WO2020177652A1 - 一种由劣质油生产低碳烯烃的方法和系统 - Google Patents

一种由劣质油生产低碳烯烃的方法和系统 Download PDF

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WO2020177652A1
WO2020177652A1 PCT/CN2020/077389 CN2020077389W WO2020177652A1 WO 2020177652 A1 WO2020177652 A1 WO 2020177652A1 CN 2020077389 W CN2020077389 W CN 2020077389W WO 2020177652 A1 WO2020177652 A1 WO 2020177652A1
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Prior art keywords
oil
weight
reaction
separation
hydro
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PCT/CN2020/077389
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English (en)
French (fr)
Inventor
侯焕娣
魏晓丽
龙军
董明
张久顺
侯栓弟
陈学峰
梁家林
李吉广
王翠红
申海平
龚剑洪
戴立顺
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中国石油化工股份有限公司
中国石油化工股份有限公司石油化工科学研究院
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Priority claimed from CN201910159559.1A external-priority patent/CN111647431B/zh
Priority claimed from CN201910159576.5A external-priority patent/CN111647432B/zh
Priority claimed from CN201910159674.9A external-priority patent/CN111647433B/zh
Application filed by 中国石油化工股份有限公司, 中国石油化工股份有限公司石油化工科学研究院 filed Critical 中国石油化工股份有限公司
Priority to JP2021552137A priority Critical patent/JP7479391B2/ja
Priority to SG11202109648P priority patent/SG11202109648PA/en
Priority to CA3131283A priority patent/CA3131283A1/en
Priority to EP20766163.8A priority patent/EP3936589A4/en
Priority to US17/435,569 priority patent/US12054682B2/en
Publication of WO2020177652A1 publication Critical patent/WO2020177652A1/zh

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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
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
    • C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
    • C10G69/04—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one step of catalytic cracking in the absence of hydrogen
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D11/00—Solvent extraction
    • B01D11/04—Solvent extraction of solutions which are liquid
    • B01D11/0492—Applications, solvents used
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D3/00—Distillation or related exchange processes in which liquids are contacted with gaseous media, e.g. stripping
    • B01D3/10—Vacuum distillation
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D3/00—Distillation or related exchange processes in which liquids are contacted with gaseous media, e.g. stripping
    • B01D3/14—Fractional distillation or use of a fractionation or rectification column
    • B01D3/143—Fractional distillation or use of a fractionation or rectification column by two or more of a fractionation, separation or rectification step
    • 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
    • B01J19/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J19/24—Stationary reactors without moving elements inside
    • B01J19/245—Stationary reactors without moving elements inside placed in series
    • 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
    • B01J19/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J19/24—Stationary reactors without moving elements inside
    • B01J19/2455—Stationary reactors without moving elements inside provoking a loop type movement of the reactants
    • B01J19/246—Stationary reactors without moving elements inside provoking a loop type movement of the reactants internally, i.e. the mixture circulating inside the vessel such that the upward stream is separated physically from the downward stream(s)
    • 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
    • B01J2219/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J2219/00002—Chemical plants
    • B01J2219/00027—Process aspects
    • B01J2219/0004—Processes in series
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1003—Waste materials
    • C10G2300/1007—Used oils
    • 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
    • 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
    • C10G2300/202—Heteroatoms content, i.e. S, N, O, P
    • 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
    • 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
    • C10G2300/205—Metal content
    • 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
    • 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
    • C10G2300/205—Metal content
    • C10G2300/206—Asphaltenes
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/30—Physical properties of feedstocks or products
    • C10G2300/301—Boiling range
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/30—Physical properties of feedstocks or products
    • C10G2300/308—Gravity, density, e.g. API
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4006—Temperature
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4012—Pressure
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4018—Spatial velocity, e.g. LHSV, WHSV
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/70—Catalyst aspects
    • 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

