WO2021083302A1 - 一种加工富芳馏分油的方法和*** - Google Patents

一种加工富芳馏分油的方法和*** Download PDF

Info

Publication number
WO2021083302A1
WO2021083302A1 PCT/CN2020/125068 CN2020125068W WO2021083302A1 WO 2021083302 A1 WO2021083302 A1 WO 2021083302A1 CN 2020125068 W CN2020125068 W CN 2020125068W WO 2021083302 A1 WO2021083302 A1 WO 2021083302A1
Authority
WO
WIPO (PCT)
Prior art keywords
unit
reaction
oil
reaction unit
hydrogenation
Prior art date
Application number
PCT/CN2020/125068
Other languages
English (en)
French (fr)
Inventor
杨清河
贾燕子
胡大为
牛传峰
孙淑玲
戴立顺
王振
户安鹏
任亮
李大东
Original Assignee
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
Priority date (The priority date is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the date listed.)
Filing date
Publication date
Priority claimed from CN201911054674.9A external-priority patent/CN112745952B/zh
Priority claimed from CN201911053864.9A external-priority patent/CN112745949B/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 KR1020227017073A priority Critical patent/KR20220091510A/ko
Priority to US17/772,317 priority patent/US20220403263A1/en
Priority to JP2022525049A priority patent/JP2023501181A/ja
Publication of WO2021083302A1 publication Critical patent/WO2021083302A1/zh

Links

Images

Classifications

    • 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
    • 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
    • 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
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/02—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils characterised by the catalyst used
    • 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
    • C10G67/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only
    • C10G67/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only plural serial stages only
    • 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
    • C10G2400/02—Gasoline
    • 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
    • C10G2400/04—Diesel oil
    • 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
    • C10G2400/08—Jet fuel
    • 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
    • C10G2400/30—Aromatics

Definitions

  • the invention relates to the field of hydrocarbon oil processing, in particular to a method for processing aromatic-rich distillate oil and a system for processing aromatic-rich distillate oil.
  • Efficient conversion of residual oil is the core of oil refining enterprises.
  • the fixed-bed residual oil hydrogenation is a key technology for high-efficiency conversion of residual oil, which has the characteristics of good product quality and mature technology.
  • the residual solvent deasphalting (demetal)-hydrotreating-catalytic cracking combined process technology (SHF) developed by the Sinopec Research Institute of Petroleum and Chemical Technology is to maximize the production of automotive use from low-value vacuum residues.
  • SHF residual solvent deasphalting
  • DOA deoiled asphaltene
  • the new combined process of residue hydrogenation-catalytic cracking (DCC) to produce more propylene in the transition to chemical industry is also limited by the influence of asphaltenes and metals in the residue.
  • the hydrogen content of the hydrogenation residue is low, and the residue is hydrogenated.
  • the operation cycle is short and the DCC propylene yield is low, which affects the economic benefits of the combined technology.
  • the purpose of the present invention is to provide a new method for processing aromatic-rich distillate oil, which can be carried out even at a relatively low hydrogen partial pressure and a relatively low hydrogen-to-oil ratio and at a relatively high space velocity. Obtain better hydrotreating effect and long-term stable operation of the device.
  • the first aspect of the present invention provides a method for processing aromatic-rich distillate oil, the method comprising:
  • the aromatic-rich distillate is introduced into the third reaction unit for hydrogenation saturation and then fractionated to obtain the first light component and the first heavy component, and the cutting of the first light component and the first heavy component
  • the point is 100-250°C, and the aromatic hydrocarbon content in the first heavy component is greater than or equal to 20% by mass;
  • the first reaction unit contains a rich ore precursor material and/or a hydrogenation catalyst, the first reaction unit is a liquid phase hydrogenation reaction unit, and the rich ore precursor material is capable of adsorbing V, Ni, Fe, Ca And at least one metal in Mg, the amount ratio of the deoiled pitch and the aromatic hydrocarbon-containing stream is such that the mixed raw material formed by the deoiled pitch and the aromatic hydrocarbon-containing stream is liquid at not higher than 400°C;
  • the second aspect of the present invention provides a system for processing aromatic-rich distillate oil, which includes:
  • the third reaction unit which is used to hydrogenate and fractionate the aromatic-rich distillate oil therein to obtain the first light component and the first heavy component;
  • a hydrogen dissolving unit which is kept in fluid communication with the third reaction unit, and is used for mixing deoiled pitch and an aromatic hydrocarbon-containing stream containing the first heavy component from the third reaction unit with hydrogen;
  • a first reaction unit which is a liquid phase hydrogenation reaction unit and is kept in fluid communication with the hydrogen dissolving unit, and is used for hydrogenating the mixture of the hydrogen dissolving unit therein;
  • a separation unit which is kept in fluid communication with the first reaction unit, and is used for fractionating the liquid phase product from the first reaction unit;
  • the second reaction unit the second reaction unit is kept in fluid communication with the separation unit, and is used to react the second light component obtained in the separation unit therein, and the second reaction unit is selected from hydrocracking At least one of a unit, a catalytic cracking unit, and a diesel hydro-upgrading unit;
  • a delayed coking unit which is kept in fluid communication with the separation unit, and is used for reacting the second heavy component obtained in the separation unit in it to obtain a coking gasoline, coking diesel, coking gas oil, and low-carbon coking oil. At least one product of sulfur petroleum coke;
  • An outlet which is kept in fluid communication with the separation unit, and is used to draw the second heavy component obtained from the separation unit as a low-sulfur marine fuel oil component out of the system.
