitable speed of reaction is realized at the reflux temperature of mixture.
Spherical gel particles oil drop method prepared is aging in this wise, namely usual in oil bath, the time of at least 10 ~ 16 hours, then in suitable alkali or alkaline medium at least 3 ~ 10 hours, finally washes.In oil bath, the suitable gelling of mixture and the aging subsequently of gel ball are difficult to carry out under lower than 48.9 DEG C (120 °F), and under 98.9 DEG C (210 °F), the quick release of gas is easy to make globules rupture or weaken spheroid.By being formed and keeping enough superatmospheric pressure so that water is remained liquid phase during Aging Step, higher temperature can be adopted, usually there is the result of improvement.If gel particle is aging under superatmospheric pressure, then without the need to alkali Aging Step.
Spheroid is washed, preferably with the water containing a small amount of ammonium hydroxide and/or ammonium nitrate.After washing, by spheroid drying 6 to 24 hours or more of a specified duration under 93.3 DEG C (200 °F) to the temperature of 315 DEG C (600 °F), then at the temperature lower calcination 2 to 12 hours or more of a specified duration of 426.67 DEG C (800 °F) to 760 DEG C (1400 °F).
By any suitable common dipping technique by group VIII component and VIB component and cogelled silicaalumina carrier Material cladding.Thus, solid support material can be soaked, dipping, suspend or be immersed in the aqueous impregnation solution containing group VIII soluble salt and group vib soluble salt.A kind of suitable method comprises and is immersed in dipping solution by solid support material, in rotary drum steam dryer, be evaporated to drying, and it is 0.1 ~ 0.3 that the concentration of dipping solution guarantees that final catalyst composition comprises the atomic ratio that nickel and nickel adds tungsten.Another kind of suitable method comprises and at room temperature immerses in aqueous impregnation solution solid support material until realize the completely infiltration of solution to carrier.After absorbing dipping solution, by carrier venting surface free fluid, and dry in mobile belt calciner.
Usually by catalyst complex before calcination in 93.3 DEG C (200 °F) to 260 DEG C of (500 °F) temperature under dry 1 ~ 10 hour.According to the present invention, calcine and carry out to the temperature of 650 DEG C (1200 °F) at 371 DEG C (700 °F) in oxidizing atmosphere.Oxidizing atmosphere is suitably air, although can adopt other gases comprising molecular oxygen.
Suitable other catalyzer is the oil drop method silica-alumina ball type carrier of diameter 3.175mm (0.125 inch), it is soaked with 0.5 ~ 15 % by weight nickel and 0 ~ 12 % by weight tungsten.Other melts combine methods are also suitable and it is expected to.The Suitable Density scope of other catalyzer can between 0.60 and 0.70g/ml.
Rare ethylene raw can contact with oligomerisation catalyst under the temperature between 200 DEG C with 400 DEG C.Reaction is main in the gas phase based on ethene GHSV 50 ~ 1000hr -1under carry out.Be surprisingly found out that, make the impurity of poisoning of catalyst and in raw material, ethene is diluted although exist in raw material, in feedstream at least 40 % by weight and many ethylene conversion as 75 % by weight become heavier hydrocarbon.Ethene is polymerized to heavier alkene by low first on a catalyst.Some heavier alkene can cyclisation on a catalyst, and the existence of hydrogen can promote that alkene is to the conversion of paraffinic hydrocarbons, and paraffinic hydrocarbons is all the hydrocarbon heavier than ethene.
Although impure in charging, catalyzer still keeps stable, can be regenerated during inactivation.Suitable regeneration condition comprises makes catalyzer, such as original position, experiences the warm air 3 hours of 500 DEG C.Activity and selectivity and the live catalyst of regenerated catalyst are suitable.
Oligomerization product stream from oligomerization reactor in pipeline 160 can be delivered to oligomeric separator 162, it can be simple flash drum, with separation bubble and liquid stream.Oligomeric separator 162 and oligomerization reactor 156 communicating downstream.Fuel element 166 can be delivered to generate the steam in pipeline 167 by comprising in overhead line 164 gaseous product stream of light gas as hydrogen, methane, ethane, unreacted alkene and light impurities.Or, can by the gaseous product burning in overhead line 164 to give well heater (not shown) for fire and/or to provide flue gas source to make gas-turbine (not shown) rotary electrification.Overhead line 164 is communicated with fuel element 166 upstream.Can make be downward through valve from oligomeric separator 162, the residue liquid stream that comprises heavier hydrocarbon and loop back product separation section 90 in pipeline 168.Bottom 169 communicating downstream of circulation line 168 and oligomeric separator 162.Like this, king-tower 92 is communicated with oligomerization reactor 156 upstream and downstream.Residual logistics loops back king-tower 92 preferably by circulation line 168 at heavy naphtha outlet 96a and the light cycle oil position exported between 95a.Or circulation line 168 is light cycle oil pipeline 95 or heavy naphtha pipeline 96 feed.Circulation line is with oligomerization reactor 156 communicating downstream, be communicated with king-tower 92 upstream.Or the oligomerization product in pipeline 160 or 168 can be saturated or undersaturated, can not loop back product separation district 90 and deliver to tanks.