Definitions

  • This application relates to the catalytic conversion of hydrocarbon oils, and more specifically, to a method and system for producing low-carbon olefins by catalytic cracking of inferior oil through hydrogen catalytic upgrading.
  • Low-carbon olefins represented by ethylene and propylene are the most basic raw materials for the chemical industry. At present, about 98% of the ethylene in the world comes from tube furnace steam cracking technology. Of the raw materials for ethylene production, naphtha accounts for 46% and ethane accounts for 34%. Approximately 62% of propylene comes from the co-production of ethylene by steam cracking. Steam cracking technology has been perfected day by day, and it is a process that consumes a lot of energy, and is limited by the use of high temperature resistant materials, and the potential for further improvement is small.
  • Chinese Patent Application Publication CN101045884A discloses a method for producing clean diesel and low-carbon olefins from residual oil and heavy distillate oil.
  • the residual oil and the optional catalytic cracking oil slurry enter the solvent deasphalting unit, the obtained deasphalted oil and the optional heavy distillate oil enter the hydrogenation unit, and the hydrocracking reaction is carried out in the presence of hydrogen to separate the products.
  • WO2015084779A1 discloses a method for producing low-carbon olefins, especially propylene, using a combined solvent deasphalting and high-severity catalytic cracking process.
  • the method includes: solvent deasphalting treatment after mixing vacuum residue and solvent to obtain solvent-rich deasphalted oil and deoiled asphalt; the solvent-rich deasphalted oil enters the heavy oil deep catalytic cracking device after separating the solvent.
  • Cracking reaction obtains target products rich in low-carbon olefins, especially propylene.
  • the method of this patent application first conducts solvent deasphalting treatment on the residual oil, and then realizes the efficient conversion of deasphalted oil and generates low-carbon olefins through a combined process, but the deoiled asphalt is not used and processed.
  • Chinese Patent Announcement CN106701185B discloses a residual oil treatment method, including a solvent deasphalting device, a hydrotreating reaction zone, a hydrotreating reaction zone and a catalytic cracking reaction zone; the process method includes the following content: the residual oil raw material is obtained by fractionation Light fractions and heavy fractions, heavy fractions enter the solvent deasphalting unit for treatment to obtain deasphalted oil and deoiled asphalt.
  • the light fractions, deasphalted oil and hydrogen are mixed and passed through the series hydrotreating reaction zone and hydroprocessing reaction zone in sequence ,
  • the reaction effluent from the hydrotreating reaction zone undergoes gas-liquid separation, the gas phase is recycled back to the hydrotreating reaction zone and/or the hydrotreating reaction zone, the liquid phase directly enters the catalytic cracking reaction zone for catalytic cracking reaction, and the catalytic cracking reaction flows out
  • the product is separated to obtain dry gas, liquefied petroleum gas, catalytic cracking gasoline fraction, catalytic cracking diesel fraction, catalytic cracking heavy cycle oil and catalytic cracking oil slurry.
  • the method of this patent can prolong the stable operation period of the device.
  • Chinese Patent Publication CN1171978C discloses a high-sulfur and high-metal residue conversion method.
  • the product, of which heavy cycle oil can be recycled to the hydrotreating unit, and the oil slurry can be recycled to the solvent deasphalting unit.
  • the method reduces the investment and operating costs of the hydrotreating unit, and improves the yield and quality of light oil.
  • the existing technology uses a combination of solvent deasphalting and hydrotreating to provide high-quality raw materials for catalytic cracking.
  • the yield of deasphalted oil is low. Economically, the benefits are limited.
  • the deoiled bitumen has not been well used, resulting in the low utilization rate of inferior oil in the current technology, and more residues are still produced. Therefore, it is necessary to develop a green and high-efficiency conversion technology to produce low-carbon olefins from inferior oil, and to produce more high-value-added ethylene and propylene while improving the utilization rate of inferior oil.
  • the purpose of this application is to provide a method and system for producing low-carbon olefins from inferior oil.
  • the method and system provided in this application can not only realize the high-efficiency and green conversion of inferior oil, but also can realize the production of chemical raw material-low-carbon olefin from inferior oil .
  • this application provides a method for producing low-carbon olefins from poor-quality oil, which includes the following steps:
  • step 5 Perform a third separation on the hydro-upgraded oil obtained in step 4) to obtain hydro-upgraded heavy oil;
  • step 6) Catalytically crack the hydro-upgraded heavy oil obtained in step 5) to obtain a catalytic cracking product containing low-carbon olefins;
  • step 3) is returned to step 1) for the thermal conversion reaction.
  • this application also provides a system for producing low-carbon olefins from low-quality oil, including a thermal conversion reaction unit, a first separation unit, a second separation unit, a hydro-upgrading unit, a third separation unit, and catalytic cracking Unit, where:
  • the thermal conversion reaction unit is configured to make the low-quality oil raw material undergo a thermal conversion reaction in the presence of hydrogen to obtain a conversion product;
  • the first separation unit is configured to separate the conversion product therein to obtain a first separation product, wherein the content of the components having a boiling point below 350° C. in the first separation product is not more than about 5% by weight, and the boiling point is The content of the components between 350-524°C is about 20-60% by weight;
  • the second separation unit is configured to separate the first separation product therein to obtain upgraded oil and residue, and the second separation unit is selected from a vacuum distillation unit, a solvent extraction unit, or a combination thereof;
  • the hydro-upgrading unit is configured to cause the modified oil to undergo a hydro-upgrading reaction therein to obtain a hydro-upgraded oil;
  • the third separation unit is configured to separate the hydro-upgraded oil therein to obtain hydro-upgraded heavy oil
  • the catalytic cracking unit is configured to allow the hydro-upgraded heavy oil to undergo a catalytic cracking reaction therein to obtain a catalytic cracking product containing low-carbon olefins.
  • the method and system of the present application provide one or more of the following advantages:
  • the overall conversion rate of the inferior oil feedstock may be greater than 90% by weight, or even greater than 95% by weight, and the amount of residue discharged may be less than 10% by weight, or even less than 5% by weight.
  • the method and system provided in this application optimize the distillation range and composition of the material to be subjected to the second separation, so that the second separation process is easy to operate.
  • the present application can efficiently upgrade low-quality oil raw materials to provide a catalytic cracking unit with a modified oil rich in saturated structure and substantially free of heavy metals and asphaltenes.
  • the content of heavy metals (based on the total weight of nickel and vanadium) in the obtained modified oil can be less than 10 micrograms/g, or even less than 5 micrograms/g, and the content of asphaltenes in the modified oil can be less than 2.0 weight. %, or even less than 0.5% by weight.
  • This application can further process the modified oil to produce low-carbon olefins, a chemical raw material, and achieve a low-carbon olefin yield greater than 36%.
  • Figure 1a shows a schematic diagram of a preferred embodiment of the method and system of the present application
  • Figure 1b shows a schematic diagram of another preferred embodiment of the method and system of the present application.
  • Figure 2a shows a schematic diagram of another preferred embodiment of the method and system of the present application.
  • Figure 2b shows a schematic diagram of another preferred embodiment of the method and system of the present application.
  • boiling point, distillation range (sometimes also referred to as boiling range), final boiling point and initial boiling point or similar parameters all refer to values under normal pressure (101325 Pa).
  • any specific numerical value (including the end point of the numerical range) disclosed in this document is not limited to the precise value of the numerical value, but should be understood to also encompass values close to the precise value.
  • any combination can be made to obtain one or more new Numerical ranges, these new numerical ranges should also be considered as specifically disclosed herein.
  • this application provides a method for producing low-carbon olefins from inferior oil, which includes the following steps:
  • step 5 Perform a third separation on the hydro-upgraded oil obtained in step 4) to obtain hydro-upgraded heavy oil;
  • step 6) Perform a catalytic cracking reaction on the hydro-upgraded heavy oil obtained in step 5) to obtain a catalytic cracking product containing low-carbon olefins;
  • step 3) is returned to step 1) for the conversion reaction.
  • the method of the present application can maintain the long-term operation of the system under the condition of reducing residue rejection as much as possible and improving resource utilization.
  • the transformation reaction and each separation step are the key to determining whether it can operate for a long time.
  • the rate is very important for the stability control of the system and the stability of the separation operation.
  • the inventors have found through a lot of experiments that, in the conversion reaction, the conversion rate of components with a boiling point above 524° C.
  • the conversion reaction of step 1) is essentially a thermal conversion reaction, in which inferior oil macromolecules, especially asphaltene supramolecules, undergo supramolecular dissociation, cracking reactions of macromolecules, and S and N heteroatoms.
  • the thermal conversion reaction makes the conversion rate of the components with a boiling point above 524° C. in the inferior oil reach about 30-70% by weight, preferably about 30-60% by weight.
  • This application does not have strict restrictions on the conditions (including catalysts) and reactors of the conversion reaction, as long as the above conversion rate can be achieved.
  • the conversion reaction can be carried out with or without the presence of a conversion catalyst.
  • the conversion reaction is carried out in the presence of a conversion catalyst, and the conversion catalyst may include at least one selected from the group consisting of a group VB metal compound, a group VIB metal compound, and a group VIII metal compound, preferably At least one of Mo compound, W compound, Ni compound, Co compound, Fe compound, V compound, and Cr compound is included.
  • the conversion catalyst is an unsupported catalyst, such as a dispersed catalyst.
  • the conversion catalyst may be selected from a solid substance containing a sulfide of the aforementioned metal, an organic complex or a chelate containing the aforementioned metal, or an aqueous solution containing an oxide of the aforementioned metal.
  • the conversion catalyst may be, for example, molybdenum octoate, molybdenum naphthenate, nickel naphthenate, tungsten naphthenate, iron oleate, molybdenum dialkylthioformate, etc.
  • Such organic-metal complexes/chelates One or more of the above-mentioned metal oxides and/or sulfides, such as one or more of hematite, molybdenite, molybdenum sulfide, iron sulfide, etc.; or It is an aqueous solution containing the above-mentioned metal oxides and/or inorganic acid salts capable of decomposing the above-mentioned metal oxides, such as ammonium molybdate, molybdenum sulfate, molybdenum nitrate, nickel nitrate, cobalt nitrate, molybdenum oxide, iron oxide, nickel oxide, Aqueous solutions of tungsten oxide, vanadium oxide, etc.
  • the conversion catalyst is in a highly dispersed form in the reaction system, and its particle size is about 2 nm to about 50 ⁇ m, preferably about 2 nm to about 1 ⁇ m.
  • the conversion reaction of step 1) is carried out in a slurry bed reactor, wherein the liquid reaction raw materials are reacted under the action of a catalyst in the form of a solid suspension.
  • the conversion reaction can be carried out under the following conditions: the temperature is about 380-470°C, preferably about 400-440°C; the hydrogen partial pressure is about 10-25 MPa, preferably about 13-20 MPa; The volumetric space velocity of the inferior oil is about 0.01-2h -1 , preferably about 0.1-1.0h -1 ; the volume ratio of hydrogen to the inferior oil is about 500-5000, preferably about 800-2000, in the conversion catalyst The amount of the conversion catalyst used is about 10-50000 micrograms/g, preferably about 30-25000 micrograms/g, based on the weight of the inferior oil based on the active metal.
  • the inferior oil may be selected from low-quality feedstock oils containing asphaltenes, where asphaltenes refer to small non-polar n-alkanes (such as n-pentane or n-heptane) that are insoluble in the feedstock oil.
  • asphaltenes refer to small non-polar n-alkanes (such as n-pentane or n-heptane) that are insoluble in the feedstock oil.
  • a substance soluble in benzene or toluene A substance soluble in benzene or toluene.
  • the inferior oil meets one or more of the following indicators: API degree less than about 27, boiling point greater than about 350°C (preferably greater than about 500°C, more preferably greater than about 524°C), and asphaltene content greater than about 2% by weight (preferably greater than about 5% by weight, more preferably greater than about 10% by weight, and even more preferably greater than about 15% by weight), and the heavy metal content based on the total weight of nickel and vanadium is greater than about 100 micrograms/g.
  • the inferior oil may be selected from at least one of inferior crude oil, heavy oil, deoiled asphalt, coal-derived oil, shale oil, and petrochemical waste oil.
  • Other low-quality feedstock oils well-known to those skilled in the art can also be used alone or mixed as the inferior oil feedstock for conversion reaction, which will not be repeated in this application.
  • the "quality crude oil” may be, for example, heavy oil, wherein “heavy oil” refers to a higher content of asphaltenes and resins, high viscosity oil, typically ground 20 °C density greater than 0.943 g / cm 3 , Underground crude oil with a viscosity greater than 50 centipoise is called heavy oil.
  • the "heavy oil” refers to distillate oil or residual oil with a boiling point above 350°C
  • distillate oil generally refers to crude oil or secondary processed oil obtained by normal pressure distillation and vacuum distillation Distillate products, such as heavy diesel oil, heavy gas oil, lubricating oil fractions or cracking raw materials, etc.
  • residual oil refers to the bottom distillate of crude oil through atmospheric and vacuum distillation.
  • the bottom distillate of atmospheric distillation is called It is atmospheric residue (generally a fraction with a boiling point greater than 350°C), and the bottom distillate of a vacuum distillation column is generally called a vacuum residue (generally a fraction with a boiling point greater than 500°C or 524°C).
  • the residual oil may be at least one selected from the group consisting of topped crude oil, heavy oil obtained from oil sands pitch, and heavy oil with an initial boiling point greater than 350° C., where "topped crude oil” refers to the treatment of crude oil in an atmospheric and vacuum distillation process. During fractional distillation, the oil discharged from the bottom of the preliminary distillation column or the bottom of the flash distillation column.
  • the “deasphalted oil” refers to the asphaltene-rich and aromatic-rich extractant obtained by contacting with solvent, dissolving and separating, and extracting the bottom of the tower in the solvent deasphalting device. According to the different types of solvents, it can be divided into propane deoiled pitch, butane deoiled pitch, pentane deoiled pitch and so on.
  • the “coal-derived oil” refers to a liquid fuel obtained by chemical processing using coal as a raw material, and may be at least one selected from coal liquefied oil produced by coal liquefaction and coal tar produced by coal pyrolysis .
  • the “shale oil” refers to synthetic crude oil obtained by low-temperature dry distillation or other heat treatment of kerogen shale, or a brown viscous paste, which has a pungent odor and high nitrogen content.
  • the "said petrochemical waste oil” may be at least one selected from petrochemical waste oil sludge, petrochemical oil residue and refined products thereof.
  • step 2) the conversion product is first separated to obtain a first separated product, wherein the content of components with a boiling point below 350° C. in the first separated product is not more than about 5 wt%, It is preferably less than about 3% by weight, and the content of components having a boiling point of 350-524°C (preferably 355-500°C or 380-524°C, more preferably 400-500°C) is about 20-60% by weight, preferably It is about 25-55% by weight.
  • the initial boiling point of the first separated product is not lower than about 300°C, preferably not lower than about 330°C, more preferably not lower than about 350°C.
  • the first separation product is generally composed of components with higher boiling points in the conversion product, which includes the residue obtained in step 3) and upgraded oil.
  • the main component of the residue is asphaltene, which is also It includes some glue and aromatic components necessary to maintain fluidity; the modified oil can be used as a high-quality raw material for subsequent processing to obtain other oil products.
  • the remaining components with lower boiling points in the conversion product can be separated from the first separation product in step 2), such as gas products under standard conditions (such as dry gas and liquefied gas, etc.) and other components with boiling points below 350°C. Minute.
  • the first separation in step 2) is used to obtain the first separated product conforming to the above-mentioned distillation range composition, and the present application does not specifically limit its specific implementation.
  • the first separation is physical separation, such as extraction, distillation, evaporation, flashing, condensation, and the like.
  • the first separation in step 2) includes:
  • step 2a separating the conversion product obtained in step 1) at a first pressure and a first temperature to obtain a gas component and a liquid component;
  • step 2a gaseous products such as hydrogen are preferably separated, and the obtained gas component is rich in hydrogen, and the hydrogen content is preferably more than 85% by weight.
  • the first pressure in step 2a) may be about 10-25 MPa, preferably about 13-20 MPa.
  • the first pressure generally refers to the outlet pressure when the gas component leaves the separation device;
  • the first temperature may be about 380-470°C, preferably about 400-440°C.
  • the first temperature generally refers to the outlet temperature of the liquid component when it leaves the separation device.
  • the separation method in step 2a) can be selected from distillation, fractionation, flash evaporation, etc., preferably distillation.
  • the distillation can be carried out in a distillation tower, where the gas component can be obtained from the top of the distillation tower, and the liquid component can be obtained from the bottom of the distillation tower;
  • step 2b) it is preferable to separate the components with a boiling point below 350°C and to keep the components with a boiling point of 350-524°C as far as possible.
  • the second pressure in step 2b) is lower than the first pressure, preferably 4-24 MPa lower than the first pressure, more preferably 7-19 MPa lower; specifically, the second pressure may be about 0.1-5MPa, preferably 0.1-4MPa, for convenience of measurement, the second pressure generally refers to the outlet pressure of the second separated product when it leaves the separation device; the second temperature may be about 150-390°C, preferably 200-370 °C, for the convenience of measurement, the second temperature generally refers to the outlet temperature of the first separated product when it leaves the separation device.
  • the separation method of step 2b) can be distillation and/or fractional distillation, preferably atmospheric or pressurized fractional distillation, which can be carried out in an atmospheric distillation tank or a pressurized distillation column.
  • the second separated product obtained in step 2b) may contain light components separated under the second pressure and second temperature conditions that have a lower boiling point than the first separated product.
  • the first separation in step 2) may further include:
  • step 2d) returning at least a part of the gas components obtained in step 2a) to step 1) for the conversion reaction;
  • step 2e Return at least a part of the gas components obtained in step 2a) to step 4) for the hydro-upgrading.
  • the cutting of step 2c) can be carried out by fractional distillation or distillation, preferably by fractional distillation, for example, in a fractionation tower.
  • the operating pressure can be 0.05-2.0 MPa, preferably about 0.1-1.0 MPa, and the operating temperature can be It is 50-350°C, preferably 150-330°C.
  • step 2d) and step 2e) return at least a part of the gas components obtained in step 2a) to step 1) and/or step 4).
  • the gas components can be used directly or separated as recycled hydrogen.
  • the first separation in step 2) may further include:
  • step 2f) Return at least a part of the second separation product obtained in step 2b) and/or at least a part of the atmospheric gas oil obtained in step 2c) to step 4), and undergo hydro-upgrading together with the upgraded oil.
  • the second separation in step 3) is used to separate the easily processed upgraded oil from the residue in the first separated product, and the resulting residue is thrown outside or returned to step 1) in step 7) for conversion reaction.
  • the second separation in step 3) can be performed at a third temperature and a third pressure, using one or more of vacuum distillation and solvent extraction.
  • the vacuum distillation can be carried out in a distillation column with or without packing.
  • the third pressure is about 1-20 mmHg in vacuum
  • the third temperature is about 250-350°C.
  • the solvent extraction is preferably a countercurrent contact extraction between the extraction solvent and the first separated product, which can be carried out in any extraction device, for example, in an extraction tower.
  • the third pressure can be about 3-12 MPa, preferably about 3.5-10MPa
  • the third temperature may be about 55-300°C, preferably about 70-220°C
  • the extraction solvent may be C 3 -C 7 hydrocarbons, preferably C 3 -C 5 alkanes and C 3 -C 5 alkenes At least one, more preferably at least one of C 3 -C 4 alkanes and C 3 -C 4 alkenes
  • the weight ratio of the extraction solvent to the first isolated product is about 1:1 to about 7:1, Preferably it is about 1.5:1 to about 5:1.
  • Those skilled in the art can also adopt other conventional extraction methods for extraction, which will not be repeated in this application.
  • the residue obtained in step 3) is the component with the highest boiling point in the conversion product.
  • the softening point of the residue obtained in step 3) is preferably less than about 150°C, more preferably less than about 120°C.
  • the conversion catalyst therein when the conversion reaction is carried out in a slurry-bed reactor, the conversion catalyst therein will enter the subsequent separation step along with the conversion product and remain in the residue. With the increase in the amount of catalyst added and the inferior oil With the accumulation of metal components, the metal in the entire reaction system will continue to increase.
  • the residues need to be discharged intermittently or continuously, preferably part of the residues are thrown outside, and the ratio of the thrown residues to the total residues is preferably about 5-70% by weight, more preferably about 10-50% by weight; at the same time, in order to fully use the inferior oil, it is preferable to return part of the residue to step 1) in step 7), and the ratio of returned residue is preferably about 30-95% by weight, more preferably about 50-90% by weight .
  • Those skilled in the art can also adjust the ratio of residue rejection and recycling according to the different metal content of the inferior oil, which will not be repeated in this application.
  • the obtained modified oil is hydro-upgraded in step 4), and the obtained hydro-upgraded oil is cut and separated into hydro-upgraded in step 5)
  • Light oil and hydro-modified heavy oil the cut point between the hydro-modified light oil and the hydro-modified heavy oil may be about 340-360°C, preferably about 345-355°C, more preferably about 350°C; and
  • the obtained hydro-upgraded heavy oil is catalytically cracked to obtain a catalytic cracking product containing low-carbon olefins.
  • the catalytic cracking product can be separated to obtain dry gas, low-carbon olefin, gasoline, circulating oil and oil slurry.
  • the "cycle oil” generally includes light cycle oil and heavy cycle oil.
  • Light cycle oil can also be called diesel, which refers to the fraction with a boiling point between 205°C and 350°C obtained from the catalytic cracking reaction.