  • the method for processing aromatic-rich distillates provided by the present invention treats residual oil, even if it is carried out at a lower hydrogen partial pressure and a lower hydrogen-to-oil ratio and at a higher space velocity, a better refueling can be obtained. Hydrogen treatment effect and long-term stable operation of the device.
  • the invention is particularly suitable for the hydrogenation conversion of normal slag and reduced slag, and is especially suitable for the hydrogenation conversion of inferior residues with high metal, high carbon residue, high fused ring substances and high nitrogen content.
  • the process method of the present invention for hydrotreating deoiled asphalt (DOA) enables the efficient conversion of heavy oil and can produce gasoline and BTX raw materials, as well as a system and method that can flexibly produce low-sulfur ship fuel and low-sulfur petroleum coke.
  • Fig. 1 is a process flow diagram of processing aromatic-rich distillate in a preferred embodiment of the present invention.
  • Fig. 2 is a process flow diagram of processing aromatic-rich distillate oil according to a specific embodiment of the first variant of the present invention.
  • the third response unit 22 The first component
  • the first aspect of the present invention provides a method for processing aromatic-rich distillates, the method comprising:
  • the aromatic-rich distillate is introduced into the third reaction unit for hydrogenation saturation and then fractionated to obtain the first light component and the first heavy component, and the cutting of the first light component and the first heavy component
  • the point is 100-250°C, and the aromatic hydrocarbon content in the first heavy component is greater than or equal to 20% by mass;
  • the first reaction unit contains a rich ore precursor material and/or a hydrogenation catalyst, the first reaction unit is a liquid phase hydrogenation reaction unit, and the rich ore precursor material is capable of adsorbing V, Ni, Fe, Ca And at least one metal in Mg, the amount ratio of the deoiled pitch and the aromatic hydrocarbon-containing stream is such that the mixed raw material formed by the deoiled pitch and the aromatic hydrocarbon-containing stream is liquid at not higher than 400°C;
  • the second light component is introduced into the second reaction unit for reaction to obtain at least one product selected from the group consisting of gasoline components, diesel components and BTX raw material components, wherein, the second reaction
  • the unit is selected from at least one of a hydrocracking unit, a catalytic cracking unit, and a diesel hydro-upgrading unit;
  • the amount ratio of the deoiled bitumen and the aromatic hydrocarbon-containing stream is such that the mixed raw material formed by the deoiled bitumen and the aromatic hydrocarbon-containing stream is liquid at no higher than 280°C; further preferably, the deoiled bitumen and the aromatic hydrocarbon stream are liquid.
  • the amount ratio of the aromatic hydrocarbon-containing stream is such that the mixed raw material formed by the deoiled asphalt and the aromatic hydrocarbon-containing stream is liquid at not higher than 100°C.
  • the hydrogenation saturation reaction performed in the third reaction unit is partial hydrogenation saturation, and it is particularly preferable that the cutting point of the first light component and the first heavy component is 180°C.
  • the operating conditions in the hydrogen dissolving unit of the present invention include: the volume ratio of the amount of hydrogen fed to the mixed raw material formed by the deoiled asphalt and the aromatic hydrocarbon stream (that is, the volume ratio of hydrogen to oil) is 30 -200, more preferably 50-150, the operating temperature of the hydrogen dissolving unit is 300-450°C, and the pressure is 2-20 MPa.
  • the mixed material obtained after mixing with hydrogen in the hydrogen dissolving unit can enter the first reaction unit in an upward flow manner, or enter the first reaction unit in a downward flow manner.
  • the mixed material obtained after mixing with hydrogen in the hydrogen dissolving unit enters the first reaction unit in an upward flow manner.
  • the hydrogen dissolved and dispersed in the oil basically does not aggregate to form large bubbles. Escape, which can provide sufficient hydrogen source for the hydrogenation reaction, obtain better hydroprocessing effect, and further reduce the tendency of catalyst coking, keep the catalyst high catalytic activity, and further extend the service life of the catalyst and the stability of the device Operation cycle.