Embodiment
Application of the present invention is illustrated by following examples.
Embodiment 1
The spherical suprabasil nickel of oil drop method amorphous silica-alumina and tungsten are by synthesizing the step given by the alternative catalyzer of the present invention above.Metal of the present invention comprises the nickel of 1.5 % by weight and the tungsten of 11 % by weight.Spherical substrate has the diameter of 3.175mm.The ratio of the silica-alumina of catalyzer is 3, and density is 0.641g/ml, and surface-area is 371m 2/ g.
Embodiment 2
The amorphous silica-alumina extruded is combined with the pseudobochmite provided with trade name Catapal by the amorphous silica-alumina that the ratio of the silica-alumina provided by CCIC with weight ratio 85 ~ 15 is 2.6 and synthesizes.Before mixing with amorphous silica-alumina, with nitric acid by pseudobochmite peptization.The water of the tensio-active agent provided with trade name Antarox and the amount being enough to wetting dough is provided to this mixture.At being 550 DEG C before calcining, catalyzer group being extruded through the 1.59mm perforate on cylindrical mouth template and being broken into block.Final catalyzer is made up of 85 % by weight silica-aluminas and 15 % by weight aluminum oxide, and the ratio of silica-alumina is 1.92 and surface-area is 368m 2/ g.
Embodiment 3
By Ni (NO 3) 26H 2o is dissolved in 32.08 grams of deionized waters.By four times add nickel solutions and interpolation between concuss, the amorphous silica-alumina that nickel solution and embodiment 2 are extruded contacts.Obtain jade-green extrudate.Then, before being cooled to room temperature, keeping 500 DEG C calcining for 3 hours by extrudate dry at 110 DEG C 3 hours, afterwards with 2 DEG C/m temperature programming to 500 DEG C, is oxide form by nickel convert metals.Form light grey extrudate, containing the nickel of 1.5 % by weight.
Embodiment 4
From the MTT zeolite sample that the ratio of Zeolyst Corporation acquisition silica-alumina is 40.MTT zeolite is combined with pseudobochmite, at 550 DEG C before calcining, extrudes it through the perforate of 3.175mm on cylindrical mouth template.Final catalyzer is made up of 80 % by weight MTT zeolites and 20 % by weight aluminum oxide.
Embodiment 5
In fixed bed on 10mL catalyzer, 280 DEG C, the olefin oligomerization of testing example 1 catalyzer under 6,895kPa (1000psig), 58OGHSV (olefin gas space-time speed).Raw material is by 30 % by weight C 2h 4with 70 % by weight CH 4form.Result is shown in Table I.
Embodiment 6
In fixed bed on 10mL catalyzer, 280 DEG C, the olefin oligomerization of testing example 2 catalyzer under 6,895kPa (1000psig), 586 O GHSV.Raw material is by 23 % by weight C 2h 4, 14 % by weight C 2h 6, and 35 % by weight CH 4, 13 % by weight H 2, 13 % by weight N 2, 1 % by weight CO, 1.5 % by weight CO 2, 10 weight ppm H 2s is formed, and with 25 DEG C, the steam-laden of 3,447kPa (500psig) before feed is to oligomerization.Result is shown in Table I and Fig. 2.
Embodiment 7
In fixed bed on 10mL catalyzer, 280 DEG C, the olefin oligomerization of testing example 3 catalyzer under 6,895kPa (1000psig), 586OGHSV.Raw material is by 23 % by weight C 2h 4, 14 % by weight C 2h 6, and 35 % by weight CH 4, 13 % by weight H 2, 13 % by weight N 2, 1 % by weight CO, 1.5 % by weight CO 2, 10 weight ppm H 2s is formed, and with 25 DEG C, the steam-laden of 3,447kPa (500psig) before feed is to oligomerization.Result is shown in Table I and Fig. 2.Produce 27 ~ 44 hours periods, also add the NH of 1ppm to raw material 3.Find transformation efficiency and selectivity unchanged.
Embodiment 8
Except H in raw material 2the concentration of S is outside 50 weight ppm instead of 10 weight ppm, repeats the experiment of embodiment 7.Result is shown in Table I and Fig. 2.