  • Heavy cycle oil refers to the boiling point The fraction between 343°C and 500°C;
  • the "slurry” generally refers to the bottom oil obtained from the fractionation step of the catalytic cracking reaction, and the product discharged from the bottom of the settler after separation by the settler, and from the sedimentation
  • the product discharged from the upper part of the device is generally called clarified oil.
  • the obtained slurry can be returned to step 1) for conversion reaction; after the obtained C3 hydrocarbons and C4 hydrocarbons are separated from alkanes-olefins, the C3 and C4 alkanes are sent to step 3) for use as extraction solvents; and/or,
  • the obtained circulating oil is hydro-upgraded separately or together with the modified oil.
  • the method of the present application can realize the return of oil slurry for conversion reaction, on the one hand, it can improve the utilization rate of raw materials, and convert the low value-added oil slurry into high value-added gasoline products rich in aromatic hydrocarbons; on the other hand, because the oil slurry is rich in Aromatic components can improve the stability of the conversion unit and extend the operation cycle of the device.
  • step 6 at least a part of the second separation product obtained in step 2b) and/or the atmospheric gas oil obtained in step 2c) can be catalytically cracked together with the hydro-upgraded heavy oil.
  • Step 6) and the above steps can realize the maximum production of chemical raw materials from inferior oil, and improve the utilization rate of the upgraded oil and the second separated product.
  • step 4 the hydro-upgrading of step 4) is well-known to those skilled in the art, and can be carried out in any manner known in the art, and there is no particular limitation. It can be any hydrogenation known in the art. It is carried out in a treatment device (such as a fixed bed reactor, a fluidized bed reactor), and those skilled in the art can make a reasonable choice.
  • a treatment device such as a fixed bed reactor, a fluidized bed reactor
  • the hydro-upgrading can be carried out under the following conditions: the partial pressure of hydrogen is about 5.0-20.0 MPa, preferably about 8-15 MPa; the reaction temperature is about 330-450°C, preferably about 350-420°C; The space velocity is about 0.1-3h -1 , preferably about 0.3-1.5h -1 ; the volume ratio of hydrogen to oil is about 300-3000, preferably about 800-1500; the catalyst used in the hydro-upgrading may include hydrogenation Refined catalyst and/or hydrocracking catalyst.
  • any catalyst conventionally used for this purpose in the art can be cited, or it can be manufactured according to any manufacturing method conventionally known in the art, and the hydrorefining catalyst
  • the amount of the hydrocracking catalyst used in the step can refer to the conventional knowledge in the field, and there is no particular limitation.
  • the hydrorefining catalyst may include a support and an active metal component
  • the active metal component is selected from the group VIB metals and/or the group VIII non-noble metals, especially the combination of nickel and tungsten, nickel , The combination of tungsten and cobalt, the combination of nickel and molybdenum, or the combination of cobalt and molybdenum.
  • the carrier include alumina, silica, and amorphous silica alumina. These carriers can be used singly or in combination of multiple types at any ratio.
  • the hydrorefining catalyst may include about 30-80% by weight of alumina support, about 5-40% by weight of molybdenum oxide, and about 5-15% by weight. Of cobalt oxide and about 5-15 wt% nickel oxide. Those skilled in the art can also use hydrorefining catalysts with other compositions.
  • the hydrocracking catalyst generally includes a carrier, an active metal component and a cracking active component.
  • the active metal component for example, sulfides of metals of group VIB of the periodic table, sulfides of base metals of group VIII of the periodic table, or noble metals of group VIII of the periodic table, etc. , Especially Mo sulfide, W sulfide, Ni sulfide, Co sulfide, Fe sulfide, Cr sulfide, Pt and Pd, etc.
  • the cracking active component include amorphous silica alumina and molecular sieves.
  • the hydrocracking catalyst may include about 3-60% by weight of zeolite, about 10-80% by weight of alumina, and about 1-15% by weight of oxidation. Nickel and about 5-40% by weight of tungsten oxide, wherein the zeolite is a Y-type zeolite.
  • zeolite aluminum oxide, silicon oxide, titanium oxide, activated carbon, and the like.
  • the hydrocracking catalyst may include about 3-60% by weight of zeolite, about 10-80% by weight of alumina, and about 1-15% by weight of oxidation. Nickel and about 5-40% by weight of tungsten oxide, wherein the zeolite is a Y-type zeolite.
  • Those skilled in the art can also use hydrocracking catalysts with other compositions.
  • the catalyst used in the hydro-upgrading includes both a hydrorefining catalyst and a hydrocracking catalyst, and the filling volume ratio of the hydrorefining catalyst to the hydrocracking catalyst is about 1: 1 to about 5:1, according to the flow direction of the reaction materials, the hydrorefining catalyst is packed upstream of the hydrocracking catalyst.
  • the catalytic cracking in step 6) can be carried out in various forms of catalytic cracking reactors, preferably in a variable diameter dilute phase conveying bed reactor and/or a combined catalytic cracking reactor.
  • the catalytic cracking of step 6) is carried out in a variable diameter dilute phase transport bed reactor, wherein the variable diameter dilute phase transport bed reactor includes first reaction zones with different diameters from bottom to top And the second reaction zone, the ratio of the diameter of the second reaction zone to the diameter of the first reaction zone is about 1.2:1 to about 2.0:1.
  • the reaction conditions in the first reaction zone may include: a reaction temperature of about 500-620°C, a reaction pressure of about 0.2-1.2 MPa, and a reaction time of about 0.1-5.0 seconds
  • the weight ratio of the catalyst to the cracked raw material is about 5-15, and the weight ratio of steam to the cracked raw material is about 0.05:1 to about 0.3:1
  • the reaction conditions in the second reaction zone may include: a reaction temperature of about 450-550°C, The reaction pressure is about 0.2-1.2MPa, and the reaction time is about 1.0-20.0 seconds.
  • the catalytic cracking of step 6) is carried out in a combined catalytic cracking reactor, wherein the combined reactor has a first reaction zone and a second reaction zone connected in series from bottom to top, and The first reaction zone is a riser reactor, and the second reaction zone is a fluidized bed reactor.
  • the fluidized bed reactor is located downstream of the riser reactor and is connected to the outlet of the riser reactor.
  • it may be A composition reactor obtained by connecting fluidized bed reactors in series with conventional catalytic cracking riser reactors known to those skilled in the art.
  • the riser reactor may be selected from an equal diameter riser reactor and/or an equal linear velocity riser reactor, and an equal diameter riser is preferably used.
  • the riser reactor includes a pre-lift section and at least one reaction zone from bottom to top.
  • the number of reaction zones can be 2-8, preferably For 2-3.
  • the reaction conditions in the first reaction zone may include: a reaction temperature of about 560-750°C, preferably about 580-730°C, more preferably about 600-700°C;
  • the reaction time is about 1-10 seconds, preferably about 2-5 seconds;
  • the agent-to-oil ratio is about 1:1 to about 50:1, preferably about 5:1 to about 30:1;
  • the conditions may include: the reaction temperature is about 550-730°C, preferably about 570-720°C; the weight space velocity is about 0.5-20 h -1 , preferably about 2-10 h -1 .
  • water vapor may be injected into the riser reactor, the water vapor is preferably injected in the form of atomized steam, and the weight ratio of the injected water vapor to the feed oil may be about 0.01: 1 to about 1:1, preferably about 0.05:1 to about 0.5:1.
  • the method of the present application may further include separating the spent catalyst in the catalytic cracking product from the reaction oil and gas to obtain the spent catalyst and the reaction oil and gas, and then separating the obtained reaction oil and gas into dry gas through a subsequent separation system.
  • separating the spent catalyst in the catalytic cracking product from the reaction oil and gas to obtain the spent catalyst and the reaction oil and gas, and then separating the obtained reaction oil and gas into dry gas through a subsequent separation system.
  • Liquefied gas, gasoline, diesel and other distillates and further separate dry gas and liquefied gas through gas separation equipment to obtain ethylene and propylene.
  • the method for separating ethylene, propylene, etc. from the reaction product can adopt conventional techniques in this field, which is not particularly limited in this application and will not be described in detail here.
  • the method of the present application may further include regenerating the spent catalyst; and preferably, at least a part of the catalyst used in the catalytic cracking reaction is a regenerated catalyst, for example, all may be regenerated catalysts. .
  • the method of the present application may further include stripping (generally steam stripping) on the regenerated catalyst obtained by regeneration to remove impurities such as gas.
  • an oxygen-containing gas is generally introduced from the bottom of the regenerator, and the oxygen-containing gas may be air, for example.
  • the spent catalyst is contacted with oxygen for coking regeneration.
  • the flue gas generated after the coking and regeneration of the catalyst undergoes gas-solid separation in the upper part of the regenerator, and the flue gas enters the subsequent energy recovery system.
  • the regeneration operating conditions of the spent catalyst may be: the regeneration temperature is about 550-750°C, preferably about 600-730°C, more preferably about 650-700°C; the apparent linear velocity of the gas is about 0.5 -3 meters/second, preferably about 0.8-2.5 meters/second, more preferably about 1-2 meters/second, the average residence time of the spent catalyst is about 0.6-3 minutes, preferably about 0.8-2.5 minutes, more preferably About 1-2 minutes.
  • the catalytic cracking catalyst suitable for step 6) may be various catalytic cracking catalysts conventionally used in the art.
  • the catalytic cracking catalyst may comprise: about 1-60% by weight of zeolite, about 5-99% by weight of inorganic oxides, and about 0-70% by weight of clay.
  • the zeolite in the catalytic cracking catalyst, is used as an active component.
  • the zeolite is selected from medium pore zeolite and/or large pore zeolite.
  • the medium pore zeolite accounts for about 50-100% by weight of the total weight of the zeolite, preferably about 70-100% by weight
  • the large pore zeolite accounts for about 0-50% by weight of the total weight of the zeolite, preferably about 0-30% by weight. weight%.
  • the medium pore zeolite and the large pore zeolite have the meaning commonly understood in the art, that is, the average pore diameter of the medium pore zeolite is 0.5-0.6 nm, and the average pore diameter of the large pore zeolite is 0.7-1.0 nm.
  • the large pore zeolite may be selected from one or a mixture of two or more of rare earth Y (REY), rare earth hydrogen Y (REHY), ultra stable Y obtained by different methods, and high silicon Y. .
  • the medium pore zeolite may be selected from zeolites having an MFI structure, such as ZSM series zeolites and/or ZRP zeolites.
  • MFI structure such as ZSM series zeolites and/or ZRP zeolites.
  • the above-mentioned mesoporous zeolite can be modified with non-metal elements such as phosphorus and/or transition metal elements such as iron, cobalt, and nickel.
  • non-metal elements such as phosphorus and/or transition metal elements such as iron, cobalt, and nickel.
  • ZSM series zeolite can be selected from ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM- 35, ZSM-38, ZSM-48 and other similar structure of zeolite or a mixture of more of them.
  • ZSM-5 please refer to US Patent 3,702,886, the contents of which are hereby quoted in full Incorporated into this article.
  • the inorganic oxide is used as a binder, preferably selected from silica (SiO 2 ) and/or aluminum oxide (Al 2 O 3 ).
  • the clay is used as a substrate (i.e., a support), preferably selected from kaolin and/or hallucinite.
  • the method of the present application includes the following steps:
  • step 2) Perform a first separation on the conversion product obtained in step 1) to obtain a first separated product, wherein the content of the components with a boiling point below 350° C. in the first separated product is not more than about 5% by weight, preferably less than
  • the content of components with a boiling point of 350-524°C (preferably 355-500°C or 380-524°C, and more preferably 400-500°C) is about 20-60% by weight, preferably about 25-55% by weight
  • the initial boiling point of the first separated product is not lower than about 300°C, preferably not lower than about 330°C, more preferably not lower than about 350°C;
  • step 2) Performing a second separation on the first separation product obtained in step 2) to obtain upgraded oil and residue, the second separation being selected from vacuum distillation, solvent extraction or a combination thereof;
  • step 5 Perform a third separation on the hydro-upgraded oil obtained in step 4) to obtain hydro-upgraded heavy oil;
  • step 6) The hydro-upgraded heavy oil obtained in step 5) is preheated and then enters the bottom of the variable-diameter dilute phase conveying bed reactor, contacts the regenerated catalyst for catalytic cracking reaction, and flows upwards into the cyclone separator for gas-solid separation and separation After further separation of the reaction oil and gas to obtain products including propylene and high-octane gasoline; the separated spent catalyst is stripped and then enters the catalyst regenerator for coking regeneration, and the regenerated catalyst is returned to the reactor for recycling; or
  • the hydro-upgraded heavy oil obtained in step 5 enters the first reaction zone of the combined catalytic cracking reactor, contacts with the regenerated catalyst to carry out the catalytic cracking reaction, and flows upward into the second reaction zone to continue the catalytic cracking reaction.
  • the reaction gas and spent catalyst at the outlet of the reactor enter the cyclone separator for gas-solid separation.
  • the separated reaction gas and gas are further separated to obtain products containing low-carbon olefins;
  • the separated spent catalyst is stripped into the catalyst regenerator Coke regeneration, regenerated catalyst is returned to the reactor for recycling, wherein the low-carbon olefins include ethylene, propylene, and butene; and
  • step 7) Return the residue obtained in step 3) to step 1) for the conversion reaction; alternatively, the residue obtained in step 3) is thrown outside; or, a part of the residue obtained in step 3) is returned to step 1) for the conversion reaction; The conversion reaction is described, and the remaining part of the obtained residue is thrown out.
  • the present application provides a system for producing low-carbon olefins from poor-quality oil, including a conversion reaction unit, a first separation unit, a second separation unit, a hydro-upgrading unit, a third separation unit, and a catalytic cracking unit ,among them:
  • the conversion reaction unit is configured to make the low-quality oil raw material undergo a thermal conversion reaction in the presence of hydrogen to obtain a conversion product;
  • the first separation unit is configured to separate the conversion product therein to obtain a first separation product, wherein the content of the components having a boiling point below 350° C. in the first separation product is not more than about 5% by weight, and the boiling point is The content of the components between 350-524°C is about 20-60% by weight;
  • the second separation unit is configured to separate the first separation product therein to obtain upgraded oil and residue, and the second separation unit is selected from a vacuum distillation unit, a solvent extraction unit, or a combination thereof;
  • the hydro-upgrading unit is configured to cause the modified oil to undergo a hydro-upgrading reaction therein to obtain a hydro-upgraded oil;
  • the third separation unit is configured to separate the hydro-upgraded oil therein to obtain hydro-upgraded heavy oil
  • the catalytic cracking unit is configured to allow the hydro-upgraded heavy oil to undergo a catalytic cracking reaction therein to obtain a catalytic cracking product containing low-carbon olefins.
  • inferior oil, hydrogen, and the conversion catalyst are reacted in the conversion reactor to obtain the conversion reaction product and sent to the first separation unit.
  • the conversion reactor is a slurry bed reactor.
  • the conversion reaction product in the first separation unit, is first separated into a gas product and a liquid product, and then the liquid product is further separated to obtain a heavy fraction with a distillation range greater than about 350°C.
  • the first separation product is sent to the second separation unit.
  • the first separation product in the second separation unit, is separated in a vacuum distillation column, or is subjected to countercurrent contact with an extraction solvent in an extraction column for extraction and separation to obtain a modified oil And the residue, or in a combination of vacuum distillation and extraction separation to obtain a modified oil and a residue, and the modified oil is sent to the hydro-upgrading unit.
  • the residue is returned to the conversion reaction unit for further conversion.
  • the modified oil in the hydro-upgrading unit, is reacted under the action of a hydroprocessing catalyst to obtain the hydro-upgraded oil and sent to the third separation unit.
  • the hydro-upgraded oil is cut and separated into hydro-upgraded light oil and hydro-upgraded heavy oil, and the hydro-upgraded heavy oil is sent to the catalyst Cracking unit.
  • the catalytic cracking unit includes a variable diameter dilute phase transport bed reactor and/or a combined catalytic cracking reactor, wherein the variable diameter dilute phase transport bed reactor includes different diameters from bottom to top.
  • the ratio of the diameter of the second reaction zone to the diameter of the first reaction zone is about 1.2:1 to about 2.0:1;
  • the combined catalytic cracking reactor includes the first reaction zone from bottom to top
  • the reaction zone and the second reaction zone, the first reaction zone is a riser reactor, and the second reaction zone is a fluidized bed reactor.
  • the catalytic cracking catalyst in the catalytic cracking unit, enters the pre-lift section of the first reaction zone of the variable-diameter dilute phase conveying bed reactor, and flows upward under the action of the pre-lifting medium,
  • the preheated hydro-upgraded heavy oil is injected into the first reaction zone together with atomized steam, and it contacts the regenerated catalyst for catalytic cracking reaction while flowing upwards, and enters the second reaction zone to continue the reaction to obtain a catalytic cracking containing low-carbon olefins product.
  • the catalytic cracking product is separated in a subsequent separation system to obtain fractions such as ethylene, propylene, gasoline with high octane number; the separated spent catalyst enters the regenerator to be burnt and regenerated, and the regenerated catalyst that restores activity is returned The variable diameter dilute phase conveying bed reactor is recycled.
  • the catalytic cracking catalyst in the catalytic cracking unit, enters the pre-lifting section of the first reaction zone of the combined catalytic cracking reactor, flows upward under the action of the pre-lifting medium, and after preheating
  • the hydro-upgraded oil is injected into the first reaction zone together with atomized steam, and it contacts the regenerated catalyst for catalytic cracking reaction while flowing upwards, and enters the second reaction zone to continue the reaction to obtain a catalytic cracking product containing low-carbon olefins.
  • the catalytic cracking product is separated in a subsequent separation system to obtain fractions such as ethylene, propylene, and pyrolysis gasoline; the separated spent catalyst enters the regenerator to be burnt and regenerated, and the regenerated catalyst that restores activity is returned to the combined catalytic cracking reactor Used in recycling.
  • fractions such as ethylene, propylene, and pyrolysis gasoline
  • inferior raw materials are transported to conversion via pipeline 1, conversion catalyst via pipeline 2, fresh hydrogen via pipeline 3, circulating hydrogen via pipeline 4, catalytic oil slurry via pipeline 57, and residue via pipeline 5.
  • the thermal conversion reaction is carried out in the reactor 6.
  • the conversion product is transported to the high-pressure separation unit 8 via line 7 for pressure distillation, separated into gaseous components and liquid components, and then the gaseous components are transported as circulating hydrogen to the conversion reactor 6 via line 9 and line 4, or as hydrogen
  • the source is transported to the hydro-upgrading unit 23 via the pipeline 9 and the pipeline 11.
  • the liquid component is transported to the low pressure separation unit 12 through the pipeline 10 to undergo a sudden pressure drop, and is separated into a second separation product and a first separation product.
  • the second separated product enters the hydro-upgrading unit 23 through the line 14, and the first separated product is sent to the second separation unit 17 through the line 15 for vacuum distillation to separate the upgraded oil and residue (see Figures 1a and 2a), or
  • the countercurrent contact with the extraction solvent from the line 16 or/and the extraction solvent from the line 55 is carried out for extraction and separation in the second separation unit 17 to obtain upgraded oil and residue (see Figures 1b and 2b).
  • a part of the residue is thrown outside through the pipeline 19 and the pipeline 20, and the remaining part is recycled to the conversion reactor 6 through the pipeline 19 and the pipeline 5 to continue the conversion reaction together with the low-quality oil feedstock.
  • all the residues may be thrown outside through the pipeline 19 and the pipeline 20 without being recycled.
  • the upgraded oil is mixed with the second separation product from the pipeline 14 and the catalytic diesel from the pipeline 21 through the pipeline 18 and enters the hydro-upgrading unit 23 through the pipeline 22 for hydro-upgrading.
  • the hydro-upgraded product is separated and lighter.
  • the components and the hydro-upgraded light oil are respectively led out via pipeline 24 and pipeline 25, or the hydro-upgraded light oil is mixed with the hydro-upgraded heavy oil discharged via pipeline 26 via pipeline 25 and sent to the catalytic cracking unit via pipeline 28 (Figure The variable-diameter dilute phase transport bed reactor shown in 1a and 1b or the combined catalytic cracking reactor shown in Figs. 2a and 2b) in the first reaction zone 29.
  • the pre-lifting medium also enters the first reaction zone 29 through the pipeline 50, and the regenerated catalyst from the pipeline 48 is adjusted by the regeneration slide valve 49 and then enters the first reaction zone 29. Under the lifting action of the pre-lifting medium, it accelerates upward along the riser.
  • the preheated hydro-modified oil is injected into the first reaction zone 29 through the line 28 together with the atomized steam from the line 27, and is mixed with the existing stream in the first reaction zone 29.
  • the feed oil is catalyzed on the hot catalyst
  • the cracking reaction is accelerated upwards and enters the second reaction zone 30 of the catalytic cracking unit to continue the reaction.
  • the generated reaction product oil and gas and the deactivated spent catalyst enter the cyclone separator 34 in the settler 33 to separate the spent catalyst and the reaction product oil and gas.
  • the reaction product oil and gas enter the gas collecting chamber 35, and the fine catalyst powder is returned to the settler.
  • the spent catalyst in the settler flows to the stripping section 32 and contacts the steam from the pipeline 31.
  • the reaction product oil gas stripped from the spent catalyst enters the gas collecting chamber 35 after passing through the cyclone separator. After the stripped spent catalyst is adjusted by the standby slide valve 38, it enters the regenerator 39.
  • the air from the pipeline 44 is distributed by the air distributor 43 and then enters the regenerator 39, where it is burned out of the dense bed at the bottom of the regenerator 39
  • the coke on the spent catalyst regenerates the deactivated spent catalyst, and the flue gas enters the subsequent energy recovery system through the upper gas flue gas pipe 41 of the cyclone separator 40.
  • the pre-lifting medium may be dry gas, water vapor or a mixture thereof.
  • the regenerated catalyst enters the degassing tank 46 through the pipeline 45 connected to the catalyst outlet of the regenerator 39, and contacts the stripping medium from the pipeline 47 at the bottom of the degassing tank 46 to remove the flue gas entrained by the regenerated catalyst.