  • the first light component preferably enters a catalytic cracking unit to produce light olefins.
  • the present invention does not specifically limit the specific operating conditions for the first light component to enter the catalytic cracking unit to produce low-carbon olefins.
  • the cutting point of the second light component and the second heavy component is 350°C.
  • the ratio of the amount of the deoiled bitumen and the aromatic hydrocarbon-containing stream is such that the 100°C viscosity of the mixed raw material formed by the deoiled bitumen and the aromatic hydrocarbon stream is not greater than 400 mm 2 /s. It is preferably not more than 200 mm 2 /s, and more preferably not more than 100 mm 2 /s.
  • the aromatic hydrocarbon-containing stream also contains aromatic compounds and/or aromatic oil, and the aromatic oil is selected from the group consisting of LCO, HCO, FGO (catalytic heavy distillate oil), ethylene tar, coal At least one of tar, coker diesel, and coker wax oil.
  • aromatic oil is selected from the group consisting of LCO, HCO, FGO (catalytic heavy distillate oil), ethylene tar, coal At least one of tar, coker diesel, and coker wax oil.
  • the aromatic hydrocarbon compound is selected from one or more of benzene, toluene, xylene, naphthalene, methyl naphthalene, multi-branched naphthalene and aromatic hydrocarbons above bicyclic rings, preferably polycyclic aromatic hydrocarbons with ring numbers not exceeding three rings Or a mixture of them.
  • the aromatic hydrocarbon compound is selected from at least one of benzene, toluene, xylene, naphthalene, naphthalene substituted with at least one C 1-6 alkyl group, and aromatic hydrocarbon with three or more rings.
  • the aromatic content in the aromatic-rich distillate oil is greater than or equal to 20% by mass, preferably greater than or equal to 25% by mass, preferably greater than or equal to 40% by mass, and more preferably greater than or equal to 60% by mass.
  • the deoiled asphalt is the deoiled asphalt obtained after the heavy oil raw material enters the solvent deasphalting unit for solvent deasphalting treatment.
  • the mass fraction of the yield of the deoiled asphalt is not more than 50%, more preferably not more than 40%, and further preferably not more than 30%.
  • the amount-to-mass ratio of the deoiled asphalt to the aromatic hydrocarbon-containing stream is 1:10-50:10, more preferably 2:10-30:10 ; More preferably, 3:10-15:10.
  • the method of the present invention further comprises: recycling the coking diesel oil and/or the coking wax oil obtained in step (42) back to the first reaction unit in step (1) for hydrogenation saturation.
  • the third reaction unit is at least one of a fixed bed reactor, a moving bed reactor, and a fluidized bed reactor.
  • the operating conditions in the third reaction unit include: a reaction temperature of 200-420°C, a reaction pressure of 2-18 MPa, a liquid hourly volumetric space velocity of 0.3-10h -1 , and a hydrogen-to-oil volume ratio of 50-5000 More preferably, the operating conditions in the third reaction unit include: a reaction temperature of 220-400° C., a reaction pressure of 2-15 MPa, a liquid hourly volumetric space velocity of 0.3-5 h -1 , and a hydrogen-to-oil volume ratio of 50 -4000.
  • the conditions for partial hydrogenation saturation of aromatic-rich distillates with hydrogen are generally as follows:
  • the partial hydrogenation saturation technology of aromatic-rich distillates is a fixed bed/ebullating bed/moving bed hydroprocessing technology.
  • the reactor or reaction bed layer includes at least one hydrorefining catalyst.
  • the hydrorefining catalyst used in the partial hydrogenation saturation of aromatic-rich distillates preferably has good and moderate hydrogenation saturation activity to avoid further saturation of the tetralin structure into decalin or naphthenic structure with lower hydrogen supply capacity .
  • These catalysts are generally based on porous refractory inorganic oxides such as alumina as the support, and the oxides of Group VIB and/or Group VIII metals such as W, Mo, Co, Ni, etc. are used as active components, and other components are selectively added.
  • a variety of additives such as P, Si, F, B and other elements of the catalyst, for example, the RS series pretreatment catalyst developed by the Research Institute of Petrochemical Industry belongs to this type of catalyst.
  • RS series catalyst is a kind of NiMo catalyst.
  • the first reaction unit is a residue liquid phase hydrogenation reactor.
  • the operating conditions in the first reaction unit include: a reaction temperature of 260 to 500° C., a reaction pressure of 2.0 to 20.0 MPa, circulating oil and feed oil at the inlet of the first reaction unit
  • the volume ratio is 0.1:1 to 15:1
  • the liquid hourly volumetric space velocity is 0.1 to 1.5h -1
  • the liquid hourly volumetric space velocity is 0.1 to 1.5h -1 .
  • Liquid hourly volumetric space velocity and reaction pressure can be selected according to the characteristics of the material to be treated, the required conversion rate and the refining depth.