Fig. 2 is the C of embodiment 6 ~ 8 2h 4transformation efficiency is to the graphic representation of production time.With regard to conversion of ethylene, the performance of the catalyzer of embodiment 3 nickel on amorphous silica-alumina is better than the substrate that embodiment 2 is only silica-alumina.The catalyzer of embodiment 2 and 3 is subject to the impact of raw material impurity all hardly.
Embodiment 9
In fixed bed on 10mL catalyzer, 280 DEG C, the olefin oligomerization of testing example 4 catalyzer under 6,895kPa (1000psig), 586OGHSV.Raw material is by 23 % by weight C 2h 4, 14 % by weight C 2h 6, and 35 % by weight CH 4, 13 % by weight H 2, 13 % by weight N 2, 1 % by weight CO, 1.5 % by weight CO 2, 10 weight ppm H 2s is formed, and with 25 DEG C, the steam-laden of 3,447kPa (500psig) before feed is to oligomerization.Result is shown in Table I and Fig. 3.
Fig. 3 is the C of embodiment 9 2h 4transformation efficiency to the graphic representation of production time, the impact of the impurity showing to make the MTT zeolite catalyst of embodiment 4 poisoning.After reacting 20 hours, transformation efficiency is brought down below 10 % by weight, and in embodiment 7 and 8, the transformation efficiency of the catalyzer of embodiment 3 remains on about 60 % by weight.
Embodiment 10
In fixed bed on 10mL catalyzer, 280 DEG C, the olefin oligomerization of testing example 4 catalyzer under 6,895kPa (1000psig), 613OGHSV.Raw material is by 30 % by weight C 2h 4with 70 % by weight CH 4form.Produce 21 constantly little, add hydrogen, obtain by 27 % by weight C 2h 4, 63 % by weight CH 4with 10 % by weight H 2the raw material formed.Produce 45 constantly little, add H 2in the NH of 500 weight ppm 3, obtain by 27 % by weight C 2h 4, 63 % by weight CH 4, 10 % by weight H 2with 500 weight ppmNH 3the raw material formed.Result is shown in Table I and Fig. 4.
Fig. 4 is the graphic representation of conversion of ethylene to the production time of embodiment 10, shows impurity H 2and NH 3on the impact of the MTT zeolite catalyst of embodiment 4.As seen in fig. 3, decline to conversion of ethylene during raw material introducing hydrogen constantly when production 21 is little.In addition, when production 45 little introducing constantly ammonia, conversion of ethylene promptly significantly declines.
Table I
Embodiment Transformation efficiency C 2-C 4Selectivity C 5-C 10Selectivity C 10+ selectivity
Embodiment 5 92 30 53 17
Embodiment 6 38 8 34 58
Embodiment 7 75 9 18 73
Embodiment 8 74 9 17 74
Embodiment 9 ? ? ? ?
Produce 5 hours 77 1 31 68
Produce 20 hours 7 10 35 55
Embodiment 10 ? ? ? ?
Produce 10 hours 90 1 31 68
Produce 35 hours 65 4 31 65
Produce 45 hours 63 4 31 65
Produce 70 hours 30 6 24 70
Embodiment 11
Ammonia temperature programmed control desorption (ammonia TPD) test comprises: the temperature 250 milligrams of catalyst samples being first heated in the helium-atmosphere of existence 20 volume % oxygen, flow velocity per minute 100 milliliters 550 DEG C with the speed of per minute 5 DEG C.After keeping 1 hour, by helium purge system 15 minutes, sample is cooled to 150 DEG C.Then in the helium of per minute 40 milliliters with ammonia pulse by saturated for this sample.The total amount of ammonia used substantially exceeds and makes all saturated required amount of all acidic sites on sample.With this sample of helium purge 8 hours of per minute 40 milliliters to remove the ammonia of physical adsorption (physiosorb).Along with the carrying out of helium purge, temperature is elevated to the outlet temperature of 600 DEG C with the speed of per minute 10 DEG C.By the ammonia amount of calibration thermal conductivity detector monitoring desorption.Ammonia total amount is determined by integration.
The ammonia total amount of desorption obtains total acidity with the ratio of sample dry weight.As used herein, the numerical value of total acidity provides with the unit of mmole ammonia every gram of dry sample.Have for acidity to the activated catalyzer of rare ethylene stream oligomerization tool, namely there is the total acidity of at least 0.15, preferably at least 0.25, as ammonia TPD survey.
Low temperature peak is obtained from the ammonia total amount of sample desorption and the ratio of sample dry weight before reaching the temperature of 300 DEG C.As used herein, the numerical value of low temperature peak provides with the unit of mmole ammonia every gram of dry sample.To the activated catalyzer of rare ethylene stream oligomerization tool, there is such low temperature peak, namely have at least 0.05, preferably at least 0.06 low temperature peak, as ammonia TPD survey.