  • the regenerated catalyst is circulated to the bottom of the first reaction zone 29 through the line 48, the catalyst circulation can be controlled by the regeneration slide valve 49, the gas returns to the regenerator 39 through the line 42, and the reaction product oil and gas in the gas collection chamber 35 passes through the large oil and gas line 36 Into the subsequent separation system 58, the separated H 2 and C1-C2 alkanes are led out from the line 53, and the obtained light olefins (including C2, C3, C4 alkenes) are sent out of the system via the line 54; C3 and C4 alkanes are sent out through the line 55
  • the system or the second separation unit 17 is used as an extraction solvent, and the obtained aromatic hydrocarbon-rich gasoline is drawn from the pipeline 56 as a product, and the obtained circulating oil is drawn from the pipeline 21 and combined with the upgraded oil from the pipeline 18 and the refined oil from the pipeline 14. After the second separated products are mixed, they are sent to the hydro-upgrading unit 23 for hydro-upgrading, and the obtained oil slurry is led
  • the C4 or light gasoline fraction obtained by the separation of the catalytic cracking product can be sent to the variable diameter dilute phase transport bed reactor as a catalytic cracking unit through line 52 together with steam through line 51
  • the second reaction zone 30 is re-refined to further crack and increase the production of low-carbon olefins.
  • this application provides the following technical solutions:
  • a method for producing low-carbon olefins from inferior oil comprising:
  • Inferior oil enters the conversion reaction unit for conversion reaction, and the resulting reaction product is separated to obtain a heavy fraction with a boiling point greater than about 350°C;
  • the preheated hydro-modified oil enters the first reaction zone of the catalytic cracking reactor, contacts the regenerated catalyst for catalytic cracking reaction, and flows upward into the second reaction zone at the same time, and continues the catalytic cracking reaction.
  • the reaction oil and gas and the spent catalyst enter the cyclone separator for gas-solid separation, and the separated reaction gas and gas extraction device is further separated to obtain products containing low-carbon olefins; the separated spent catalyst is stripped and enters the catalyst regenerator for coke burning Regeneration, the regenerated catalyst is returned to the reactor for recycling.
  • the inferior oil includes at least one selected from the group consisting of inferior crude oil, heavy oil, deoiled asphalt, coal-derived oil, shale oil and petrochemical waste oil.
  • the modified raw material meets the requirements of heavy metals selected from the group consisting of heavy metals with an API degree less than about 27, a distillation range greater than about 350°C, an asphaltene content greater than about 2% by weight, and the total weight of nickel and vanadium The content is greater than one or more of about 100 micrograms/gram.
  • the volume ratio of hydrogen to inferior oil is about 500-5000, based on the metal in the conversion catalyst and based on the weight of the inferior oil, the amount of the conversion catalyst is about 10-50000 micrograms/g.
  • the operating conditions of the extraction separation unit include: pressure of about 3-12 MPa, temperature of about 55-300°C, extraction solvent of C3-C7 hydrocarbon, solvent and heavy fraction
  • the weight ratio is (1-7):1, or
  • the operating conditions of the vacuum distillation separation unit include: a vacuum of about 1-20 mmHg and a temperature of about 250-350°C.
  • the reaction conditions of the hydro-upgrading unit include: a hydrogen partial pressure of about 5.0-20.0 MPa, a reaction temperature of about 330-450°C, and a volumetric space velocity of about 0.1-3 Hour -1 , the volume ratio of hydrogen to oil is about 300-3000.
  • the catalyst used in the hydro-upgrading unit includes a hydrorefining catalyst and a hydrocracking catalyst
  • the hydrorefining catalyst includes a carrier and an active metal component, and the active metal component It is selected from group VIB metals and/or group VIII non-noble metals
  • the hydrocracking catalyst includes zeolite, alumina, at least one group VIII metal component and at least one group VIB metal component.
  • the hydrocracking catalyst is based on a catalyst, and its composition is: 3-60% by weight of zeolite, 10-80% by weight of alumina, 1-15% by weight of nickel oxide, and tungsten oxide 5-40% by weight.
  • the reactor of the catalytic cracking unit includes a first reaction zone and a second reaction zone, the first reaction zone is a riser reactor, and the second reaction zone is a fluidized bed reaction Device.
  • the conditions of the first reaction zone include: a reaction temperature of 560 to 750°C, a time of 1 to 10 seconds, and a catalyst-to-oil ratio of 1 to 50:1; the second reaction zone
  • the conditions include: the reaction temperature is 550-700°C, and the space velocity is about 0.5-20 h -1 .
  • the catalyst in step (4) contains: 1 to 60% by weight of zeolite, 5 to 99% by weight of inorganic oxides, and 0 to 70% by weight of clay, all based on the catalyst Based on the total weight of the zeolite, the zeolite is selected from medium pore zeolite and optional large pore zeolite, the medium pore zeolite accounts for 50-100% by weight of the total weight of the zeolite, and the large pore zeolite accounts for 0-50% by weight of the total zeolite weight.
  • the method according to item A13 characterized in that the medium pore zeolite accounts for 70-100% by weight of the total weight of the zeolite, and the large pore zeolite accounts for 0-30% by weight of the total zeolite weight.
  • step (2) The method according to item A1, characterized in that the residue described in step (2) is returned to step (1) for the conversion reaction; or, the residue obtained in step (2) is thrown outside; or Part of the residue obtained in step (2) is returned to step (1) for the conversion reaction, and the remaining part of the residue is subjected to external rejection.
  • a system for producing low-carbon olefins from inferior oil includes a conversion reaction unit, an extraction or vacuum distillation separation unit, a hydro-upgrading unit and a catalytic cracking unit, wherein the conversion reaction unit is combined with vacuum distillation or/and extraction
  • the separation unit is connected, the vacuum distillation or/and extraction separation unit is connected with the hydro-upgrading unit, and the hydro-upgrading unit is connected with the catalytic cracking unit.
  • a upgrading method for producing low-carbon olefins from low-quality oil comprising:
  • step (2) The hydrogen conversion product obtained in step (1) is subjected to separation treatment to obtain at least the first separated product; wherein, in the first separated product, the content of the components with a boiling point below 350° C. is not more than about 5 wt% , The content of components with boiling point between 350-524°C is about 20-60% by weight;
  • step (3) The first separated product obtained in step (2) is separated by vacuum distillation in a vacuum distillation separation unit or/and is extracted and separated by an extraction solvent in the extraction separation unit to obtain upgraded oil and residue;
  • step (3) Return the residue obtained in step (3) to step (1) to carry out the hydroconversion reaction; alternatively, the residue obtained in step (3) is thrown out; or, part of the residue in step (3) The obtained residue is returned to step (1) to perform the hydroconversion reaction, and the remaining part of the residue is subjected to external rejection;
  • step (6) Separate the hydro-upgraded oil obtained in step (5), and subject the obtained hydro-upgraded heavy oil to a catalytic conversion reaction to obtain a product containing low-carbon olefins.
  • step (1) the conversion rate of the hydrogen conversion reaction is about 30-60% by weight.
  • step (1) the hydrogen conversion reaction is carried out in a slurry-bed reactor.
  • step (1) the hydroconversion reaction is carried out in the presence or absence of a hydroconversion catalyst, and the hydroconversion catalyst contains selected from the group VB At least one of a metal compound, a group VIB metal compound, and a group VIII metal compound.
  • the conditions of the hydrogen conversion reaction include: a temperature of about 380-470°C, a hydrogen partial pressure of 10-25 MPa, The volumetric space velocity is about 0.01-2 hours -1 , and the volume ratio of hydrogen to the reformed raw material is about 500-5000, based on the metal in the hydrogen conversion catalyst and based on the weight of the reformed raw material.
  • the dosage is about 10-50000 micrograms/g.
  • the modified raw material includes at least one selected from the group consisting of inferior crude oil, heavy oil, deoiled pitch, coal-derived oil, shale oil and petrochemical waste oil kind.
  • step (2) in the first separated product, the content of components with a boiling point below 350°C is less than about 3% by weight, and the boiling point is between 350-524°C.
  • the content of ingredients is about 25-55% by weight.
  • step (2) the separation process includes:
  • step (1) Separating the hydrogen conversion product obtained in step (1) at a first pressure and a first temperature to obtain a gas component and a liquid component;
  • step (2-4) Return the gas components obtained in step (2-1) to step (1) for hydroconversion reaction and/or step (5) for hydrogenation upgrading.
  • Metal components, the active metal components are selected from group VIB metals and/or group VIII non-noble metals; the hydrocracking catalyst includes zeolite, alumina, at least one group VIII metal component and at least one Group VIB metal components.
  • the hydrocracking catalyst includes about 3-60% by weight of zeolite and about 10-80% by weight of oxidation Aluminum, about 1-15% by weight of nickel oxide, and about 5-40% by weight of tungsten oxide.
  • step (3) the extraction and separation are performed in an extraction solvent at a third temperature and a third pressure; wherein the third pressure is about 3-12 trillion Pa, the third temperature is about 55-300 deg.] C, the extraction solvent is a C 3 -C 7 hydrocarbons, the extraction solvent weight ratio of the first separation of the product (1-7): 1.
  • step (3) the softening point of the residue is less than about 150°C.
  • step (4) the ratio of the residue returned to step (1) to the total residue is 30-95% by weight, preferably 50-90% by weight.
  • step (6) the hydro-upgraded oil is cut and separated into hydro-upgraded light oil and hydro-upgraded heavy oil.
  • the cutting point between the heavy oil is 340°C to 360°C, preferably about 345-355°C, more preferably about 350°C.
  • a upgrading system for producing low-carbon olefins from low-quality oil includes a hydroconversion reaction unit, a vacuum distillation or/and extraction separation unit, a hydro-upgrading unit, and a catalytic conversion unit.
  • the reaction unit is connected with the vacuum distillation or/and extraction separation unit
  • the vacuum distillation or/and extraction separation unit is connected with the hydro-upgrading unit
  • the hydro-upgrading unit is connected with the catalytic conversion unit.
  • a method for producing propylene and high-octane gasoline from low-quality oil comprising:
  • Inferior oil enters the conversion reaction unit for conversion reaction, and the resulting reaction product is separated to obtain a heavy fraction with a distillation range greater than about 350°C;
  • the preheated hydromodified oil enters the bottom of the variable-diameter dilute phase conveying bed reactor, contacts the regenerated catalyst for catalytic cracking reaction, and flows upward into the cyclone separator for gas-solid separation, and the separated reaction oil and gas extraction device ,
  • the product containing propylene and high-octane gasoline is further separated; the separated spent catalyst is stripped and then enters the catalyst regenerator for coking regeneration, and the regenerated catalyst is returned to the reactor for recycling.
  • inferior oil includes at least one selected from the group consisting of inferior crude oil, heavy oil, deoiled asphalt, coal-derived oil, shale oil and petrochemical waste oil.
  • the modified raw material meets the requirements selected from the group consisting of an API degree less than about 27, a distillation range greater than about 350°C, an asphaltene content greater than about 2% by weight, and a total weight of nickel and vanadium.
  • the content of heavy metals is greater than one or more of about 100 micrograms/g.
  • the volume ratio of hydrogen to inferior oil is about 500-5000, based on the metal in the reforming catalyst and based on the weight of the upgraded raw material, the amount of the reforming catalyst is about 10-50000 micrograms/g.
  • the reaction conditions of the extraction separation unit include: a pressure of about 3-12 MPa, a temperature of about 55-300°C, the extraction solvent is a C3-C7 hydrocarbon, a solvent and a heavy fraction
  • the weight ratio is (1-7):1, or
  • the operating conditions of the vacuum distillation separation unit include: a vacuum of about 1-20 mmHg and a temperature of about 250-350°C.
  • the reaction conditions of the hydro-upgrading unit include: a hydrogen partial pressure of about 5.0-20.0 MPa, a reaction temperature of about 330-450°C, and a volumetric space velocity of about 0.1-3 Hour -1 , the volume ratio of hydrogen to oil is about 300-3000.
  • the catalyst used in the hydro-upgrading unit includes a hydrorefining catalyst and a hydrocracking catalyst, the hydrorefining catalyst includes a carrier and an active metal component, and the active metal
  • the components are selected from Group VIB metals and/or Group VIII non-noble metals;
  • the hydrocracking catalyst includes zeolite, alumina, at least one Group VIII metal component and at least one Group VIB metal component.
  • the hydrocracking catalyst is based on a catalyst, and its composition is: 3-60% by weight of zeolite, 10-80% by weight of alumina, 1-15% by weight of nickel oxide, and tungsten oxide 5-40% by weight.
  • variable-diameter dilute phase transport bed includes two reaction zones, and the ratio of the diameter of the second reaction zone to the diameter of the first reaction zone is 1.2-2.0:1.
  • reaction conditions of the first reaction zone in the reduced diameter dilute phase transport bed include: reaction temperature 500-620°C, reaction pressure 0.2-1.2 MPa, reaction time 0.1-5.0 Second, the weight ratio of catalyst to raw material is 5-15, and the weight ratio of steam to raw oil is 0.05-0.3:1.
  • reaction conditions of the second reaction zone in the reduced diameter dilute phase transport bed include: reaction temperature 450-550°C, reaction pressure 0.2-1.2 MPa, reaction time 1.0-20.0 second.
  • the catalyst contains: 1 to 60% by weight of zeolite, 5 to 99% by weight of inorganic oxides, and 0 to 70% by weight of clay, wherein zeolite It is selected from medium pore zeolite and optional large pore zeolite.
  • the medium pore zeolite accounts for 50-100% by weight of the total weight of the zeolite
  • the large pore zeolite accounts for 0-50% by weight of the total weight of the zeolite.
  • step (2) is returned to step (1) for the conversion reaction; or, the residue obtained in step (2) is thrown outside; or Part of the residue obtained in step (2) is returned to step (1) for the conversion reaction, and the remaining part of the residue is subjected to external rejection.
  • a system for producing propylene and high-octane gasoline from low-quality oil includes a conversion reaction unit, a vacuum distillation or/and extraction separation unit, a hydro-upgrading unit, and a catalytic cracking unit.
  • the pressure distillation or/and extraction separation unit is connected, the vacuum distillation or/and extraction separation unit is connected with the hydro-upgrading unit, and the hydro-upgrading unit is connected with the catalytic cracking unit.
  • the determination method for the content of heavy metals (based on the total weight of nickel and vanadium) in oil products is carried out in accordance with ASTM D5708;
  • Residue rejection rate weight of rejection residue/weight of inferior oil ⁇ 100% by weight
  • Inferior oil conversion rate (1-residue rejection rate) ⁇ 100% by weight
  • LPG+liquid product yield (the total weight of LPG and liquid product obtained by the first separation and the second separation of the conversion product)/inferior oil weight ⁇ 100 wt%;
  • Yield of toluene insoluble matter (weight of toluene insoluble matter in the first separated product/weight of the first separated product) ⁇ yield of the first separated product ⁇ 100%;
  • the softening point of the residue is determined by the GB/T 4507-84 method
  • the special component refers to the component with the boiling point between 350-524°C in the first separated product
  • Ethylene yield ethylene weight/catalytic cracking feedstock oil weight ⁇ 100% by weight
  • Propylene yield weight of propylene/weight of catalytic cracking feed oil ⁇ 100% by weight
  • Butene yield butene weight/catalytic cracking feed oil weight ⁇ 100% by weight
  • the inferior oil B used is vacuum residue, the properties of which are shown in Table 1.
  • the conversion reaction is carried out in a slurry-bed reactor with the inferior oil B as the raw material, and then the first separation is carried out, and the first separation is carried out in two fractionation towers to obtain the first separation product and the second separation product.
  • the first separation product is then subjected to a second separation (the extraction separation shown in Figures 1b and 2b in Example 1, and the vacuum distillation shown in Figures 1a and 2a in Example 3) to obtain upgraded oil and residue.
  • the specific conditions and results of each step are listed in Table 2-1 and Table 2-2.
  • the conversion reaction is carried out in a slurry-bed reactor with the inferior oil B as the raw material, and then the first separation is carried out, and the first separation is carried out in two fractionation towers to obtain the first separation product and the second separation product.
  • the first separation product is then subjected to a second separation (the extraction separation shown in Figures 1b and 2b in Example 2 and the vacuum distillation shown in Figures 1a and 2a in Example 4) to obtain upgraded oil and residue.
  • the basic process is the same as that of Example 1, except that the conversion reaction and the first separation are not performed.
  • the specific conditions and results of each step are listed in Table 2-1 and Table 2-2.
  • Example 2 The basic process is the same as in Example 2, except that the conversion reaction and the first separation are not carried out.
  • the specific conditions and results of each step are listed in Table 2-1 and Table 2-2.
  • Example 2 The basic process is the same as in Example 2, except that different conversion catalysts and operating conditions are used. The specific conditions and results of each step are listed in Table 2-1 and Table 2-2.
  • Molybdenum octoate self-made in the laboratory, purity greater than 90%;
  • Molybdenum naphthenate self-made in the laboratory, purity greater than 85%;
  • Nickel naphthenate self-made in the laboratory, purity greater than 90%
  • Comparative Example 3 show that when the content of components with a boiling point of less than 350°C in the first separated product exceeds the scope defined in this application, the conversion rate of inferior oil is reduced by 12%, and the yield of LPG+liquid product is reduced by 11%. At the same time, the heavy metal content of the modified oil reaches 20 ⁇ g/g, and the yield of toluene insoluble matter increases by about 1%.
  • Example 2 and Example 4 were sent to the hydro-upgrading unit, and the hydrorefining and cracking temperatures were 380-386°C, the volumetric space velocity was 0.5h -1 , the hydrogen-to-oil volume ratio was 1000 and Hydro-upgrading was carried out under a hydrogen partial pressure of 15 MPa to obtain hydro-upgraded oil.
  • the hydro-upgraded oil was simply separated to obtain hydro-upgraded heavy oil.
  • the test conditions and properties of the hydro-upgraded heavy oil are shown in Table 3.
  • the hydro-upgraded heavy oil obtained in Example 5 and Example 6 was subjected to catalytic cracking (the reactor type is shown in Figure 1a and 1b).
  • the catalytic cracking reaction was carried out on a medium-sized device.
  • the catalyst was a commercial product produced by Qilu Catalyst Branch.
  • the catalyst is CGP.
  • the preheated hydro-modified oil enters the first reaction zone of the variable diameter dilute phase conveying bed reactor at a reaction temperature of 535°C, a reaction time of 1.8 seconds, the weight ratio of catalyst to feed oil 8, and the weight ratio of water vapor to feed oil The reaction was performed under the condition of 0.10.
  • the oil-gas mixture and the catalyst continue to move upward into the second reaction zone, and continue to react at a reaction temperature of 510°C and a reaction time of 2.5 seconds.
  • the reaction oil and gas and the spent catalyst enter the closed cyclone separator from the reactor outlet, and the reaction gas and the spent catalyst are quickly separated.
  • the reaction gas and gas are cut according to the distillation range in the separation system to obtain fractions such as propylene and gasoline; the spent catalyst is in gravity Under the action, it enters the stripping section, and steam strips out the hydrocarbon products adsorbed on the spent catalyst.
  • the stripped catalyst enters the regenerator and contacts the air for regeneration; the regenerated catalyst enters the degassing tank to remove the regeneration The non-hydrocarbon gas impurities adsorbed and carried by the catalyst; the regenerated catalyst after degassing is returned to the variable diameter dilute phase conveying bed reactor for recycling.
  • the operating conditions and product distribution of the catalytic cracking unit are listed in Table 4.
  • the process is basically the same as in Examples 7-8, except that the raw materials are the hydro-upgraded heavy oils obtained in Comparative Examples 4-5.
  • the operating conditions and product distribution of the catalytic cracking unit are listed in Table 4.
  • Catalytic cracking was performed on the hydro-upgraded heavy oil obtained in Example 5 and Example 6 (the reactor type is shown in Figure 2a and 2b).
  • the catalytic cracking reaction was carried out on a medium-sized device.
  • the catalyst was a commercial product produced by Qilu Catalyst Branch. The brand name is MMC-2 catalytic cracking catalyst.
  • the preheated hydromodified oil enters the first reaction zone of the combined catalytic cracking reactor.
  • the temperature at the outlet of the riser is 580°C
  • the reaction time is 1.8 seconds
  • the weight ratio of catalytic cracking catalyst to feed oil is 15, and the ratio of water vapor to feed oil
  • the cleavage reaction was carried out at a weight ratio of 0.25.
  • the oil-gas mixture and the catalyst continue to move upward into the second reaction zone, and continue to react at a reaction temperature of 565°C and a bed weight space velocity of 4h -1 .
  • the reaction oil and gas and the spent catalyst enter the closed cyclone separator from the reactor outlet, and the reaction gas and the spent catalyst are quickly separated.
  • the reaction oil and gas are cut according to the distillation range in the separation system to obtain fractions such as ethylene, propylene and pyrolysis gasoline;
  • the catalyst enters the stripping section under the action of gravity, and the steam is stripped out of the hydrocarbon products adsorbed on the spent catalyst.
  • the stripped catalyst enters the regenerator and contacts with air for regeneration; the regenerated catalyst enters the degassing tank. To remove the non-hydrocarbon gas impurities adsorbed and carried by the regenerated catalyst; the regenerated catalyst after degassing is returned to the riser reaction for recycling.
  • Table 5 The operating conditions and product distribution of the catalytic cracking unit are listed in Table 5.
  • the process is basically the same as in Examples 9-10, except that the raw material is the hydro-upgraded heavy oil obtained in Comparative Examples 4-5.
  • the operating conditions and product distribution of the catalytic cracking unit are listed in Table 5.
  • the light cycle oil fraction with a distillation range of less than 350°C in the cycle oil obtained in Example 10 and the upgraded oil obtained in Example 4 were subjected to hydro-upgrading according to the process of Example 6.
  • the conditions and product properties of the hydro-upgrading were listed In Table 6.
  • Example 11 The hydro-upgraded heavy oil obtained in Example 11 was subjected to catalytic cracking in a conventional riser reactor.
  • the catalytic cracking catalysts were obtained from Qilu Branch of Sinopec Catalyst Co., Ltd. The conditions and results of catalytic cracking are listed in Table 7.
  • Example 2 On a medium-sized plant, the operation was performed referring to Example 2, in which the oil slurry obtained in Example 10 was recycled back to the conversion reaction, mixed with the inferior oil B and the circulating residue, and then subjected to the conversion reaction together, and then the conversion product was first separated to obtain The first separation product and the second separation product. The first separation product is then subjected to a second separation (extraction separation) to obtain upgraded oil and residue. Part of the residue circulates, and the rest is thrown away.
  • the specific operating conditions of each step are the same as in Example 2, and the reaction results obtained are shown in Table 8.
  • the method and system of the present application can greatly increase the yield of LPG+ liquid products obtained by upgrading inferior oil, and at the same time improve the quality of the raw materials of the catalytic cracking unit, with obvious yields of ethylene and propylene. It also has the advantage of high yield of high-octane gasoline.