  • the mixed raw material formed by the deoiled asphalt and aromatic hydrocarbon stream of the present invention can enter from the top of the reactor of the first reaction unit after being mixed with hydrogen, and pass through the catalyst bed from top to bottom; or from the first reaction unit It enters from the bottom of the reactor and passes through the catalyst bed from bottom to top.
  • the rich ore precursor material contains a carrier and an active component element supported on the carrier, and the carrier is selected from at least one of aluminum hydroxide, aluminum oxide and silicon oxide.
  • the active component element is selected from at least one of group VIB and group VIII metal elements. More preferably, the active components in the rich ore precursor material are oxides and/or sulfides selected from the group VIB and VIII metal elements.
  • the ignition loss of the rich ore precursor material is not less than 3% by mass, the specific surface area is not less than 80 m 2 /g, and the water absorption rate is not less than 0.9 g/g.
  • the ignition reduction refers to the percentage of the mass of the rich ore precursor material after roasting treatment at 600°C/2h, which accounts for the percentage of the mass before roasting;
  • the water absorption refers to the immersion of the rich ore precursor material in water for half an hour at room temperature (for example, 25°C) The added mass accounts for the percentage of the mass before soaking.
  • step (2) according to the direction of the reactant flow, the first reaction unit is sequentially filled with a first rich ore precursor material and a second rich ore precursor material, and the second The ignition loss of the rich ore precursor material is greater than or equal to the ignition loss of the first rich ore precursor material.
  • the ignition loss of the first rich ore precursor material is 3-15% by mass, and the ignition loss of the second rich ore precursor material is not less than 15% by mass.
  • the filling volume ratio of the first rich ore precursor material to the second rich ore precursor material is 5:95 to 95:5.
  • the hydrogenation catalyst of the present invention may be a graded combination of different catalysts.
  • the hydrogenation catalyst can at least catalyze the hydrodemetalization reaction and the hydrodesulfurization reaction.
  • the present invention does not specifically limit the specific types of catalysts that can catalyze the hydrodemetalization reaction, the hydrodesulfurization reaction, the hydrodeasphalting reaction, and the hydrodecarbonization reaction. Conventionally used in the field can be used to catalyze the above reaction. Catalyst.
  • the hydrogenation catalyst of the present invention may be, for example, a porous refractory inorganic oxide as a support, a group VIB and/or group VIII metal oxide or sulfide as an active component, and a catalyst optionally added with an auxiliary agent.
  • the rich ore precursor material can be transformed into a vanadium-rich material, and the vanadium content in the vanadium-rich material is not less than 10% by mass; particularly preferably, the rich ore precursor
  • the bulk material is transformed into a vanadium-rich material with a V content of more than 20% by mass, which can directly refine high-value V 2 O 5 .
  • the raw material hydrotreating technology involved in the first reaction unit of the present invention is a liquid-phase hydrotreating technology, and the reactor or reaction bed layer at least includes a rich ore precursor material and/or a hydrogenation catalyst,
  • the rich ore precursor material is mainly composed of two parts: one is the carrier with strong ability to adsorb vanadium-containing organic compounds in the oil, and the other is the active component with the function of hydrogenation activity.
  • the carrier is mainly obtained by extruding and drying silicon oxide, aluminum hydroxide or aluminum hydroxide/alumina mixture.
  • the surface is rich in -OH and has strong adsorption capacity for vanadium-containing organic compounds in the oil. It is calcined at 600°C. 2h, the ignition loss is not less than 5% by mass.
  • the active components are mainly oxides or sulfides of Group VIB and/or Group VIII metals such as W, Mo, Co, Ni, etc.
  • the hydrogenation catalyst involved in the foregoing preferred embodiments is generally a heavy residue hydrogenation catalyst.
  • the heavy residue hydrogenation catalyst refers to the functions of heavy and residual oil hydrodemetalization, hydrodesulfurization, and hydrodecarbonization.
  • the combination of catalysts. These catalysts are generally based on porous refractory inorganic oxides such as alumina as the carrier, and the oxides or sulfides of Group VIB and/or Group VIII metals such as W, Mo, Co, Ni, etc. as the active components, selectively Add other various additives such as P, Si, F, B and other elements of the catalyst, such as RDM, RCS series of heavy and residual oil hydrodemetalization catalysts and desulfurization catalysts developed by the Research Institute of Petrochemical Sciences.
  • the liquid phase hydroprocessing technology multiple catalysts are often used in conjunction.
  • the filling sequence is generally such that the raw materials are sequentially contacted with the ore-rich precursor materials, hydrodesulfurization, and hydrodesulfurization catalysts.
  • one or two catalysts should be installed less, for example, only the rich ore precursor material and the hydrodesulfurization catalyst are installed, and the hydrodemetalization desulfurization catalyst is not installed.