Low temperature peak and the ratio of total acidity obtain being called low temperature acidity ratio without unit ratio.Tolerate raw material impurity in rare ethylene stream and have at least 0.15 to the activated catalyzer of rare ethylene stream oligomerization tool, aptly at least 0.2, be preferably more than 0.25 low temperature acidity ratio, as ammonia TPD survey.
Table II
As can be seen from embodiment, with regard to conversion of ethylene, by impurity effect in raw material, efficiency reduces zeolite catalyst greatly; Although and there is the impurity usually making poisoning of catalyst in raw material, catalyzer of the present invention keeps efficient ethylene oligomerization catalyst.Catalyzer of the present invention keeps the transformation efficiency of at least 40 % by weight, is generally 60 % by weight, preferably more than 70 % by weight.
Without the need to being described in further detail, it is believed that those skilled in the art farthest can utilize the present invention according to description before.Preferred embodiment is only exemplarily illustrated thus before, and limits all the other contents herein never by any way.
Above, all temperature degree Celsius to list, all parts and per-cent all by weight, except as otherwise noted.
By aforementioned, those skilled in the art can easily determine essential characteristic of the present invention, and can not depart from its spirit and scope and make various change of the present invention and modification, make it to adapt to different purposes and condition.

Claims (9)

1. for implementing the device of following methods,
Described method comprises:
Supply raw materials stream, this feedstream comprises the impurity that ethene between 5 to 50 % by weight and at least one are selected from the group of following composition: the carbon monoxide of at least 0.1 % by weight, the hydrogen sulfide of at least 1 weight ppm, the ammonia of at least 1 weight ppm, the hydrogen of at least 5 % by weight and at least 0.1 % by weight carbonic acid gas; Under the pressure of the emergent pressure higher than ethene, make feedstream and catalyst exposure, the ratio that this catalyzer comprises silica-alumina is no more than the amorphous silica-alumina substrate of 30 and is selected from the metal of the group that group VIII and group vib metal form in periodictable; With
By in feedstream at least 40 % by weight ethylene conversion be heavier hydrocarbon;
Described device comprises:
Fluid catalytic cracking reactor, for making cracking catalyst contact with hydrocarbon feed stream, hydrocarbon feed to be cracked into the crackate with lower molecular weight, and makes sedimentation of coke on the cracking catalyst, to provide the cracking catalyst of coking;
Product exit, for discharging described crackate from described reactor;
Revivifier, for the cracking catalysis burn off coke by contacting with oxygen from described coking;
The product recovery section be communicated with described product exit, described product recovery section is used for described crackate to be separated into multiple product stream, comprises the stream containing ethene;
The compressor be communicated with described product recovery section, for compressing the described stream containing ethene to provide described feedstream; And
The fixed bed oligomerization reactor be communicated with described compressor, for becoming heavier hydrocarbon by the ethylene oligomerization in described feedstream.
2. the device of claim 1, wherein said catalyzer has the ratio of the silica-alumina being no more than 20.
3. the device of claim 1, wherein said catalyzer is the amorphous silica-alumina substrate with 0.5 ~ 15 % by weight nickel dipping.
4. the device of claim 1, wherein said catalyzer has the low temperature acidity ratio by ammonia temperature programmed control desorption thermometrically at least 0.15.
5. the device of claim 1, wherein said feedstream comprises the propylene being no more than 0.5 % by weight.
6. the device of claim 1, wherein said feedstream comprises the methane of 25 to 55 % by weight further.
7. the device of claim 1, it comprises further:
Cracking catalyst is made to contact with hydrocarbon feed stream with the crackate hydrocarbon becoming to have lower molecular weight by hydrocarbon cracking and make sedimentation of coke on the cracking catalyst to provide the cracking catalyst of coking;
The cracking catalyst of described coking is separated with described crackate;
Cracking catalyst to described coking adds oxygen;
The coke burnt in the cracking catalyst of described coking with oxygen is to make described regenerated cracking catalyst;
Be separated described crackate with obtain being formed described feedstream rare ethylene stream;
Compress this rare ethene and flow to 4,826 to 7, the pressure between 584kPa;
And
This rare ethylene stream is contacted with oligomerisation catalyst.
8. the device of claim 7, wherein said contact procedure is carried out in the fixed bed of described catalyzer.
9. the device any one of claim 1-8, comprises the king-tower susceptor be communicated with described product exit, the elementary adsorption tower be communicated with king-tower susceptor further, for the primary exhaust stream of providing package containing the described stream containing ethene.
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