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Abstract

一种由劣质油生产低碳烯烃的方法和系统，所述方法包括如下步骤：在氢气存在下对劣质油进行热转化反应得到转化产物；对所得转化产物进行分离得到第一分离产物；对所述第一分离产物进行分离得到改质油和残渣；对所得改质油进行加氢改质得到加氢改质油；对加氢改质油进行分离得到加氢改质重油；以及，对加氢改质重油进行催化裂解得到包含低碳烯烃的催化裂解产物。所述方法和系统不仅可实现劣质油的高效绿色转化，还可实现劣质油的高价值利用——生产低碳烯烃。

Description

一种由劣质油生产低碳烯烃的方法和系统
相关申请的交叉引用
本申请要求申请人于2019年3月4日提交的申请号为201910159559.1、名称为“一种劣质油生产低碳烯烃的方法和系统”的专利申请，于2019年3月4日提交的申请号为201910159576.5、名称为“一种由低品质油生产低碳烯烃的改质方法和系统”的专利申请，以及于2019年3月4日提交的申请号为201910159674.9、名称为“一种劣质油生产丙烯和高辛烷值汽油的方法和系统”的专利申请的优先权，其内容经此引用全文并入本文。
技术领域
本申请涉及烃油的催化转化，更具体地说，涉及一种劣质油通过临氢催化改质后进行催化裂解生产低碳烯烃的方法和系统。
背景技术
以乙烯、丙烯为代表的低碳烯烃是化学工业的最基本原料。目前，世界上约98％的乙烯来自于管式炉蒸汽裂解技术，在乙烯生产原料中，石脑油占46％，乙烷占34％。大约62％丙烯来自蒸汽裂解制乙烯的联产。蒸汽裂解技术已日臻完善，并且是大量消耗能源的过程，又受使用耐高温材质的局限，进一步改进的潜力已很小。
随着世界经济的缓慢复苏，石油需求增长放缓，世界石油市场供需基本面保持宽松。国际能源机构认为，在供给侧，未来几年以美国为代表的非欧佩克国家原油产量将持续上升，在2022年全球原油需求将会趋紧；在需求侧，未来5年全球原油需求持续攀升，2019年或将突破1亿桶/日；其中，非常规油、劣质重油的加工量将逐年增加。因此，利用非常规油或劣质油来最大量生产低碳烯烃之类的化工原料，是石化企业拓宽低碳烯烃生产原料来源，产品结构调整、提质增效的关键和重点。
中国专利申请公开CN101045884A公开了一种由渣油和重馏分油生产清洁柴油和低碳烯烃的方法。该方法是渣油与任选的催化裂解油浆进入溶剂脱沥青单元，所得的脱沥青油与任选的重馏分油进入加氢 单元，在氢气存在的条件下进行加氢裂化反应，分离产物得到轻重石脑油馏分、柴油馏分和加氢尾油；加氢尾油进入催化裂解单元，进行催化裂解反应，分离产物得到低碳烯烃、汽油馏分、柴油馏分和油浆；柴油循环回催化裂解单元，全部或部分的油浆返回溶剂脱沥青单元。该方法加工减压渣油和催化裂解油浆的混合物，可获得27.3重量％丙烯和10.6重量％乙烯。
国际申请公开WO2015084779A1公开了一种采用溶剂脱沥青和高苛刻度的催化裂化组合工艺来生产低碳烯烃，尤其是丙烯的方法。该方法包括：减压渣油和溶剂混合后进行溶剂脱沥青处理，得到富含溶剂的脱沥青油和脱油沥青；富含溶剂的脱沥青油经分离溶剂之后进入重油深度催化裂解装置进行深度裂解反应，得到富含低碳烯烃尤其是丙烯的目标产物。该专利申请的方法首先将渣油进行溶剂脱沥青处理，然后通过组合工艺实现了脱沥青油高效转化且生成低碳烯烃，但对于脱油沥青未进行使用和加工。
中国专利公告CN106701185B公开了一种渣油处理方法，包括溶剂脱沥青装置、加氢预处理反应区、加氢处理反应区和催化裂化反应区；所述工艺方法包括以下内容：渣油原料分馏得到轻馏分和重馏分，重馏分进入溶剂脱沥青装置处理后得到脱沥青油和脱油沥青，轻馏分和脱沥青油及氢气混合后依次经过串联的加氢预处理反应区和加氢处理反应区，加氢处理反应区的反应流出物进行气液分离，气相循环回加氢预处理反应区和/或加氢处理反应区，液相直接进入催化裂化反应区进行催化裂化反应，催化裂化反应流出物分离得到干气、液化气、催化裂化汽油馏分、催化裂化柴油馏分、催化裂化重循环油和催化裂化油浆。该专利的方法可以延长装置的稳定运作周期。
中国专利公告CN1171978C公开了一种高硫高金属渣油转化方法，渣油、油浆和溶剂经抽提得到的脱沥青油与重循环油、任选的溶剂精制抽出油一起进入加氢处理装置，在氢气和加氢催化剂存在下反应，分离产物得到气体、石脑油、加氢柴油和加氢尾油，其中加氢尾油进入催化裂化装置，在裂化催化剂存在下进行裂化反应，分离反应产物，其中重循环油可循环至加氢处理装置，油浆循环至溶剂脱沥青装置。该方法降低加氢处理装置投资和操作费用，提高轻质油的收率和质量。
为了从劣质油中更多地获取低碳烯烃，现有技术采用了溶剂脱沥 青与加氢处理组合的技术方法为催化裂解提供优质原料，但脱沥青油的收率较低，从全流程的经济性看，收益有限，另外脱油沥青也并未得到很好的利用，导致现在技术中劣质油利用率不高，仍产生较多的残渣。因此，有必要开发一种劣质油生产低碳烯烃的绿色高效转化技术，在提高劣质油利用率的同时，更多地生产高附加值的乙烯和丙烯等。
发明内容
本申请的目的是提供一种由劣质油生产低碳烯烃的方法和系统，本申请提供的方法和系统不仅可实现劣质油的高效绿色转化，而且可实现由劣质油生产化工原料-低碳烯烃。
为了实现上述目的，一方面，本申请提供了一种由劣质油生产低碳烯烃的方法，包括如下步骤：
1)在氢气存在下对劣质油原料进行热转化反应，得到转化产物；
2)对所述转化产物进行第一分离，得到第一分离产物，其中所述第一分离产物中沸点在350℃以下的组分的含量为不大于约5重量％，沸点在350-524℃之间的组分的含量为约20-60重量％；
3)对所述第一分离产物进行第二分离得到改质油和残渣，所述第二分离选自减压蒸馏、溶剂萃取或它们的组合；
4)对步骤3)所得的改质油进行加氢改质，得到加氢改质油；
5)对步骤4)所得的加氢改质油进行第三分离，得到加氢改质重油；
6)对步骤5)所得的加氢改质重油进行催化裂解，得到包含低碳烯烃的催化裂解产物；以及
7)任选地，将步骤3)所得的残渣的至少一部分返回步骤1)中进行所述热转化反应。
另一方面，本申请还提供了一种由劣质油生产低碳烯烃的系统，包括热转化反应单元、第一分离单元、第二分离单元、加氢改质单元、第三分离单元和催化裂解单元，其中：
所述热转化反应单元设置为使劣质油原料在氢气存在下在其中进行热转化反应，得到转化产物；
所述第一分离单元设置为使所述转化产物在其中分离得到第一分 离产物，其中所述第一分离产物中沸点在350℃以下的组分的含量为不大于约5重量％，沸点在350-524℃之间的组分的含量为约20-60重量％；
所述第二分离单元设置为使所述第一分离产物在其中分离得到改质油和残渣，所述第二分离单元选自减压蒸馏单元、溶剂萃取单元或它们的组合；
所述加氢改质单元设置为使所述改质油在其中进行加氢改质反应，得到加氢改质油；
所述第三分离单元设置为使所述加氢改质油在其中分离得到加氢改质重油；以及
所述催化裂解单元设置为使所述加氢改质重油在其中进行催化裂解反应，得到包含低碳烯烃的催化裂解产物。
与现有技术相比，本申请的方法和系统提供以下优点中的一个或多个：
1、可以加工高金属、高沥青质含量的劣质油，并且可实现劣质油原料的高效率转化，大幅度减少残渣量。在优选情况下，劣质油原料的总体转化率可大于90重量％，乃至大于95重量％，外排残渣量可小于10重量％，乃至小于5重量％。
2、本申请提供的方法和系统优化了待进行第二分离的物料的馏程与组成，使第二分离过程易于操作。
3、本申请能够将劣质油原料高效率改质，为催化裂解单元提供富含饱和结构、且基本上不含重金属和沥青质的改质油。在优选情况下，所得改质油中重金属(以镍和钒的总重量计)的含量可小于10微克/克，乃至小于5微克/克，并且改质油中沥青质的含量可小于2.0重量％，乃至小于0.5重量％。
4、本申请能够将改质油进行进一步加工，生产化工原料-低碳烯烃，实现低碳烯烃收率大于36％。
本申请的其他特征和优点将在随后的具体实施方式部分予以详细说明。
附图说明
附图是用来提供对本申请的进一步理解，并且构成说明书的一部 分，与下面的具体实施方式一起用于解释本申请，但并不构成对本申请的限制。在附图中：
图1a显示了本申请方法和系统的一种优选实施方式的示意图；
图1b显示了本申请方法和系统的另一优选实施方式的示意图；
图2a显示了本申请方法和系统的另一优选实施方式的示意图；
图2b显示了本申请方法和系统的另一优选实施方式的示意图。
附图标记说明
1管线              2管线               3管线
4管线              5管线               6热转化反应器
7管线              8高压分离单元       9管线
10管线             11管线              12低压分离单元
13管线             14管线              15管线
16管线             17第二分离单元      18管线
19管线             20管线              21管线
22管线             23加氢改质单元      24管线
25管线             26管线              27管线
28管线             29第一反应区        30第二反应区
31管线             32汽提段            33沉降器
34旋风分离器       35集气室            36大油气管线
37待生斜管         38待生滑阀          39再生器
40旋风分离器       41烟气管道          42管线
43空气分配器       44管线              45管线
46脱气罐           47管线              48再生斜管
49再生滑阀         50管线              51管线
52管线             53管线              54管线
55管线             56管线              57管线
58分离器
具体实施方式
以下结合附图对本申请的具体实施方式进行详细说明。应当理解的是，此处所描述的具体实施方式仅用于说明和解释本申请，并不用 于限制本申请。
在本申请的上下文中，沸点、馏程(有时也称为沸程)、终馏点和初馏点或者类似参数均指的是常压(101325Pa)下的值。
在本申请的上下文中，若无特殊说明，所给压力均指表压。
本说明书中提到的所有出版物、专利申请、专利和其它参考文献均通过引用方式全文并入本文。
除非另有定义，本说明书所用的所有技术和科学术语都具有本领域技术人员常规理解的含义。在有冲突的情况下，以本说明书的定义为准。
当本说明书以词头“本领域常规使用”、“本领域常规已知”或其类似用语来导出材料、物质、方法、步骤、装置或部件等时，该词头导出的对象涵盖本申请提出时本领域常规使用或已知的那些，但也包括目前还不常用或还不普遍知道，却将变成本领域公认为适用于类似目的的那些。
在没有明确指明的情况下，本说明书内所提到的所有百分数、份数、比率等都是以重量为基准的，除非以重量为基准时不符合本领域技术人员的常规认识。
在本文中所披露的任何具体数值(包括数值范围的端点)都不限于该数值的精确值，而应当理解为还涵盖了接近该精确值的值。并且，对于所披露的数值范围而言，在该范围的端点值之间、端点值与范围内的具体点值之间，以及各具体点值之间可以任意组合而得到一个或多个新的数值范围，这些新的数值范围也应被视为在本文中具体公开。
本申请中，除了明确说明的内容之外，未提到的任何事宜或事项均直接适用本领域已知的那些而无需进行任何改变。而且，本文描述的任何实施方式均可以与本文描述的一种或多种其他实施方式自由结合，由此形成的技术方案或技术思想均视为本发明原始公开或原始记载的一部分，而不应被视为是本文未曾披露或预期过的新内容，除非本领域技术人员认为该结合明显不合理。
在第一方面，本申请提供了一种由劣质油生产低碳烯烃的方法，包括如下步骤：
1)在氢气存在下对劣质油原料进行转化反应，得到转化产物；
2)对所述转化产物进行第一分离，得到第一分离产物，其中所述 第一分离产物中沸点在350℃以下的组分的含量为不大于约5重量％，沸点在350-524℃之间的组分的含量为约20-60重量％；
3)对所述第一分离产物进行第二分离得到改质油和残渣，所述第二分离选自减压蒸馏、溶剂萃取或它们的组合；
4)对步骤3)所得的改质油进行加氢改质，得到加氢改质油；
5)对步骤4)所得的加氢改质油进行第三分离，得到加氢改质重油；
6)对步骤5)所得的加氢改质重油进行催化裂解反应，得到包含低碳烯烃的催化裂解产物；以及
7)任选地，将步骤3)所得的残渣的至少一部分返回步骤1)中进行所述转化反应。
本申请的方法在尽可能减少残渣外甩，提高资源利用率的情况下，可以维持系统的长时间运转，而转化反应和各分离步骤是决定是否能够长期运转的关键，所述转化反应的转化率对于体系稳定性控制和分离操作的稳定性至关重要。发明人经过大量实验发现，在所述转化反应中，劣质油中沸点在524℃以上的组分的转化率(本文中也称为“转化反应的转化率”)可以为约30-70重量％，优选为约30-60重量％，其中所述转化率＝(劣质油中沸点在524℃以上的组分的重量-转化产物中沸点在524℃以上的组分的重量)/劣质油中沸点在524℃以上的组分的重量×100重量％
根据本申请，步骤1)的所述转化反应本质上是热转化反应，劣质油大分子，尤其是沥青质超分子在其中发生超分子的解离、大分子的裂化反应以及S、N杂原子的脱除反应，所述热转化反应使得劣质油中沸点在524℃以上的组分的转化率达到约30-70重量％，优选达到约30-60重量％。本申请对所述转化反应的条件(包括催化剂)和反应器并无严格的限制，只要能够达到上述转化率即可。
根据本申请，所述转化反应可以在有或没有转化催化剂存在的条件下进行。在优选的实施方式中，所述转化反应在转化催化剂存在下进行，所述转化催化剂可以包含选自第VB族金属化合物、第VIB族金属化合物和第VIII族金属化合物中的至少一种，优选包含Mo化合物、W化合物、Ni化合物、Co化合物、Fe化合物、V化合物和Cr化合物中的至少一种。进一步优选地，所述转化催化剂为非负载型催化 剂，如分散型催化剂。例如，所述转化催化剂可以选自包含上述金属的硫化物的固体物质，包含上述金属的有机配合物或螯合物，或者包含上述金属的氧化物的水溶液。特别地，所述转化催化剂例如可以是辛酸钼、环烷酸钼、环烷酸镍、环烷酸钨、油酸铁、二烷基硫代甲酸钼等这类有机-金属配合物/螯合物中的一种或多种；或者是含有上述金属的氧化物和/或硫化物的固体粉末，如赤铁矿、辉钼矿、硫化钼、硫化铁等中的一种或多种；或者是含有上述金属的氧化物和/或能够分解生成上述金属氧化物的无机酸盐的水溶液，如钼酸铵、硫酸钼、硝酸钼、硝酸镍、硝酸钴、氧化钼、氧化铁、氧化镍、氧化钨、氧化钒等的水溶液。转化催化剂在反应体系中为高度分散的形式，其粒度为约2nm至约50μm，优选约2nm至约1μm。
在优选的实施方式中，步骤1)的所述转化反应在浆态床反应器中进行，其中液体反应原料在以固体悬浮物形式存在的催化剂作用下反应。
在优选的实施方式中，所述转化反应可以在如下条件下进行：温度为约380-470℃，优选为约400-440℃；氢分压为约10-25MPa，优选为约13-20MPa；劣质油的体积空速为约0.01-2h -1，优选为约0.1-1.0h -1；氢气与劣质油的体积比为约500-5000，优选为约800-2000，以所述转化催化剂中的活性金属计、并以劣质油的重量为基准，所述转化催化剂的用量为约10-50000微克/克，优选为约30-25000微克/克。
根据本申请，所述劣质油可选自含沥青质的低品质原料油，其中沥青质是指原料油中不溶于非极性的小分子正构烷烃(比如正戊烷或者正庚烷)而溶于苯或者甲苯的物质。优选地，所述劣质油满足以下指标中的一项或多项：API度小于约27、沸点大于约350℃(优选为大于约500℃，更优选大于约524℃)、沥青质含量大于约2重量％(优选大于约5重量％，更优选大于约10重量％，进一步优选大于约15重量％)、以及以镍和钒的总重量计的重金属含量大于约100微克/克。在某些具体实施方式中，所述劣质油可以选自劣质原油、重油、脱油沥青、煤衍生油、页岩油和石化废油中的至少一种。本领域技术人员所熟知的其它低品质原料油也可以单独或混合后作为所述劣质油原料进行转化反应，本申请不再赘述。
根据本申请，所述“劣质原油”例如可以为稠油，其中“稠油” 是指沥青质和胶质含量较高、黏度较高的原油，一般将地面20℃密度大于0.943克/厘米 3、地下原油黏度大于50厘泊的原油叫稠油。
根据本申请，所述“重油”是指沸点在350℃以上的馏分油或者渣油，其中“馏分油”一般指的是原油或二次加工油经常压精馏和减压精馏得出的馏分产品，比如重柴油、重瓦斯油、润滑油馏分或者裂化原料等；“渣油”是指原油经过常减压蒸馏得到的塔底馏出物，一般将常压蒸馏塔底馏出物称为常压渣油(一般为沸点大于350℃的馏分)，一般将减压蒸馏塔底馏出物称为减压渣油(一般为沸点大于500℃或524℃的馏分)。所述渣油可以为选自拔头原油、由油砂沥青得到的重油和初馏点大于350℃的重油中的至少一种，其中“拔头原油”是指在常减压蒸馏工艺中对原油进行分馏时，从初馏塔的塔底或者闪蒸塔的塔底排出的油。
根据本申请，所述“脱沥青油”是指原料油在溶剂脱沥青装置中，通过与溶剂接触、溶解分离、萃取塔塔底得到的富沥青质、富含芳香组分的萃佘物，根据溶剂种类的不同，可分为丙烷脱油沥青、丁烷脱油沥青、戊烷脱油沥青等。
根据本申请，所述“煤衍生油”是指以煤为原料，经过化学加工得到的液体燃料，可以为选自煤液化产生的煤液化油和煤热解生成的煤焦油中的至少一种。
根据本申请，所述“页岩油”是指将油母页岩经低温干馏或其它热处理得到的合成原油，或呈褐色黏稠状膏状物，其有刺激性臭味，氮含量较高。
根据本申请，所述“所述石化废油”可以为选自石化废油泥、石化油渣及其炼制产品中的至少一种。
根据本申请，在步骤2)中对所述转化产物进行第一分离，得到第一分离产物，其中所述第一分离产物中沸点在350℃以下的组分的含量不大于约5重量％，优选小于约3重量％，沸点在350-524℃(优选为355-500℃或380-524℃，进一步优选为400-500℃)之间的组分的含量为约20-60重量％，优选为约25-55重量％。优选地，所述第一分离产物的初馏点不低于约300℃，优选不低于约330℃，更优选不低于约350℃。
根据本申请，所述第一分离产物一般由转化产物中沸点较高的组 分所组成，其包括步骤3)中得到的残渣和改质油，所述残渣的主要成分是沥青质，其中也包括保持流动性所必须的一些胶质和芳香分组分；所述改质油可以作为后续处理的优质原料进行加工得到其它油品。转化产物中其余沸点较低的组分可以在步骤2)中与所述第一分离产物分离，例如标准状态下的气体产物(例如干气和液化气等)以及沸点在350℃以下的其它组分。
根据本申请，步骤2)的所述第一分离用于获得符合上述馏程组成的第一分离产物，本申请对其具体实施方式并无特别限制。在某些具体实施方式中，所述第一分离为物理分离，比如萃取、蒸馏、蒸发、闪蒸和冷凝等。
在优选的实施方式中，步骤2)的所述第一分离包括：
2a)将步骤1)所得的转化产物在第一压力和第一温度下进行分离，得到气体组分和液体组分；和
2b)将所得液体组分在第二压力和第二温度下进行分离，得到所述第一分离产物和第二分离产物，其中所述第一压力大于所述第二压力。
根据本申请，在步骤2a)中优选分离出氢气等气体产物，所得气体组分富含氢气，优选氢气含量在85重量％以上。优选地，步骤2a)中的所述第一压力可以为约10-25MPa，优选为约13-20MPa，为了方便测量，该第一压力一般指气体组分离开分离装置时的出口压力；所述第一温度可以为约380-470℃，优选为约400-440℃，为了方便测量，该第一温度一般指液体组分离开分离装置时的出口温度。所述步骤2a)的分离方式可以选自蒸馏、分馏和闪蒸等，优选为蒸馏。该蒸馏可以在蒸馏塔中进行，其中气体组分可以从蒸馏塔塔顶得到，而液体组分可以从蒸馏塔塔底得到；
根据本申请，在步骤2b)中优选分离出沸点在350℃以下的组分而尽量保留沸点在350-524℃的组分。优选地，步骤2b)的所述第二压力低于所述第一压力，优选比所述第一压力低4-24MPa，更优选低7-19MPa；具体地，所述第二压力可以为约0.1-5MPa，优选为0.1-4MPa，为了方便测量，该第二压力一般指第二分离产物离开分离装置时的出口压力；所述第二温度可以为约150-390℃，优选为200-370℃，为了方便测量，该第二温度一般指第一分离产物离开分离装置时的出口温 度。所述步骤2b)的分离方式可以为蒸馏和/或分馏，优选为常压或加压分馏，可以在常压蒸馏罐或加压蒸馏塔中进行。根据本申请，步骤2b)所得的第二分离产物可以包含在第二压力和第二温度条件下分离得到的沸点低于所述第一分离产物的轻组分。
在进一步优选的实施方式中，步骤2)的所述第一分离可进一步包括：
2c)对步骤2b)所得的第二分离产物的至少一部分进行切割，得到石脑油和常压瓦斯油；
2d)将步骤2a)所得的气体组分的至少一部分返回步骤1)中进行所述转化反应；和/或
2e)将步骤2a)所得的气体组分的至少一部分返回步骤4)中进行所述加氢改质。
根据本申请，步骤2c)的切割可以通过分馏或蒸馏来进行，优选通过分馏进行，例如在分馏塔中进行，其操作压力可以为0.05-2.0MPa，优选为约0.1-1.0MPa，操作温度可以为50-350℃，优选为150-330℃。
根据本申请，步骤2d)和步骤2e)将步骤2a)中所得气体组分的至少一部分返回步骤1)和/或步骤4)中，该气体组分可以直接使用或经分离作为循环氢气使用。
在更进一步优选的实施方式中，步骤2)的所述第一分离可进一步包括：
2f)将步骤2b)所得的第二分离产物的至少一部分和/或步骤2c)所得的常压瓦斯油的至少一部分返回步骤4)中，与所述改质油一起进行加氢改质。
根据本申请，步骤3)的所述第二分离用于将第一分离产物中的易于加工的改质油与残渣分离，所得残渣外甩或在步骤7)中返回步骤1)进行转化反应。在具体实施方式中，步骤3)的所述第二分离可以在第三温度和第三压力下、采用减压蒸馏和溶剂萃取两者中的一个或多个进行。具体地，所述减压蒸馏可以在带有或不带填料的蒸馏塔中进行，此时所述第三压力为真空度约1-20mmHg，第三温度为约250-350℃。所述溶剂萃取优选为萃取溶剂与第一分离产物逆流接触萃取，其可以在任意萃取装置中进行，例如在萃取塔中进行，此时所述第三压力可以为约3-12MPa，优选为约3.5-10MPa，第三温度可以为约55-300℃， 优选为约70-220℃，萃取溶剂可以为C 3-C 7烃，优选为C 3-C 5烷烃和C 3-C 5烯烃中至少一种，进一步优选为C 3-C 4烷烃和C 3-C 4烯烃中至少一种，所述萃取溶剂与所述第一分离产物的重量比为约1∶1至约7∶1，优选为约1.5∶1至约5∶1。本领域技术人员也可以采取其它常规的萃取方式进行萃取，本申请不再赘述。
根据本申请，步骤3)所得的残渣是转化产物中沸点最高的组分，其软化点越高则转化产物中易于加工的组分分离得越完全，但是为了维持残渣在管线输送时的流动性以及返回转化反应器时的溶解性，步骤3)所得的残渣的软化点优选小于约150℃，更优选小于约120℃。