  • the second reaction unit is a hydrocracking unit
  • the operating conditions in the hydrocracking unit include: the reaction temperature is 360-420°C, and the reaction The pressure is 10.0 ⁇ 18.0MPa, the volume ratio of hydrogen to oil is 600 ⁇ 2000, and the liquid hourly volumetric space velocity is 1.0 ⁇ 3.0h -1 .
  • the hydrocracking unit is filled with at least one hydrotreating catalyst and at least one hydrocracking catalyst.
  • the hydrocracking unit is a fixed bed hydrocracking unit.
  • the second reaction unit is a hydrocracking unit
  • the following provides a preferred specific embodiment of the second reaction unit of the present invention:
  • the second light component is introduced into the second reaction unit for reaction, and the hydrocracking technology used is a fixed bed hydrocracking technology.
  • the reactor or reaction bed layer includes at least two hydrocracking catalysts, one is a pretreatment catalyst and the other is a hydrocracking catalyst. Since the metal content, sulfur, nitrogen content and carbon residue value of the materials obtained by liquid phase hydrotreating technology and fractional distillation are high, the pretreatment catalyst preferably has strong demetallization activity and good desulfurization and desulfurization. Nitrogen activity to ensure the activity of the subsequent hydrocracking catalyst.
  • the hydrocracking catalyst preferably has good hydrocracking activity and high VGO conversion and HDS activity.
  • These catalysts are generally porous refractory inorganic oxides such as alumina or molecular sieves as the carrier, and the oxides of Group VIB and/or Group VIII metals such as W, Mo, Co, Ni, etc. are used as active components, which are selectively added
  • Various other additives such as P, Si, F, B and other elemental catalysts, such as the RS series pretreatment catalysts and RHC series hydrocracking catalysts developed by the Research Institute of Petrochemical Industry, belong to this category of catalysts.
  • the RS series catalyst is a NiW catalyst
  • the RHC series catalyst is a NiMo molecular sieve catalyst.
  • the second reaction unit is a catalytic cracking unit
  • the catalytic cracking unit is a fluidized catalytic cracking (FCC) unit.
  • the second light component catalytic cracking technology used in the catalytic cracking unit is fluidized-bed catalytic cracking (FCC) technology, preferably the LTAG technology developed by the Research Institute of Petrochemical Technology, which mainly produces gasoline Distillate and liquefied petroleum gas.
  • FCC fluidized-bed catalytic cracking
  • the operating conditions in the fluidized catalytic cracking unit include: a reaction temperature of 500 to 600° C., a catalyst-to-oil ratio of 3 to 12, and a residence time of 0.6 to 6 s.
  • agent-to-oil ratio in the present invention all means the agent-to-oil mass ratio.
  • the second reaction unit is a diesel hydro-upgrading unit
  • the operating conditions in the diesel hydro-upgrading unit include: a reaction temperature of 330 ⁇ 420°C, reaction pressure is 5.0 ⁇ 18.0MPa, hydrogen-oil volume ratio is 500 ⁇ 2000, liquid hour volume space velocity is 0.3 ⁇ 3.0h -1 .
  • the diesel hydro-upgrading unit is filled with at least one diesel hydro-upgrading catalyst.
  • the diesel hydro-upgrading catalyst of the present invention can be, for example, the RS series pretreatment catalyst and the RHC-100 series diesel hydrocracking catalyst developed by the Research Institute of Petrochemical Industry.
  • the second heavy component is introduced into the delayed coking unit for reaction to obtain a product selected from the group consisting of coking gasoline, coking diesel, coking wax oil, and low-sulfur petroleum coke.
  • a product selected from the group consisting of coking gasoline, coking diesel, coking wax oil, and low-sulfur petroleum coke.
  • At least one product of, and the operating conditions in the delayed coking unit include: a reaction temperature of 440-520°C, and a residence time of 0.1-4h.
  • the sulfur content of the second heavy component is not more than 1.8% by mass, and the second heavy component is introduced into the delayed coking unit for reaction to obtain low Sulfur petroleum coke, more preferably, the sulfur content of the low-sulfur petroleum coke is not more than 3% by mass.
  • the second heavy component is used as the low-sulfur marine fuel oil component, and the conditions are controlled such that the sulfur content in the low-sulfur marine fuel oil component is not more than 0.5% by mass.
  • the present invention does not particularly limit the specific operation of the solvent deasphalting treatment, and it can be carried out by using a conventional solvent deasphalting process in the field.
  • the operating parameters of the solvent deasphalting process are exemplarily listed in the examples of the present invention, and those skilled in the art should not be understood as limiting the present invention.