根据本申请，当所述转化反应在浆态床反应器中进行时，其中的转化催化剂会随着转化产物一起进入后续分离步骤并保留在残渣中，随着催化剂加入量的增加以及劣质油中金属组分的累积，整个反应系统中的金属会不断增加。为了维持系统中金属的平衡，需要间断或持续地将残渣进行外排，优选将部分残渣外甩，外甩的残渣占所述残渣总量的比例优选为约5-70重量％，更优选约10-50重量％；同时为了使劣质油得到充分使用，优选在步骤7)中将部分残渣返回步骤1)，返回的残渣比例优选为约30-95重量％，更优选约50-90重量％。本领域技术人员也可以根据劣质油的不同金属含量来调整残渣外甩和循环的比例，本申请不再赘述。
根据本申请，为了方便生产化工原料-低碳烯烃，在步骤4)中对所得改质油进行加氢改质，在步骤5)中将得到的加氢改质油切割分离为加氢改质轻油和加氢改质重油，加氢改质轻油和加氢改质重油之间的切割点可以为约340-360℃，优选为约345-355℃，更优选为约350℃；以及，在步骤6)中对所得加氢改质重油进行催化裂解，得到包含低碳烯烃的催化裂解产物。所述催化裂解产物经分离可以得到干气、低碳烯烃、汽油、循环油和油浆。所述“循环油”一般包括轻循环油和重循环油，轻循环油也可以称为柴油，是指催化裂化反应所得沸点介于205℃至350℃之间的馏分，重循环油是指沸点介于343℃至500℃之间的馏分；所述“油浆”一般是指催化裂化反应的分馏步骤获得的塔底油，经过沉降器分离后，从沉降器底部排出的产品，而从沉降器上部排出的产品一般称为澄清油。
任选地，可以将所得油浆返回步骤1)进行转化反应；将所得C3 烃和C4烃经过烷-烯烃分离后将C3、C4烷烃送至步骤3)中作为萃取溶剂使用；和/或，对所得循环油单独进行加氢改质或者将其与所述改质油一起进行加氢改质。本申请方法能够实现将油浆返回进行转化反应，一方面可提高原料利用率，将低附加值的油浆转化为高附加值的富芳烃的汽油产品；另一方面，由于油浆中富含芳烃组分，可提高转化单元稳定性，延长装置操作周期。同时，也可以在步骤6)中，将至少一部分步骤2b)所得的第二分离产物和/或步骤2c)所得的常压瓦斯油与所述加氢改质重油一起进行催化裂解。步骤6)和上述步骤可以实现由劣质油最大化地生产化工原料，提高改质油和第二分离产物的利用率。
根据本申请，步骤4)的所述加氢改质是本领域技术人员所熟知的，可以按照本领域已知的任何方式进行，并没有特别的限制，可以在本领域已知的任何加氢处理装置(比如固定床反应器、流化床反应器)中进行，本领域技术人员可以对此进行合理选择。例如，所述加氢改质可以在如下条件下进行：氢气分压为约5.0-20.0MPa，优选为约8-15MPa；反应温度为约330-450℃，优选为约350-420℃；体积空速为约0.1-3h -1，优选为约0.3-1.5h -1；氢油体积比为约300-3000，优选为约800-1500；所述加氢改质采用的催化剂可包括加氢精制催化剂和/或加氢裂化催化剂。作为所述加氢精制催化剂和加氢裂化催化剂，可以举出本领域为此目的而常规使用的任何催化剂，或者可以按照本领域常规已知的任何制造方法进行制造，而且所述加氢精制催化剂和加氢裂化催化剂在所述步骤中的用量可以参照本领域的常规认识，并没有特别的限制。
具体举例而言，所述加氢精制催化剂可包括载体和活性金属组分，所述活性金属组分选自第VIB族金属和/或第VIII族非贵金属，特别是镍与钨的组合，镍、钨与钴的组合，镍与钼的组合，或者钴与钼的组合。这些活性金属组分可以单独使用一种，或者以任意的比例组合使用多种。另外，作为所述载体，比如可以举出氧化铝、二氧化硅和无定形硅铝等。这些载体可以单独使用一种，或者以任意的比例组合使用多种。优选地，以加氢精制催化剂的干基重量为基准，所述加氢精制催化剂可以包括约30-80重量％的氧化铝载体，约5-40重量％的氧化钼、约5-15重量％的氧化钴和约5-15重量％的氧化镍。本领域技术人 员也可以采用具有其它组成的加氢精制催化剂。
所述加氢裂化催化剂一般包含载体、活性金属组分和裂化活性组元。更为具体举例而言，作为所述活性金属组分，比如可以举出元素周期表第VIB族金属的硫化物、元素周期表第VIII族贱金属的硫化物或者元素周期表第VIII族贵金属等，特别是Mo硫化物、W硫化物、Ni硫化物、Co硫化物、Fe硫化物、Cr硫化物、Pt和Pd等。这些活性金属组分可以单独使用一种，或者以任意的比例组合使用多种。另外，作为所述裂化活性组元，比如可以举出无定形硅铝和分子筛等。这些裂化活性组元可以单独使用一种，或者以任意的比例组合使用多种。再者，作为所述载体，比如可以举出氧化铝、氧化硅、氧化钛和活性炭等。这些载体可以单独使用一种，或者以任意的比例组合使用多种。本申请对所述载体、所述活性金属组分和所述裂化活性组元各自的含量没有特别的限定，可以参照本领域的常规认识。优选地，以加氢裂化催化剂的干基重量为基准，所述加氢裂化催化剂可以包括约3-60重量％的沸石、约10-80重量％的氧化铝、约1-15重量％的氧化镍和约5-40重量％的氧化钨，其中所述沸石为Y型沸石。本领域技术人员也可以采用具有其它组成的加氢裂化催化剂。
在一优选的实施方式中，所述加氢改质采用的催化剂同时包含加氢精制催化剂和加氢裂化催化剂两者，所述加氢精制催化剂与加氢裂化催化剂的装填体积比为约1∶1至约5∶1，按照反应物料流向，所述加氢精制催化剂装填于所述加氢裂化催化剂上游。
根据本申请，步骤6)的所述催化裂解可以在各种形式的催化裂解反应器中进行，优选在变径稀相输送床反应器和/或组合催化裂解反应器中进行。
在一优选实施方式中，步骤6)的所述催化裂解在变径稀相输送床反应器中进行，其中所述变径稀相输送床反应器自下至上包括具有不同直径的第一反应区和第二反应区，所述第二反应区的直径与第一反应区的直径之比为约1.2∶1至约2.0∶1。优选地，在所述变径稀相输送床反应器中，第一反应区内的反应条件可以包括：反应温度约500-620℃、反应压力约0.2-1.2MPa、反应时间约0.1-5.0秒、催化剂与裂解原料的重量比约5-15、水蒸汽与裂解原料的重量比约0.05∶1至约0.3∶1；第二反应区内的反应条件可以包括：反应温度约450-550℃、反应压力约 0.2-1.2MPa、反应时间约1.0-20.0秒。
在另一优选实施方式中，步骤6)的所述催化裂解在组合催化裂解反应器中进行，其中所述组合反应器具有自下至上串联连接的第一反应区和第二反应区，所述第一反应区为提升管反应器，所述第二反应区为流化床反应器，所述流化床反应器位于提升管反应器下游，与提升管反应器出口相连接，例如其可以为本领域技术人员公知的常规催化裂化提升管反应器串联流化床反应器得到的组合物反应器。具体地，所述提升管反应器可以选自等直径提升管反应器和/或等线速提升管反应器，优选使用等直径提升管。所述提升管反应器自下而上依次包括预提升段以及至少一个反应区，为了使原料油能够充分反应，并根据不同的目的产物品质需求，所述反应区可以为2-8个，优选为2-3个。优选地，在所述组合催化裂解反应器中，第一反应区内的反应条件可以包括：反应温度为约560-750℃，优选为约580-730℃，更优选为约600-700℃；反应时间为约1-10秒，优选为约2-5秒；剂油比为约1∶1至约50∶1，优选为约5∶1至约30∶1；第二反应区内的反应条件可以包括：反应温度为约550-730℃，优选为约570-720℃；重量空速为约0.5-20h -1，优选为约2-10h -1。
在进一步优选的实施方式中，可以在所述提升管反应器中注入水蒸气，所述水蒸气优选以雾化蒸汽的形式注入，所注入的水蒸气与原料油的重量比可以为约0.01∶1至约1∶1，优选为约0.05∶1至约0.5∶1。
在某些实施方式中，本申请的方法可进一步包括将催化裂解产物中的待生催化剂与反应油气分离得到待生催化剂和反应油气，然后将得到的反应油气经后续的分离系统分离为干气、液化气、汽油和柴油等馏分，并将干气和液化气经气体分离设备进一步分离得到乙烯、丙烯等。从反应产物中分离乙烯、丙烯等的方法可采用本领域的常规技术，本申请对此没有特别限制，在此不详细描述。
在某些实施方式中，本申请的方法可进一步包括将所述待生催化剂再生；且优选地，用于所述催化裂解反应的催化剂的至少一部分为再生后的催化剂，例如可以全部为再生催化剂。
在某些实施方式中，本申请的方法可进一步包括对再生得到的再生催化剂进行汽提(一般用水蒸气汽提)，以脱去气体等杂质。
根据本申请，在再生过程中，一般从再生器的底部引入含氧气体， 含氧气体例如可以为空气。引入再生器后，待生催化剂与氧气接触烧焦再生，催化剂烧焦再生后生成的烟气在再生器上部进行气固分离，烟气进入后续能量回收系统。
根据本申请，所述待生催化剂的再生操作条件可以为：再生温度为约550-750℃，优选为约600-730℃，更优选为约650-700℃；气体表观线速为约0.5-3米/秒、优选为约0.8-2.5米/秒、更优选为约1-2米/秒，待生催化剂的平均停留时间为约0.6-3分钟、优选约0.8-2.5分钟、更优选约1-2分钟。
根据本申请，适用于步骤6)的催化裂解催化剂可以为本领域常规采用的各种催化裂解催化剂。优选地，以催化剂的总重量计，所述催化裂解催化剂可以包含：约1-60重量％的沸石、约5-99重量％的无机氧化物和约0-70重量％的粘土。
根据本申请，在所述催化裂解催化剂中，所述沸石用作活性组分，优选地，所述沸石选自中孔沸石和/或大孔沸石。在优选的实施方式中，中孔沸石占沸石总重量的约50-100重量％，优选约70-100重量％，大孔沸石占沸石总重量的约0-50重量％，优选约0-30重量％。
根据本申请，所述中孔沸石和大孔沸石具有本领域通常理解的含义，即中孔沸石的平均孔径0.5-0.6nm，大孔沸石的平均孔径0.7-1.0nm。例如，所述大孔沸石可以选自由稀土Y(REY)、稀土氢Y(REHY)、不同方法得到的超稳Y、高硅Y构成的组的一种或其中两种或更多种的混合物。
在优选的实施方式中，所述中孔沸石可以选自具有MFI结构的沸石，例如ZSM系列沸石和/或ZRP沸石。任选地，可对上述中孔沸石用磷等非金属元素和/或铁、钴、镍等过渡金属元素进行改性。有关ZRP沸石更为详尽的描述可参见美国专利US5,232,675，其内容经此引用全文并入本文；ZSM系列沸石可选自ZSM-5、ZSM-11、ZSM-12、ZSM-23、ZSM-35、ZSM-38、ZSM-48和其它类似结构的沸石之中的一种或其中多种的混合物，有关ZSM-5更为详尽的描述可参见美国专利US3,702,886，其内容经此引用全文并入本文。
根据本申请，在所述催化裂解催化剂中，所述无机氧化物用作粘接剂，优选选自二氧化硅(SiO 2)和/或三氧化二铝(Al 2O 3)。
根据本申请，在所述催化裂解催化剂中，所述粘土用作基质(即 载体)，优选选自高岭土和/或多水高岭土。
在一特别优选的实施方式中，本申请的方法包括如下步骤：
1)在氢气存在下对劣质油原料进行热转化反应，得到转化产物，其中所述转化反应的转化率为约30-70重量％，所述转化率＝(劣质油中沸点在524℃以上的组分的重量-转化产物中沸点在524℃以上的组分的重量)/劣质油中沸点在524℃以上的组分的重量×100重量％；
2)对步骤1)中所得的转化产物进行第一分离，得到第一分离产物，其中所述第一分离产物中沸点在350℃以下的组分的含量为不大于约5重量％，优选小于约3重量％，沸点在350-524℃(优选为355-500℃或380-524℃，进一步优选为400-500℃)之间的组分的含量为约20-60重量％，优选为约25-55重量％，所述第一分离产物的初馏点不低于约300℃，优选不低于约330℃，更优选不低于约350℃；
3)对步骤2)所得的第一分离产物进行第二分离得到改质油和残渣，所述第二分离选自减压蒸馏、溶剂萃取或它们的组合；
4)对步骤3)所得的改质油进行加氢改质，得到加氢改质油；
5)对步骤4)所得的加氢改质油进行第三分离，得到加氢改质重油；
6)将步骤5)所得的加氢改质重油预热后进入变径稀相输送床反应器底部，与再生催化剂接触进行催化裂解反应，同时向上流动进入旋风分离器进行气固分离，分离出的反应油气经进一步分离得到包括丙烯和高辛烷值汽油的产物；分离出的待生催化剂经汽提后进入催化剂再生器中烧焦再生，再生催化剂返回反应器中循环使用；或者
将步骤5)所得的加氢改质重油预热后进入组合催化裂解反应器的第一反应区，与再生催化剂接触进行催化裂解反应，同时向上流动进入第二反应区，继续进行催化裂解反应，反应器出口的反应油气和待生催化剂进入旋风分离器进行气固分离，分离出的反应油气经进一步分离得到包含低碳烯烃的产物；分离出的待生催化剂经汽提后进入催化剂再生器中烧焦再生，再生催化剂返回反应器中循环使用，其中所述低碳烯烃包括乙烯、丙烯、丁烯；以及
7)将步骤3)所得的残渣返回步骤1)中进行所述转化反应；或者，将步骤3)所得的残渣外甩；或者，将步骤3)所得的残渣的一部分返回步骤1)中进行所述转化反应，所得残渣的剩余部分外甩。
在第二方面，本申请提供了一种由劣质油生产低碳烯烃的系统，包括转化反应单元、第一分离单元、第二分离单元、加氢改质单元、第三分离单元和催化裂解单元，其中：
所述转化反应单元设置为使劣质油原料在氢气存在下在其中进行热转化反应，得到转化产物；
所述第一分离单元设置为使所述转化产物在其中分离得到第一分离产物，其中所述第一分离产物中沸点在350℃以下的组分的含量为不大于约5重量％，沸点在350-524℃之间的组分的含量为约20-60重量％；
所述第二分离单元设置为使所述第一分离产物在其中分离得到改质油和残渣，所述第二分离单元选自减压蒸馏单元、溶剂萃取单元或它们的组合；
所述加氢改质单元设置为使所述改质油在其中进行加氢改质反应，得到加氢改质油；
所述第三分离单元设置为使所述加氢改质油在其中分离得到加氢改质重油；以及
所述催化裂解单元设置为使所述加氢改质重油在其中进行催化裂解反应，得到包含低碳烯烃的催化裂解产物。
根据本申请的某些实施方式，在所述转化反应单元中，劣质油、氢气和转化催化剂在转化反应器中反应，得到转化反应产物并送至第一分离单元。优选地，所述转化反应器为浆态床反应器。
根据本申请的某些实施方式，在所述第一分离单元中，所述转化反应产物首先分离成气体产物和液体产物，而后所述液体产物进一步分离得到馏程大于约350℃的重馏分作为所述第一分离产物并送至第二分离单元。
根据本申请的某些实施方式，在所述第二分离单元中，所述第一分离产物在减压蒸馏塔中分离，或者在萃取塔中与萃取溶剂逆流接触进行萃取分离，得到改质油和残渣，或者在减压蒸馏与萃取分离的组合中进行分离，得到改质油和残渣，所述改质油送至加氢改质单元。任选地，所述残渣返回所述转化反应单元做进一步转化。
根据本申请的某些实施方式，在所述加氢改质单元中，所述改质油在加氢处理催化剂作用下进行反应，得到加氢改质油并送至第三分 离单元。
根据本申请的某些实施方式，在所述第三分离单元中，所述加氢改质油被切割分离为加氢改质轻油和加氢改质重油，加氢改质重油送至催化裂解单元。
根据本申请的某些实施方式，所述催化裂解单元包括变径稀相输送床反应器和/或组合催化裂解反应器，其中所述变径稀相输送床反应器自下至上包括具有不同直径的第一反应区和第二反应区，第二反应区的直径与第一反应区的直径之比为约1.2∶1至约2.0∶1；所述组合催化裂解反应器自下至上包括第一反应区和第二反应区，所述第一反应区为提升管反应器，所述第二反应区为流化床反应器。
根据本申请的一种优选实施方式，在所述催化裂解单元中，催化裂解催化剂进入变径稀相输送床反应器的第一反应区的预提升段，在预提升介质的作用下向上流动，预热后的加氢改质重油与雾化蒸汽一起注入第一反应区，与再生催化剂接触进行催化裂化反应同时向上流动，并进入第二反应区继续进行反应，得到包含低碳烯烃的催化裂解产物。任选地，所述催化裂解产物在后续分离系统中分离得到乙烯、丙烯、具有高辛烷值的汽油等馏分；分离出的待生催化剂进入再生器中烧焦再生，恢复活性的再生催化剂返回变径稀相输送床反应器中循环使用。
根据本申请的另一优选实施方式，在所述催化裂解单元中，催化裂解催化剂进入组合催化裂解反应器的第一反应区的预提升段，在预提升介质的作用下向上流动，预热后的加氢改质油与雾化蒸汽一起注入第一反应区，与再生催化剂接触进行催化裂解反应同时向上流动，并进入第二反应区继续进行反应，得到包含低碳烯烃的催化裂解产物。任选地，所述催化裂解产物在后续分离系统中分离得到乙烯、丙烯和裂解汽油等馏分；分离出的待生催化剂进入再生器中烧焦再生，恢复活性的再生催化剂返回组合催化裂解反应器中循环使用。
以下结合附图对本申请的具体实施方式做进一步的详细说明。
如图1a、1b、2a和2b所示，劣质原料经管线1、转化催化剂经管线2、新鲜氢气经管线3、循环氢气经管线4、催化油浆经管线57以及残渣经管线5输送至转化反应器6中进行热转化反应。转化产物经管线7输送至高压分离单元8进行加压蒸馏，分离为气体组分和液体组 分，然后将气体组分作为循环氢气经管线9和管线4输送至转化反应器6，或作为氢源经管线9和管线11输送至加氢改质单元23。液体组分经管线10输送至低压分离单元12进行压力骤降，分离为第二分离产物和第一分离产物。第二分离产物经管线14进入加氢改质单元23，第一分离产物经管线15输送至第二分离单元17进行减压蒸馏，分离得到改质油和残渣(参见图1a和2a)，或与来自管线16或/和来自管线55的萃取溶剂逆流接触在第二分离单元17中进行萃取分离，得到改质油和残渣(参见图1b和2b)。残渣的一部分经管线19和管线20外甩，其余部分经管线19和管线5循环至转化反应器6与劣质油原料一起继续进行转化反应。或者，也可以将全部残渣经管线19和管线20外甩而不进行循环。改质油经管线18，与来自管线14的第二分离产物、以及来自管线21的催化柴油混合经管线22进入加氢改质单元23进行加氢改质，加氢改质产物经分离，轻组分及加氢改质轻油分别经管线24和管线25引出，或加氢改质轻油经管线25与经管线26排出的加氢改质重油混合经管线28送入催化裂解单元(图1a和1b所示的变径稀相输送床反应器或图2a和2b所示的组合催化裂解反应器)第一反应区29中。同时，预提升介质经管线50也进入第一反应区29，来自管线48的再生催化剂经再生滑阀49调节后进入第一反应区29，在预提升介质的提升作用下沿提升管向上加速运动，预热后的加氢改质油经管线28与来自管线27的雾化蒸汽一起注入第一反应区29，与第一反应区29已有的物流混合，原料油在热的催化剂上发生催化裂解反应，并向上加速运动，进入催化裂解单元的第二反应区30继续进行反应。生成的反应产物油气和失活的待生催化剂进入沉降器33中的旋风分离器34，实现待生催化剂与反应产物油气的分离，反应产物油气进入集气室35，催化剂细粉返回沉降器。沉降器中待生催化剂流向汽提段32，与来自管线31的蒸汽接触。从待生催化剂中汽提出的反应产物油气经旋风分离器后进入集气室35。汽提后的待生催化剂经待生滑阀38调节后进入再生器39，来自管线44的空气经空气分配器43分配后进入再生器39，烧去位于再生器39底部的密相床层中待生催化剂上的焦炭，使失活的待生催化剂再生，烟气经旋风分离器40的上部气体烟气管道41进入后续能量回收系统。其中，所述预提升介质可以为干气、水蒸气或它们的混合物。
再生后的催化剂经与再生器39催化剂出口连通的管线45进入脱气罐46，与来自脱气罐46底部的管线47的汽提介质接触，脱除再生催化剂夹带的烟气，脱气后的再生催化剂经管线48循环到第一反应区29的底部，可以通过再生滑阀49控制催化剂循环量，气体经管线42返回再生器39内，集气室35中的反应产物油气经过大油气管线36进入后续分离系统58，经分离得到的H 2、C1-C2烷烃从管线53引出，得到的低碳烯烃(包括C2、C3、C4烯烃)经管线54送出系统；C3、C4烷烃经管线55送出系统或送入第二分离单元17作为萃取溶剂使用，得到的富含芳烃的汽油作为产品从管线56引出，得到的循环油从管线21引出并与来自管线18的改质油、来自管线14的第二分离产物混合后一起送入加氢改质单元23进行加氢改质，得到的油浆经管线57引出并返回转化反应器6中进行热转化反应。
任选地，如图1a和1b所示，可将催化裂解产物分离得到的C4或轻汽油馏分通过管线52与水蒸气一起经由管线51送入作为催化裂解单元的变径稀相输送床反应器的第二反应区30进行回炼，以进一步裂解增产低碳烯烃。
在某些优选的实施方式中，本申请提供了如下的技术方案：
A1、一种劣质油生产低碳烯烃的方法，该方法包括：
(1)劣质油进入转化反应单元进行转化反应，生成的反应产物经分离得到沸点大于约350℃的重馏分；
(2)将该重馏分送入减压蒸馏分离单元或/和萃取分离单元中进行分离，得到改质油和残渣；
(3)将改质油送入加氢改质单元进行加氢改质，得到加氢改质油；
(4)预热后的加氢改质油进入催化裂解反应器的第一反应区，与再生催化剂接触进行催化裂解反应同时向上流动进入第二反应区，继续进行催化裂解反应，反应器出口的反应油气和待生催化剂进入旋风分离器进行气固分离，分离出的反应油气引出装置，进一步分离得到包含低碳烯烃的产物；分离出的待生催化剂经汽提后进入催化剂再生器中烧焦再生，再生催化剂返回反应器中循环使用。
A2、根据项目A1所述的方法，所述劣质油包括选自劣质原油、重油、脱油沥青、煤衍生油、页岩油和石化废油中的至少一种。
A3、根据项目A1所述的方法，所述改质原料满足选自API度小 于约27、馏程大于约350℃、沥青质含量大于约2重量％以及以镍和钒的总重量计的重金属含量大于约100微克/克中的一项或多项指标。
A4、根据项目A1所述的方法，其特征在于，转化反应单元的转化反应器为流动床反应器。
A5、根据项目A1所述的方法，其特征在于，所述转化反应单元的转化催化剂含有选自第VB族金属化合物、第VIB族金属化合物和第VIII族金属化合物中的至少一种。
A6、根据项目A1所述的方法，所述转化反应单元的反应条件包括：温度为约380-470℃，氢分压为10-25兆帕，劣质油体积空速为约0.01-2小时 -1，氢气与劣质油的体积比为约500-5000，以所述转化催化剂中金属计并以劣质油的重量为基准，所述转化催化剂的用量为约10-50000微克/克。
A7、根据项目A1所述的方法，所述的萃取分离单元的操作条件包括：压力为约3-12兆帕，温度为约55-300℃，萃取溶剂为C3-C7烃，溶剂与重馏分的重量比为(1-7)∶1，或者
所述的减压蒸馏分离单元的操作条件包括：约1-20mmHg的真空度和约250-350℃的温度。
A8、根据项目A1所述的方法，所述的加氢改质单元反应条件包括：氢分压为约5.0-20.0兆帕，反应温度为约330-450℃，体积空速为约0.1-3小时 -1，氢油体积比为约300-3000。
A9、根据项目A1所述的方法，所述加氢改质单元所用的催化剂包括加氢精制催化剂和加氢裂化催化剂，所述加氢精制催化剂包括载体和活性金属组分，所述活性金属组分选自第VIB族金属和/或第VIII族非贵金属；所述加氢裂化催化剂包括沸石、氧化铝、至少一种第VIII族金属组分和至少一种第VIB族金属组分。
A10、根据项目A1所述的方法，所述加氢裂化催化剂以催化剂为基准，其组成为：沸石3～60重量％，氧化铝10～80重量％，氧化镍1～15重量％，氧化钨5～40重量％。
A11、根据项目A1所述的方法，催化裂解单元的反应器包括第一反应区和第二反应区，所述第一反应区为提升管反应器，所述第二反应区为流化床反应器。
A12、根据项目A1所述的方法，所述第一反应区条件包括：反应 温度为560～750℃，时间为1～10秒，剂油比为1～50∶1；所述第二反应区条件包括：反应温度为550～700℃，空速为约0.5-20h -1。
A13、根据项目A1中的方法，其特征在于，步骤(4)中所述催化剂含有：沸石1～60重量％、无机氧化物5～99重量％和粘土0～70重量％，均以该催化剂的总重量计，其中沸石选自中孔沸石和任选的大孔沸石，中孔沸石占沸石总重量的50～100重量％，大孔沸石占沸石总重量的0～50重量％。