  • the invention is suitable for the hydrogenation conversion of normal slag and reduced slag, and is especially suitable for high metal (Ni+V>150 ⁇ g/g, especially Ni+V>200 ⁇ g/g), high carbon residue (mass fraction of carbon residue>17%, In particular, the mass fraction of carbon residue>20%), the inferior residue of high-density ring substances is hydroconverted.
  • the second aspect of the present invention provides a system for processing aromatic-rich distillates, which includes:
  • the third reaction unit which is used to hydrogenate and fractionate the aromatic-rich distillate oil therein to obtain the first light component and the first heavy component;
  • a hydrogen dissolving unit which is kept in fluid communication with the third reaction unit, and is used for mixing deoiled pitch and an aromatic hydrocarbon-containing stream containing the first heavy component from the third reaction unit with hydrogen;
  • a first reaction unit which is a liquid phase hydrogenation reaction unit and is kept in fluid communication with the hydrogen dissolving unit, and is used for hydrogenating the mixture of the hydrogen dissolving unit therein;
  • a separation unit which is kept in fluid communication with the first reaction unit, and is used for fractionating the liquid phase product from the first reaction unit;
  • the second reaction unit the second reaction unit is kept in fluid communication with the separation unit, and is used to react the second light component obtained in the separation unit therein, and the second reaction unit is selected from hydrocracking At least one of a unit, a catalytic cracking unit, and a diesel hydro-upgrading unit;
  • a delayed coking unit which is kept in fluid communication with the separation unit, and is used for reacting the second heavy component obtained in the separation unit in it to obtain a coking gasoline, coking diesel, coking gas oil, and low-carbon coking oil. At least one product of sulfur petroleum coke;
  • An outlet which is kept in fluid communication with the separation unit, and is used to draw the second heavy component obtained from the separation unit as a low-sulfur marine fuel oil component out of the system.
  • the delayed coking unit is kept in fluid communication with the hydrogen dissolving unit for recycling the coking diesel oil and/or the coking wax oil obtained in the delayed coking unit to the first reaction unit .
  • the system further includes a solvent deasphalting unit, which is kept in fluid communication with the hydrogen dissolving unit, and is used for solvent deasphalting the heavy oil feedstock therein, and deasphalting the solvent.
  • the deoiled asphalt obtained later is introduced into the hydrogen dissolving unit.
  • the second reaction unit is a hydrocracking unit.
  • the second reaction unit is a catalytic cracking unit
  • the catalytic cracking unit is a fluidized catalytic cracking unit
  • the second reaction unit is a diesel hydro-upgrading unit.
  • the present invention also provides a first variant of the method, in which the first variant further includes:
  • the deasphalted oil is introduced into the reaction unit of the fourth hydrogenation unit for hydrogenation reaction, and the liquid phase effluent obtained in the reaction unit of the fourth hydrogenation unit is introduced into the DCC unit for reaction, to obtain Propylene, LCO, HCO and oil slurry, wherein the fourth hydrogenation unit reaction unit is a fixed bed hydrogenation unit reaction unit;
  • the aromatic-rich distillate containing LCO and/or HCO from the DCC unit is used as the aromatic-rich distillate in the step (1).
  • the method of the present invention further comprises: recycling the coking diesel oil and/or the coking wax oil obtained in step (42) back to the third reaction unit for adding Saturated with hydrogen.
  • the operating conditions of the fourth reaction unit include: a reaction temperature of 280 to 400° C., a reaction pressure of 6.0 to 14.0 MPa, a hydrogen-to-oil volume ratio of 600 to 1200, and liquid hour volume
  • the airspeed is 0.3 ⁇ 2.0h -1 .
  • the fourth reaction unit is filled with at least two hydrogenation catalysts.
  • the hydrogenation catalyst is a catalyst capable of catalyzing at least one reaction selected from the group consisting of a hydrodemetalization reaction, a hydrodesulfurization reaction, and a hydrodecarbonization reaction.
  • the hydrogenation catalyst generally uses porous refractory inorganic oxides such as alumina as a carrier; particularly preferably, in step (12), the hydrogenation catalyst contains alumina as a carrier and as an active component. It is a group VIB and/or group VIII metal element of the element, and the hydrogenation catalyst optionally further contains at least one auxiliary element selected from the group consisting of P, Si, F and B.
  • the group VIB and group VIII metal elements may be, for example, W, Mo, Co, Ni, and the like.
  • the active component may be an oxide and/or sulfide of the above-mentioned active component element.
  • the conditions of the fourth reaction unit of deasphalted oil (DAO) with hydrogen are generally as follows:
  • the hydroprocessing technology of DAO is a fixed-bed hydroprocessing technology.
  • the reactor or reaction bed layer includes at least two hydrogenation catalysts.
  • the heavy residual oil hydrogenation catalyst used means the A combined catalyst with functions such as hydrodemetalization, hydrodesulfurization, hydrodenitrogenation, and hydrodecarbonization.
  • catalysts are generally based on porous refractory inorganic oxides such as alumina as supports, and Group VIB and/or Group VIII metals such as oxides or sulfides of W, Mo, Co, Ni, etc., as active components, selectively Add other various additives such as P, Si, F, B and other elements of the catalyst, such as RDM, RCS series of heavy and residual oil hydrodemetalization catalysts and desulfurization catalysts developed by the Research Institute of Petrochemical Sciences.
  • RDM Rasteretalization catalysts
  • hydrodesulfurization catalysts hydrodesulfurization catalysts
  • hydrodenitrogenation catalysts hydrodenitrogenation catalysts.
  • the filling order is generally such that the feedstock oil is sequentially followed by hydrogenation and denitrification.
  • Metal, hydrodesulfurization, and hydrodenitrogenation catalysts are contacted.
  • one or two catalysts can be installed less according to the situation. For example, only the hydrodemetalization catalyst and the hydrodesulfurization catalyst are installed, and the hydrodenitrogenation catalyst is not installed. .
  • the aromatic-rich distillate oil 20 is introduced into the third reaction unit 21 for hydrogenation saturation and then fractionated to obtain the first light component and the first heavy component 22; and the heavy oil feedstock 1 enters the solvent deasphalting unit 2
  • the deoiled asphalt 4 and the deasphalted oil 3 obtained after the solvent deasphalting treatment are carried out in the process; the deoiled asphalt 4 and the aromatic hydrocarbon stream containing the first heavy component 22 together form the mixed raw material 6 and enter the hydrogen dissolving unit 23 with hydrogen
  • the mixture material thus obtained enters the first reaction unit 7 for hydrogenation reaction.
  • the aromatic hydrocarbon-containing stream preferably further contains aromatic hydrocarbon compounds 5 from the outside, wherein the first reaction unit contains ore-rich precursor materials.
  • a hydrogenation catalyst capable of catalyzing at least one reaction selected from the group consisting of a hydrodemetalization reaction, a hydrodesulfurization reaction, a hydrodeasphalting reaction, and a hydrodecarbonization reaction
  • the first reaction unit is a liquid phase hydrogenation Reaction unit
  • the liquid phase product from the first reaction unit 7 enters the separation unit 19 for fractional distillation to obtain a second light component 8 and a second heavy component 9, wherein the second light component and the first
  • the cutting point of the double component is 240-450°C
  • the second light component 8 is introduced into the second reaction unit 10 for reaction to obtain a gasoline component 13, a BTX raw material component 12, and a diesel component 14
  • At least one product of the second reaction unit wherein the second reaction unit is selected from at least one of a hydrocracking unit, a catalytic cracking unit, and a diesel hydro-upgrading unit
  • the second heavy component 9 is introduced to the delay Reaction in the coking unit 11 to obtain at least one product selected from
  • the heavy oil feedstock 1 enters the solvent deasphalting unit 2 for solvent deasphalting treatment to obtain deoiled asphalt 4 and deasphalted oil 3; the deasphalted oil 3 is introduced into the fourth reaction unit 24 for adding Hydrogen reaction, and the liquid phase effluent obtained in the fourth reaction unit 24 is introduced into the DCC unit 25 for reaction to obtain propylene 26, LCO27, HCO28 and oil slurry 29; will contain LCO27 from the DCC unit 25 And/or the aromatic-rich distillate 20 of HCO28 is introduced into the third reaction unit 21 for hydrogenation saturation and fractional distillation to obtain the first heavy component 22 and the first light component; The aromatic hydrocarbon-containing stream divided into 22 together form the mixed raw material 6 and is introduced into the hydrogen dissolving unit 29 to be mixed with hydrogen, and the mixed material is introduced into the first reaction unit 7 for hydrogenation reaction.
  • the aromatic hydrocarbon-containing stream is preferably recycled Contains an aromatic compound 5 from the outside, wherein the first reaction unit 7 contains a rich ore precursor material and can catalyze selected from a hydrodemetalization reaction, a hydrodesulfurization reaction, a hydrodeasphalting reaction, and a hydrodecarbonization reaction
  • the hydrogenation catalyst for at least one reaction in the reaction the liquid phase product from the first reaction unit 7 enters the separation unit 19 for fractional distillation to obtain the second light component 8 and the second heavy component 9;
  • the second light component 8 is introduced into the second reaction unit 10 for reaction to obtain at least one product selected from the group consisting of gasoline component 13, BTX raw material component 12, and diesel component 14, or the second light component 8 is recycled back to the DCC unit 25; and the second heavy component 9 is introduced into the delayed coking unit 11 for reaction to obtain a group selected from the group consisting of coking gasoline 15, coking diesel 16, coking wax oil 17 and low-sulfur petroleum coke 18 Or use the second heavy component 9 as a low-s
  • the technology of the present invention enables the efficient conversion of heavy oil and can produce gasoline and BTX raw materials, as well as a system and method capable of flexibly producing low-sulfur ship fuel and low-sulfur petroleum coke.