A14、根据项目A13的方法，其特征在于，所述的中孔沸石占沸石总重量的70～100重量％，大孔沸石占沸石总重量的0～30重量％。
A15、根据项目A1的方法，其特征在于，步骤(2)所述的残渣返回步骤(1)中进行所述转化反应；或者，将步骤(2)中所得的残渣进行外甩；或者，将部分步骤(2)中所得的残渣返回步骤(1)中进行所述转化反应，剩余部分残渣进行外甩。
A16、根据项目A1所述的方法，其中，所述转化反应的转化率为约30-70重量％，所述转化反应的转化率＝(劣质油中馏程在524℃以上组分的重量-转化产物中馏程在524℃以上组分的重量)/劣质油中馏程在524℃以上组分的重量×100重量％；和/或所述重馏分中，馏程在350-524℃之间组分的含量为约20-60重量％。
A17、一种劣质油生产低碳烯烃的系统，该系统包括转化反应单元、萃取或减压蒸馏分离单元、加氢改质单元和催化裂解单元，其中转化反应单元与减压蒸馏或/和萃取分离单元相连，减压蒸馏或/和萃取分离单元与加氢改质单元相连，加氢改质单元和催化裂解单元相连。
B1、一种由低品质油生产低碳烯烃的改质方法，该方法包括：
(1)将作为改质原料的低品质油进行临氢转化反应，得到临氢转化产物；其中，所述临氢转化反应的转化率为约30-70重量％，所述临氢转化反应的转化率＝(改质原料中沸点在524℃以上组分的重量-临氢转化产物中沸点在524℃以上组分的重量)/改质原料中沸点在524℃以上组分的重量×100重量％；
(2)将步骤(1)中所得临氢转化产物进行分离处理，至少得到第一分离产物；其中，所述第一分离产物中，沸点在350℃以下组分的含量不大于约5重量％，沸点在350-524℃之间组分的含量为约20-60重量％；
(3)将步骤(2)中所得第一分离产物在减压蒸馏分离单元中通过减压蒸馏进行分离或/和在萃取分离单元中采用萃取溶剂进行萃取分离，得到改质油和残渣；
(4)将步骤(3)中所得的残渣返回步骤(1)中进行所述临氢转化反应；或者，将步骤(3)中所得的残渣进行外甩；或者，将部分步骤(3)中所得的残渣返回步骤(1)中进行所述临氢转化反应，剩余部分残渣进行外甩；
(5)将步骤(3)中所得改质油进行加氢改质，得到加氢改质油；
(6)将步骤(5)所得加氢改质油进行分离，所得加氢改质重油进行催化转化反应，得到包含低碳烯烃的产品。
B2、根据项目B1所述的改质方法，步骤(1)中，所述临氢转化反应的转化率为约30-60重量％。
B3、根据项目B1所述的改质方法，步骤(1)中，所述临氢转化反应在浆态床反应器中进行。
B4、根据项目B1所述的改质方法，步骤(1)中，所述临氢转化反应在临氢转化催化剂存在或不存在的条件下进行，所述临氢转化催化剂含有选自第VB族金属化合物、第VIB族金属化合物和第VIII族金属化合物中的至少一种。
B5、根据项目B1所述的改质方法，步骤(1)中，所述临氢转化反应的条件包括：温度为约380-470℃，氢分压为10-25兆帕，改质原料的体积空速为约0.01-2小时 -1，氢气与改质原料的体积比为约500-5000，以所述临氢转化催化剂中金属计并以改质原料的重量为基准，所述转化催化剂的用量为约10-50000微克/克。
B6、根据项目B1所述的改质方法，步骤(1)中，所述改质原料包括选自劣质原油、重油、脱油沥青、煤衍生油、页岩油和石化废油中的至少一种。
B7、根据项目B1所述的改质方法，所述改质原料满足选自API度小于约27、沸点大于约350℃、沥青质含量大于约2重量％、以及以镍和钒的总重量计的重金属含量大于约100微克/克中的一项或多项指标。
B8、根据项目B1所述的改质方法，步骤(2)中，所述第一分离产物中，沸点在350℃以下组分的含量小于约3重量％，沸点在350-524 ℃之间组分的含量为约25-55重量％。
B9、根据项目B1所述的改质方法，步骤(2)中，所述分离处理包括：
(2-1)将步骤(1)中所得临氢转化产物在第一压力和第一温度下进行分离，得到气体组分和液体组分；
(2-2)将液体组分在第二压力和第二温度下进行分离，得到所述第一分离产物和第二分离产物；其中，所述第一压力大于所述第二压力。
B10、根据项目B9所述的改质方法，其中，所述第一压力为10-25兆帕，第一温度为约380-470℃；所述第二压力为约0.1-5兆帕，第二温度为约150-390℃。
B11、根据项目B9所述的改质方法，其中，所述分离处理还包括：
(2-3)将步骤(2-2)中所得第二分离产物进行切割，得到石脑油和常压瓦斯油；和/或
(2-4)将步骤(2-1)中所得气体组分返回步骤(1)中进行临氢转化反应和/或步骤(5)进行加氢改质。
B12、根据项目B11所述的改质方法，其中，将所述第二分离产物和/或常压瓦斯油与所述改质油一起进行加氢改质。
B13、根据项目B1或12所述的改质方法，其中，步骤(5)所述加氢改质的条件包括：氢气分压为约5.0-20.0兆帕，反应温度为约330-450℃，体积空速为约0.1-3小时 -1，氢油体积比为约300-3000。
B14、根据项目B1或12所述的改质方法，其中，步骤(5)所述加氢改质所用的催化剂包括加氢精制催化剂和加氢裂化催化剂，所述加氢精制催化剂包括载体和活性金属组分，所述活性金属组分选自第VIB族金属和/或第VIII族非贵金属；所述加氢裂化催化剂包括沸石、氧化铝、至少一种第VIII族金属组分和至少一种第VIB族金属组分。
B15、根据项目B14所述的改质方法，其中，以加氢裂化催化剂的干基重量为基准，所述加氢裂化催化剂包括约3-60重量％的沸石、约10-80重量％的氧化铝、约1-15重量％的氧化镍和约5-40重量％的氧化钨。
B16、根据项目B1所述的改质方法，其中，步骤(6)所述催化转化反应在催化转化催化剂存在的条件下在催化转化反应器中进行，其 中，催化转化反应器选自提升管反应器、流化床反应器、下行式输送线反应器、移动床反应器中任一种反应器或任两种反应器组合的复合反应器。
B17、根据项目B1所述的改质方法，其中，步骤(6)所述催化转化反应的条件包括：反应温度为500-750℃，反应压力为0.15-0.50兆帕，反应时间为0.2-10秒，剂油比为5-40，水油比为0.05-1.0。
B18、根据项目B1所述的改质方法，其中，步骤(6)所述催化转化催化剂包括沸石、无机氧化物和任选的粘土，各组分的含量分别为：沸石1～60重量％，无机氧化物5～99重量％，粘土0～70重量％，其中沸石为中孔沸石和任选的大孔沸石的混合物，中孔沸石比例为50～100重量％，优选70～100重量％，大孔沸石的比例为0～50重量％，优选0～30重量％。
B19、根据项目B1所述的改质方法，步骤(3)中，所述萃取分离在第三温度和第三压力下的萃取溶剂中进行；其中，所述第三压力为约3-12兆帕，第三温度为约55-300℃，所述萃取溶剂为C 3-C 7烃，所述萃取溶剂与所述第一分离产物的重量比为(1-7)：1。
B20、根据项目B1所述的改质方法，步骤(3)中，所述残渣的软化点小于约150℃。
B21、根据项目B1所述的改质方法，步骤(4)中，返回步骤(1)中的残渣占残渣总量的比例为30～95重量％，优选50～90重量％。
B22、根据项目B1所述的改质方法，步骤(6)中，将加氢改质油切割分离为加氢改质轻油和加氢改质重油，加氢改质轻油和加氢改质重油之间的切割点为340℃～360℃，优选为约345-355℃，更优选为约350℃。
B23、一种由低品质油生产低碳烯烃的改质系统，该系统包括临氢转化反应单元、减压蒸馏或/和萃取分离单元、加氢改质单元和催化转化单元，其中临氢转化反应单元与减压蒸馏或/和萃取分离单元相连，减压蒸馏或/和萃取分离单元与加氢改质单元相连，加氢改质单元和催化转化单元相连。
C1、一种劣质油生产丙烯和高辛烷值汽油的方法，该方法包括：
(1)劣质油进入转化反应单元进行转化反应，生成的反应产物经分离得到馏程大于约350℃的重馏分；
(2)将该重馏分送入减压蒸馏分离单元或/和萃取分离单元中进行分离，得到改质油和残渣；
(3)将改质油送入加氢改质单元进行加氢改质，得到加氢改质油；
(4)预热后的加氢改质油进入变径稀相输送床反应器底部，与再生催化剂接触进行催化裂化反应同时向上流动进入旋风分离器进行气固分离，分离出的反应油气引出装置，进一步分离得到包含丙烯、高辛烷值汽油的产物；分离出的待生催化剂经汽提后进入催化剂再生器中烧焦再生，再生催化剂返回反应器中循环使用。
C2、根据项目C1所述的方法，所述劣质油包括选自劣质原油、重油、脱油沥青、煤衍生油、页岩油和石化废油中的至少一种。
C3、根据项目C1所述的方法，所述改质原料满足选自API度小于约27、馏程大于约350℃、沥青质含量大于约2重量％、以及以镍和钒的总重量计的重金属含量大于约100微克/克中的一项或多项指标。
C4、根据项目C1所述的方法，其特征在于，转化反应单元的转化反应器为浆态床反应器。
C5、根据项目C1所述的方法，其特征在于，所述转化反应单元的转化催化剂含有选自第VB族金属化合物、第VIB族金属化合物和第VIII族金属化合物中的至少一种。
C6、根据项目C1所述的方法，所述转化反应单元的反应条件包括：温度为约380-470℃，氢分压为10-25兆帕，劣质油体积空速为约0.01-2小时 -1，氢气与劣质油的体积比为约500-5000，以所述转化催化剂中金属计并以改质原料的重量为基准，所述转化催化剂的用量为约10-50000微克/克。
C7、根据项目C1所述的方法，所述的萃取分离单元的反应条件包括：压力为约3-12兆帕，温度为约55-300℃，萃取溶剂为C3-C7烃，溶剂与重馏分的重量比为(1-7)∶1，或者
所述的减压蒸馏分离单元的操作条件包括：约1-20mmHg的真空度和约250-350℃的温度。
C8、根据项目C1所述的方法，所述的加氢改质单元反应条件包括：氢分压为约5.0-20.0兆帕，反应温度为约330-450℃，体积空速为约0.1-3小时 -1，氢油体积比为约300-3000。
C9、根据项目C1所述的方法，所述的加氢改质单元所用的催化剂 包括加氢精制催化剂和加氢裂化催化剂，所述加氢精制催化剂包括载体和活性金属组分，所述活性金属组分选自第VIB族金属和/或第VIII族非贵金属；所述加氢裂化催化剂包括沸石、氧化铝、至少一种第VIII族金属组分和至少一种第VIB族金属组分。
C10、根据项目C1所述的方法，所述加氢裂化催化剂以催化剂为基准，其组成为：沸石3～60重量％，氧化铝10～80重量％，氧化镍1～15重量％，氧化钨5～40重量％。
C11、根据项目C1所述的方法，所述变径稀相输送床包含两个反应区，所述第二反应区的直径与第一反应区的直径比为1.2～2.0∶1。
C12、按照项目C1的方法，其特征在于所述的变径稀相输送床中第一反应区的反应条件包括：反应温度500～620℃、反应压力0.2～1.2兆帕、反应时间0.1～5.0秒、催化剂与原料的重量比5～15、水蒸汽与原料油重量比0.05～0.3∶1。
C13、按照项目C1的方法，其特征在于所述的变径稀相输送床中第二反应区的反应条件包括：反应温度450～550℃、反应压力0.2～1.2兆帕、反应时间1.0～20.0秒。
C14、根据项目C1中的方法，其特征在于，以催化剂的总重量计，所述催化剂含有：沸石1～60重量％、无机氧化物5～99重量％和粘土0～70重量％，其中沸石选自中孔沸石和任选的大孔沸石，中孔沸石占沸石总重量的50～100重量％，大孔沸石占沸石总重量的0～50重量％。
C15、根据项目C14的方法，其特征在于，所述的中孔沸石占沸石总重量的70～100重量％，大孔沸石占沸石总重量的0～30重量％。
C16、根据项目C1的方法，其特征在于，步骤(2)所述的残渣返回步骤(1)中进行所述转化反应；或者，将步骤(2)中所得的残渣进行外甩；或者，将部分步骤(2)中所得的残渣返回步骤(1)中进行所述转化反应，剩余部分残渣进行外甩。
C17、根据项目C1所述的方法，其中，所述转化反应的转化率为约30-70重量％，所述转化反应的转化率＝(劣质油中馏程在524℃以上组分的重量-转化产物中馏程在524℃以上组分的重量)/劣质油中馏程在524℃以上组分的重量×100重量％；和/或所述重馏分中，馏程在350-524℃之间组分的含量为约20-60重量％。
C18、一种劣质油生产丙烯和高辛烷值汽油的系统，该系统包括转 化反应单元、减压蒸馏或/和萃取分离单元、加氢改质单元和催化裂化单元，其中转化反应单元与减压蒸馏或/和萃取分离单元相连，减压蒸馏或/和萃取分离单元与加氢改质单元相连，加氢改质单元和催化裂化单元相连。
实施例
以下采用实施例进一步详细地说明本申请，但本申请并不限于这些实施例。
在本申请的上下文中以及包括在以下的实施例和对比例中：
油品中重金属(以镍和钒的总重量计)含量的测定方法按照ASTMD5708进行；
油品中沥青质含量的测定方法按照SH/T 0266-92(1998)进行；
残渣外甩率＝外甩残渣重量/劣质油重量×100重量％；
劣质油转化率＝(1-残渣外甩率)×100重量％；
LPG+液体产物收率＝(转化产物经第一分离和第二分离得到的LPG和液体产物的总重量)/劣质油重量×100重量％；
第一分离产物收率＝第一分离产物重量/劣质油重量×100重量％
甲苯不溶物收率＝(第一分离产物中甲苯不溶物重量/第一分离产物重量)×第一分离产物收率×100％；
残渣软化点采用GB/T 4507-84方法进行测定；
特别组分是指第一分离产物中沸点在350-524℃之间的组分；
乙烯收率＝乙烯重量/催化裂解原料油重量×100重量％；
丙烯收率＝丙烯重量/催化裂解原料油重量×100重量％；
丁烯收率＝丁烯重量/催化裂解原料油重量×100重量％；
低碳烯烃收率＝乙烯收率+丙烯收率+丁烯收率。
以下的实施例和对比例，按照附图所示的实施方式进行操作。
在以下的实施例和对比例中，所用的劣质油B为减压渣油，其性质见表1。
表1实施例和对比例中所用原料的性质
名称 劣质油B
密度(20℃)/(千克/米 3) 1060.3
API度 1.95
残炭值/重量％ 23.2
元素含量/重量％  
碳 84.62
氢 10.07
硫 4.94
氮 0.34
氧 /
四组分组成/重量％  
饱和分 9.0
芳香分 53.8
胶质 24.5
沥青质 12.7
金属含量/(微克/克)  
Ca 2.4
Fe 23.0
Ni 42.0
V 96.0
＞524℃组分含量/重量％ 100
实施例1和3
在中型装置上，以劣质油B为原料在浆态床反应器中进行转化反应，然后进行第一分离，第一分离在两个分馏塔中进行，得到第一分离产物和第二分离产物。第一分离产物再进行第二分离(实施例1中为图1b和2b所示的萃取分离，实施例3中为图1a和2a所示的减压蒸馏)，得到改质油和残渣。各步骤的具体条件和结果列于表2-1和表2-2。
实施例2和4
在中型装置上，以劣质油B为原料在浆态床反应器中进行转化反应，然后进行第一分离，第一分离在两个分馏塔中进行，得到第一分离产物和第二分离产物。第一分离产物再进行第二分离(实施例2中 为图1b和2b所示的萃取分离，实施例4中为图1a和2a所示的减压蒸馏)，得到改质油和残渣。
所得到的残渣的一部分循环，其余部分外甩。将循环的残渣与劣质油B混合后进行转化反应，其后依次进行第一分离、第二分离，得到改质油和残渣。将获得的第二分离产物进行分离，得到石脑油馏分和常压瓦斯油。各步骤的具体条件和结果列于表2-1和表2-2。
对比例1
与实施例1的基本流程相同，只是不进行转化反应和第一分离，各步骤的具体条件和结果列于表2-1和表2-2。
对比例2
与实施例2的基本流程相同，只是不进行转化反应和第一分离，各步骤的具体条件和结果列于表2-1和表2-2。
对比例3
与实施例2的基本流程相同，只是采用了不同的转化催化剂和操作条件，各步骤的具体条件和结果列于表2-1和表2-2。
表2-1各实施例和对比例中所用的反应条件
Figure PCTCN2020077389-appb-000001
Figure PCTCN2020077389-appb-000002
*各转化催化剂来源如下：
钼酸铵：北京试剂公司，试剂纯；
辛酸钼：实验室自制，纯度大于90％；
环烷酸钼：实验室自制，纯度大于85％；
环烷酸镍：实验室自制，纯度大于90％；
赤铁矿：工业用品。
表2-2各实施例和对比例的反应结果
Figure PCTCN2020077389-appb-000003
Figure PCTCN2020077389-appb-000004
表2-2的结果显示，若劣质油不经过转化反应，直接进行萃取分离，其LPG+液体产物收率仅为34.2％，残渣收率为65.8％；若将残渣循环，则LPG+液体产物收率仅为25.1％，残渣油收率高达74.9％。
另一方面，对比例3的结果表明，当第一分离产物中沸点小于350℃的组分含量超出本申请限定的范围时，使得劣质油转化率降低12个百分点，LPG+液体产物收率下降11个百分点，同时，改质油的重金属含量达到20μg/g，甲苯不溶物收率提高约1个百分点。
实施例5-6
分别将实施例2和实施例4得到的改质油送至加氢改质单元，在加氢精制与裂化温度分别为380-386℃，体积空速0.5h -1、氢油体积比1000和氢分压15MPa下进行加氢改质，得到加氢改质油，加氢改质油经过简单分离，得到加氢改质重油，试验条件及加氢改质重油性质见表3。
对比例4-5
参照实施例5-6的流程，分别对对比例1-2得到的改质油进行加氢改质，得到加氢改质油，加氢改质油经过简单分离得到加氢改质重油，试验条件及加氢改质重油性质见表3。
表3各实施例和对比例的加氢改质条件和结果
Figure PCTCN2020077389-appb-000005
*各加氢精制/裂化催化剂均由中国石化催化剂分公司提供。
实施例7-8
分别对实施例5和实施例6得到的加氢改质重油进行催化裂解(反应器型式如图1a和1b所示)，催化裂解反应在中型装置上进行，催化 剂采用齐鲁催化剂分公司生产的商业牌号为CGP的催化剂。预热的加氢改质油进入变径稀相输送床反应器第一反应区，在反应温度535℃、反应时间1.8秒，催化剂与原料油的重量比8，水蒸气与原料油的重量比为0.10的条件下进行反应。油气混合物与催化剂继续上行进入第二反应区，在反应温度510℃，反应时间2.5秒条件下继续反应。反应油气和待生催化剂从反应器出口进入密闭式旋风分离器，反应油气和待生催化剂快速分离，反应油气在分离系统按馏程进行切割，从而得到丙烯和汽油等馏分；待生催化剂在重力作用下进入汽提段，由水蒸气汽提出待生催化剂上吸附的烃类产物，汽提后的催化剂进入到再生器，与空气接触进行再生；再生后的催化剂进入脱气罐，以除去再生催化剂吸附和携带的非烃气体杂质；脱气后的再生催化剂再返回到变径稀相输送床反应器中循环使用。催化裂解单元操作条件和产品分布列于表4。
从表4的结果可以看出，加氢改质重油的丙烯产率可达9.3重量％，汽油产率约为48.5重量％，辛烷值高达98.2。
对比例6-7
采用与实施例7-8基本相同的流程，不同之处在于原料分别为对比例4-5得到的加氢改质重油。催化裂解单元操作条件和产品分布列于表4。
从表4的结果可以看出，加氢改质重油的丙烯产率仅为6.2重量％，汽油产率仅为35.8重量％，辛烷值仅为92。
表4实施例7-8和对比例6-7的催化裂解反应条件和结果
Figure PCTCN2020077389-appb-000006
Figure PCTCN2020077389-appb-000007
实施例9-10
分别对实施例5和实施例6得到的加氢改质重油进行催化裂解(反应器型式如图2a和2b所示)，催化裂解反应在中型装置上进行，催化剂采用齐鲁催化剂分公司生产的商业牌号为MMC-2催化裂解催化剂。预热的加氢改质油进入组合催化裂解反应器的第一反应区，在提升管出口温度580℃、反应时间1.8秒，催化裂解催化剂与原料油的重量比15，水蒸气与原料油的重量比为0.25的条件下进行裂解反应。油气混合物与催化剂继续上行进入第二反应区，在反应温度565℃，床层重量空速4h -1下继续反应。反应油气和待生催化剂从反应器出口进入密闭式旋风分离器，反应油气和待生催化剂快速分离，反应油气在分离系统按馏程进行切割，从而得到乙烯、丙烯和裂解汽油等馏分；待生催化剂在重力作用下进入汽提段，由水蒸气汽提出待生催化剂上吸附的烃类产物，汽提后的催化剂进入到再生器，与空气接触进行再生；再生后的催化剂进入脱气罐，以除去再生催化剂吸附和携带的非烃气体杂质；脱气后的再生催化剂再返回到提升管反应中循环使用。催化裂 解单元操作条件和产品分布列于表5。
从表5的结果可以看出，加氢改质重油的乙烯和丙烯产率分别为4.18重量％和20.50重量％，低碳烯烃产率(乙烯产率+丙烯的产率+丁烯产率，下同)约为40.83％。
对比例8-9
采用与实施例9-10基本相同的流程，不同之处在于原料为对比例4-5得到的加氢改质重油。催化裂解单元操作条件和产品分布列于表5。
从表5的结果可以看出，加氢改质重油的乙烯和丙烯产率仅分别为3.50重量％和19.87重量％，低碳烯烃产率(乙烯产率+丙烯的产率+丁烯产率，下同)仅为34.38％。
表5实施例9-10和对比例8-9的催化裂解反应条件和结果
Figure PCTCN2020077389-appb-000008
Figure PCTCN2020077389-appb-000009
实施例11
将实施例10所得循环油中馏程小于350℃的轻循环油馏分与实施例4得到的改质油一同按照实施例6的流程进行加氢改质，加氢改质的条件和产物性质列于表6。
表6实施例11的加氢改质条件和结果
Figure PCTCN2020077389-appb-000010
实施例12-13
将实施例11得到的加氢改质重油在常规提升管反应器中进行催化 裂解，催化裂解催化剂均获自中国石化催化剂有限公司齐鲁分公司，催化裂解的条件和结果列于表7。
由表7的结果可知，轻循环油与改质油一同加氢改质后获得加氢改质重油再进行催化裂解，可获得乙烯、丙烯等低碳烯烃，实施例12的低碳烯烃收率达到36.22％，实施例13的低碳烯烃收率达到36.92％。
表7实施例12-13的催化裂解反应条件和结果
Figure PCTCN2020077389-appb-000011
实施例14
在中型装置上，参照实施例2进行操作，其中将实施例10得到的油浆循环回转化反应，与劣质油B和循环残渣混合后一起进行转化反 应，然后对转化产物进行第一分离，得到第一分离产物和第二分离产物。第一分离产物再进行第二分离(萃取分离)，得到改质油和残渣。残渣的一部分循环，其余部分外甩。各步骤的具体操作条件与实施例2相同，所得反应结果列于表8。
表8结果显示，通过油浆循环，有利于提高劣质油转化率和LPG+液体产物收率，劣质油转化率和LPG+液体产物收率分别提高了1.6个百分点和1.3个百分点，甲苯不溶物收率下降了50％。
表8实施例2和14的反应结果比较
Figure PCTCN2020077389-appb-000012
由各实施例的结果可以看出，本申请的方法和系统可以使劣质油 改质得到的LPG+液体产物收率大幅度提高，同时改善催化裂解单元的原料品质，具有明显的乙烯、丙烯产率高的优势，同时还具有高辛烷值汽油产率高的优势。
以上详细描述了本申请的优选实施方式，但是，本申请并不限于上述实施方式中的具体细节，在本申请的技术构思范围内，可以对本申请的技术方案进行多种简单变型，这些简单变型均属于本申请的保护范围。
另外需要说明的是，在上述具体实施方式中所描述的各个具体技术特征，在不矛盾的情况下，可以通过任何合适的方式进行组合，为了避免不必要的重复，本申请对各种可能的组合方式不再另行说明。
此外，本申请的各种不同的实施方式之间也可以进行任意组合，只要其不违背本申请的思想，其同样应当视为本申请所公开的内容。

Claims (20)

  1. 一种由劣质油生产低碳烯烃的方法，包括如下步骤：
    1)在氢气存在下对劣质油原料进行热转化反应，得到转化产物；
    2)对所述转化产物进行第一分离，得到第一分离产物，其中所述第一分离产物中沸点在350℃以下的组分的含量为不大于约5重量％，优选小于约3重量％，沸点在350-524℃之间的组分的含量为约20-60重量％，优选约25-55重量％；
    3)对所述第一分离产物进行第二分离，得到改质油和残渣，其中所述第二分离选自减压蒸馏、溶剂萃取或它们的组合；
    4)对步骤3)所得的改质油进行加氢改质，得到加氢改质油；
    5)对步骤4)所得的加氢改质油进行第三分离，得到加氢改质重油；
    6)对步骤5)所得的加氢改质重油进行催化裂解，得到包含低碳烯烃的催化裂解产物；以及
    7)任选地，将步骤3)所得的残渣的至少一部分返回步骤1)中进行所述热转化反应。
  2. 根据权利要求1所述的方法，其中步骤1)的所述热转化反应在浆态床反应器中进行。
  3. 根据前述权利要求中任一项所述的方法，其中步骤1)的所述热转化反应在氢气及转化催化剂存在下进行，其中所述转化催化剂包含至少一种选自第VB族金属化合物、第VIB族金属化合物和第VIII族金属化合物的化合物。
  4. 根据前述权利要求中任一项所述的方法，其中步骤1)的所述热转化反应在如下条件下进行：
    温度为约380-470℃，氢分压为约10-25MPa，劣质油的体积空速为约0.01-2h -1，氢气与劣质油的体积比为约500-5000，以所述转化催化剂中的活性金属计并以劣质油的重量为基准，所述转化催化剂的用量为约10-50000微克/克。
  5. 根据前述权利要求中任一项所述的方法，其中所述劣质油选自劣质原油、重油、脱油沥青、煤衍生油、页岩油和石化废油中的至少一种，
    优选地，所述劣质油满足以下指标中的一项或多项：API度小于约27、沸点大于约350℃、沥青质含量大于约2重量％、以及以镍和钒的总重量计的重金属含量大于约100微克/克。
  6. 根据前述权利要求中任一项所述的方法，其中步骤1)的所述热转化反应进行到获得约30-70重量％的转化率，所述转化率＝(劣质油中沸点在524℃以上的组分的重量-转化产物中沸点在524℃以上的组分的重量)/劣质油中沸点在524℃以上的组分的重量×100重量％；
    优选地，步骤1)的所述热转化反应进行到获得约30-60重量％的转化率。
  7. 根据前述权利要求中任一项所述的方法，其中步骤2)的所述第一分离包括：
    2a)将步骤1)中所得的转化产物在第一压力和第一温度下进行分离，得到气体组分和液体组分；和
    2b)将所得液体组分在第二压力和第二温度下进行分离，得到所述第一分离产物和第二分离产物，其中所述第一压力大于所述第二压力，
    优选地，所述第一压力为约10-25MPa，第一温度为约380-470℃；所述第二压力为约0.1-5MPa，第二温度为约150-390℃。
  8. 根据权利要求7所述的方法，其中步骤2)的所述第一分离进一步包括：
    2c)对步骤2b)所得的第二分离产物的至少一部分进行切割，得到石脑油和常压瓦斯油；
    2d)将步骤2a)所得的气体组分的至少一部分返回步骤1)中进行所述热转化反应；和/或
    2e)将步骤2a)所得的气体组分的至少一部分返回步骤4)中进行所述加氢改质。
  9. 根据权利要求8所述的方法，进一步包括：
    2f)将步骤2b)所得的第二分离产物的至少一部分和/或步骤2c)所得的常压瓦斯油的至少一部分返回步骤4)中，与所述改质油一起进行加氢改质。
  10. 根据前述权利要求中任一项所述的方法，其中步骤4)的所述加氢改质在如下条件下进行：
    氢气分压为约5.0-20.0MPa，反应温度为约330-450℃，体积空速为约0.1-3h -1，氢油体积比为约300-3000。
  11. 根据前述权利要求中任一项所述的方法，其中步骤4)的所述加氢改质在加氢精制催化剂和/或加氢裂化催化剂的存在下进行，所述加氢精制催化剂包括载体和活性金属组分，所述活性金属组分选自第VIB族金属和/或第VIII族非贵金属；所述加氢裂化催化剂包括沸石、氧化铝、至少一种第VIII族金属组分和至少一种第VIB族金属组分，
    优选地，以加氢精制催化剂的干基重量为基准，所述加氢精制催化剂包括约30-80重量％的氧化铝载体，约5-40重量％的氧化钼、约5-15重量％的氧化钴和约5-15重量％的氧化镍；以加氢裂化催化剂的干基重量为基准，所述加氢裂化催化剂包括约3-60重量％的沸石、约10-80重量％的氧化铝、约1-15重量％的氧化镍和约5-40重量％的氧化钨。
  12. 根据权利要求1-11中任一项所述的方法，其中步骤6)的所述催化裂解在变径稀相输送床反应器和/或组合催化裂解反应器中进行，
    其中所述变径稀相输送床反应器自下至上包括具有不同直径的第一反应区和第二反应区，第二反应区的直径与第一反应区的直径之比为约1.2∶1至约2.0∶1；所述组合催化裂解反应器自下至上包括第一反应区和第二反应区，所述第一反应区为提升管反应器，所述第二反应区为流化床反应器。
  13. 根据权利要求12所述的方法，其中：
    在所述变径稀相输送床反应器中，第一反应区内的反应条件包括：反应温度约500-620℃、反应压力约0.2-1.2MPa、反应时间约0.1-5.0秒、催化剂与裂解原料的重量比约5-15、水蒸汽与裂解原料的重量比约0.05∶1至约0.3∶1；第二反应区内的反应条件包括：反应温度约450-550℃、反应压力约0.2-1.2MPa、反应时间约1.0-20.0秒，和/或
    在所述组合催化裂解反应器中，第一反应区内的反应条件包括：反应温度为约560-750℃，时间为约1-10秒，剂油比为约1∶1至约50∶1；第二反应区内的反应条件包括：反应温度为约550-700℃，重量空速为约0.5-20h -1。
  14. 根据前述权利要求中任一项所述的方法，其中步骤6)的所述催化裂解在催化裂解催化剂存在下进行，以催化剂的重量为基准，所 述催化裂解催化剂包含约1-60重量％的沸石、约5-99重量％的无机氧化物、和约0-70重量％的粘土，其中以沸石的重量为基准，所述沸石包含约50-100重量％、优选约70-100重量％的中孔沸石，和约0-50重量％、优选约0-30重量％的大孔沸石。
  15. 根据前述权利要求中任一项所述的方法，其中：
    步骤3)的所述第二分离为减压蒸馏，所述减压蒸馏在约1-20mmHg的真空度和约250-350℃的温度下进行；或者
    步骤3)的所述第二分离为溶剂萃取，其在如下条件下进行：压力为约3-12MPa，优选为约3.5-10Mpa；温度为约55-300℃，优选为约70-220℃；萃取溶剂为C 3-C 7烃，优选为C 3-C 5烷烃和C 3-C 5烯烃中的至少一种，进一步优选为C 3-C 4烷烃和C 3-C 4烯烃中的至少一种；萃取溶剂与所述第一分离产物的重量比为约1∶1至约7∶1，优选为约1.5∶1至约5∶1；或者
    步骤3)的所述第二分离为减压蒸馏与萃取分离的组合，所述减压蒸馏及萃取分离的条件如前文所限定。
  16. 根据前述权利要求中任一项所述的方法，其中在步骤7)中，将步骤3)所得的残渣的约30-95重量％、优选约50-90重量％返回步骤1)中进行所述热转化反应，
    优选地，所述残渣具有小于约150℃的软化点。
  17. 根据前述权利要求中任一项所述的方法，其中步骤5)的所述第三分离包括将所述加氢改质油切割分离为加氢改质轻油和加氢改质重油，所述加氢改质轻油与加氢改质重油之间的切割点为约340-360℃，优选为约345-355℃，更优选为约350℃。
  18. 一种由劣质油生产低碳烯烃的系统，包括热转化反应单元、第一分离单元、第二分离单元、加氢改质单元、第三分离单元和催化裂解单元，其中：
    所述热转化反应单元设置为使劣质油原料在氢气存在下在其中进行热转化反应，得到转化产物；
    所述第一分离单元设置为使所述转化产物在其中分离得到第一分离产物，其中所述第一分离产物中沸点在350℃以下的组分的含量为不大于约5重量％，优选小于约3重量％，沸点在350-524℃之间的组分的含量为约20-60重量％，优选约25-55重量％；
    所述第二分离单元设置为使所述第一分离产物在其中分离得到改质油和残渣，所述第二分离单元选自减压蒸馏单元、溶剂萃取单元或它们的组合；
    所述加氢改质单元设置为使所述改质油在其中进行加氢改质反应，得到加氢改质油；
    所述第三分离单元设置为使所述加氢改质油在其中分离得到加氢改质重油；以及
    所述催化裂解单元设置为使所述加氢改质重油在其中进行催化裂解反应，得到包含低碳烯烃的催化裂解产物。
  19. 根据权利要求18所述的系统，其中所述热转化反应单元包括浆态床反应器。
  20. 根据权利要求18或19所述的系统，其中所述催化裂解单元包括变径稀相输送床反应器和/或由提升管反应器与流化床反应器组成的组合催化裂解反应器。
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Families Citing this family (1)

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Publication number Priority date Publication date Assignee Title
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Citations (9)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US3702886A (en) 1969-10-10 1972-11-14 Mobil Oil Corp Crystalline zeolite zsm-5 and method of preparing the same
US5232675A (en) 1991-02-28 1993-08-03 Research Institute Of Petroleum Processing Rare earth-containing high-silica zeolite having penta-sil type structure and process for the same
CN1233644A (zh) * 1998-04-24 1999-11-03 中国石油化工总公司 缓和热转化－溶剂脱沥青组合工艺
CN101045884A (zh) 2006-03-31 2007-10-03 中国石油化工股份有限公司 一种由渣油和重馏分油生产清洁柴油和低碳烯烃的方法
CN102618323A (zh) * 2011-01-27 2012-08-01 中国石油化工股份有限公司 一种提高液体产物收率的重油原料加工方法
US20130112593A1 (en) * 2004-12-22 2013-05-09 Eni S.P.A. Process for the conversion of heavy charge stocks such as heavy crude oils and distillation residues
WO2015084779A1 (en) 2013-12-02 2015-06-11 Saudi Arabian Oil Company Integrated solvent-deasphalting and fluid catalytic cracking process for light olefin production
CN106701185A (zh) 2015-11-12 2017-05-24 中国石油化工股份有限公司 一种渣油处理方法
CN109486519A (zh) * 2017-09-11 2019-03-19 中国石油化工股份有限公司 一种由低品质油生产高辛烷值汽油的改质方法和系统

Family Cites Families (19)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US5013427A (en) 1989-07-18 1991-05-07 Amoco Corportion Resid hydrotreating with resins
CN1098337C (zh) * 2000-11-02 2003-01-08 中国石油天然气股份有限公司 一种采用多金属液体催化剂的常压重油悬浮床加氢新工艺
CN1171978C (zh) 2001-06-29 2004-10-20 中国石油化工股份有限公司 一种高硫高金属渣油转化方法
US20030127358A1 (en) 2002-01-10 2003-07-10 Letzsch Warren S. Deep catalytic cracking process
CN101987972B (zh) 2009-08-06 2013-06-05 中国石油化工股份有限公司 一种组合工艺加工劣质原油的方法
CN102021031B (zh) 2009-09-17 2013-11-06 中国石油化工股份有限公司 一种从劣质原料油制取优质燃料油的方法
FR2958658B1 (fr) 2010-04-13 2012-03-30 Inst Francais Du Petrole Procede d'hydroconversion de charges petrolieres via une technologie en slurry permettant la recuperation des metaux du catalyseur et de la charge mettant en oeuvre une etape de lixiviation.
CN102071054B (zh) * 2009-10-30 2013-07-31 中国石油化工股份有限公司 一种催化裂化方法
FR2964386B1 (fr) 2010-09-07 2013-09-13 IFP Energies Nouvelles Procede de conversion de residu integrant une etape de desashphaltage et une etape d'hydroconversion
US9101853B2 (en) 2011-03-23 2015-08-11 Saudi Arabian Oil Company Integrated hydrocracking and fluidized catalytic cracking system and process
CN103789036B (zh) 2012-10-26 2015-09-23 中国石油化工股份有限公司 一种劣质重油组合加工方法
FR2999599B1 (fr) 2012-12-18 2015-11-13 IFP Energies Nouvelles Procede de conversion d'une charge hydrocarbonee lourde integrant un desaphaltage selectif avec recyclage de l'huile desasphaltee
FR3014110B1 (fr) * 2013-12-03 2015-12-18 Ifp Energies Now Procede de conversion d'une charge hydrocarbonee lourde integrant un desasphaltage selectif en cascade avec recyclage d'une coupe desasphaltee
FR3019554B1 (fr) * 2014-04-07 2017-10-27 Ifp Energies Now Procede de production d'olefines legeres et de btx utilisant une unite fcc traitant une charge lourde de type vgo tres hydrotraite, couplee avec une unite de reformage catalytique et un complexe aromatique traitant une charge de type naphta.
CN104017601B (zh) * 2014-06-19 2016-11-09 北京中科诚毅科技发展有限公司 一种渣油浆态床加氢裂化方法及装置
CN106635138B (zh) 2015-10-29 2018-05-18 中国石油化工股份有限公司 一种浆态床加氢裂化方法
US10301556B2 (en) * 2016-08-24 2019-05-28 Saudi Arabian Oil Company Systems and methods for the conversion of feedstock hydrocarbons to petrochemical products
CN109385306B (zh) 2017-08-09 2021-05-14 中国石油化工股份有限公司 与加氢处理组合的催化裂化方法及装置
US11078434B2 (en) 2017-09-11 2021-08-03 China Petroleum & Chemical Corporation Process and system for upgrading low-quality oils

Patent Citations (9)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US3702886A (en) 1969-10-10 1972-11-14 Mobil Oil Corp Crystalline zeolite zsm-5 and method of preparing the same
US5232675A (en) 1991-02-28 1993-08-03 Research Institute Of Petroleum Processing Rare earth-containing high-silica zeolite having penta-sil type structure and process for the same
CN1233644A (zh) * 1998-04-24 1999-11-03 中国石油化工总公司 缓和热转化－溶剂脱沥青组合工艺
US20130112593A1 (en) * 2004-12-22 2013-05-09 Eni S.P.A. Process for the conversion of heavy charge stocks such as heavy crude oils and distillation residues
CN101045884A (zh) 2006-03-31 2007-10-03 中国石油化工股份有限公司 一种由渣油和重馏分油生产清洁柴油和低碳烯烃的方法
CN102618323A (zh) * 2011-01-27 2012-08-01 中国石油化工股份有限公司 一种提高液体产物收率的重油原料加工方法
WO2015084779A1 (en) 2013-12-02 2015-06-11 Saudi Arabian Oil Company Integrated solvent-deasphalting and fluid catalytic cracking process for light olefin production
CN106701185A (zh) 2015-11-12 2017-05-24 中国石油化工股份有限公司 一种渣油处理方法
CN109486519A (zh) * 2017-09-11 2019-03-19 中国石油化工股份有限公司 一种由低品质油生产高辛烷值汽油的改质方法和系统

Non-Patent Citations (1)

* Cited by examiner, † Cited by third party
Title
See also references of EP3936589A4

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EP3936589A4 (en) 2022-03-30
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