  • the present invention uses organic combination of residual oil hydrogenation, hydrocracking or catalytic cracking processes, which not only converts low-value DOA into a low-sulfur ship fuel group that meets environmental protection requirements. Separate and low-sulfur petroleum coke raw materials, and realize the high-efficiency, environmental protection and comprehensive utilization of heavy petroleum resources.
  • the technology provided by the present invention enables efficient conversion of DOA in a residue liquid phase hydrogenation reactor and can produce gasoline fractions, BTX raw materials, and can provide raw materials for the production of low-sulfur marine fuel and low-sulfur coke products.
  • results of Table I-3 and Table II-4 in the following examples are the average of the results obtained by sampling and testing every 25 hours during the continuous operation of the device for 100 hours.
  • the partial hydrogenation saturation experiment of aromatic-rich distillate was carried out on a medium-sized fixed-bed diesel hydrotreating unit with a total reactor volume of 200 mL.
  • the hydrogenation catalyst and materials used for partial hydrogenation saturation of aromatic-rich distillate oil are the RS-2100 series hydrogenation catalysts developed by the Research Institute of Petrochemical Sciences.
  • Fractional distillation is performed on the liquid phase stream obtained by partial hydrogenation saturation to obtain the first light component and the first heavy component with a cutting point of 180° C., and the first heavy component and DOA form a mixed raw material.
  • the hydrogenation reaction of the mixed raw materials was tested on a medium-sized heavy oil liquid phase hydrotreating device, and the total volume of the reactor was 200 mL.
  • the hydrogenation catalyst and materials used in the first reaction unit are RG-30B protective catalyst, rich ore precursor material 1, rich ore precursor material 2, RDM-33B residue demetallization and desulfurization transition catalyst developed by the Research Institute of Petrochemical Sciences, RCS-31 desulfurization catalyst.
  • the order of catalyst loading is the hydrogenation protection catalyst, the rich ore precursor material 1, the rich ore precursor material 2, the hydrodemetalization desulfurization catalyst, and the hydrodesulfurization catalyst.
  • the second reaction unit is a fixed bed hydrocracking unit, and the catalysts used are RS-2100 refined catalyst and RHC-131 hydrocracking catalyst developed by the Research Institute of Petrochemical Sciences.
  • the operating conditions of the fixed-bed hydrocracking unit are: the reaction temperature of the refining section is 370°C, the reaction temperature of the cracking section is 385°C, the reaction pressure is 10MPa, the liquid hourly volumetric space velocity is 2.0h -1 , and the hydrogen-to-oil volume ratio is 1200 :1.
  • Preparation of rich ore precursor material 1 Select 2000g of RPB110 pseudo-boehmite produced by Changling Branch of Sinopec Catalyst Co., Ltd., of which 1000g is treated at 550°C for 2h to obtain about 700g of alumina, and about 700g of alumina and another 1000g of pseudoboehmite are selected.
  • the boehmite is thoroughly mixed, then 40g sesame powder and 20g citric acid are added, and 2200g deionized water is added, kneaded and extruded, and dried at 300°C for 3h to obtain about 1730g carrier.
  • Preparation of rich ore precursor material 2 Select 2000g of RPB110 pseudo-boehmite produced by Changling Branch of Sinopec Catalyst Co., Ltd., add 30g of sesame powder and 30g of citric acid, and add 2400g of deionized water, knead and extrude into After drying at 120°C for 5 hours, about 2040g of carrier is obtained. 2200mL of solution containing Mo and Ni is added for saturated impregnation. The Mo content in the solution is 7.5% by mass of MoO 3 , and the Ni content is 1.7% by mass of NiO. Soak for half an hour. Afterwards, it was treated at 200°C for 3 hours to obtain the rich ore precursor material 2, whose properties are shown in Table I-6.
  • Preparation of rich ore precursor material 3 select 2000g of commercially available silicon oxide, add 30g of sesame powder and 30g of sodium hydroxide, and add 2400g of deionized water, knead and extrude, dry at 120°C for 5h to obtain a carrier, add 2200mL solution containing Mo and Ni is saturated immersed, the Mo content in the solution is 4.5% by weight of MoO 3 , Ni content is 1.0% by weight of NiO, immersed for half an hour, and then treated at 200°C for 3 hours to obtain a rich ore precursor Material 3, the properties are shown in Table I-6.
  • the aromatic-rich distillate used in this example I is LCO, which comes from the Shanghai Petrochemical RLG plant.
  • the LCO hydrogenation operating conditions are: reaction temperature is 290°C, reaction pressure is 4MPa, liquid hourly volumetric space velocity is 1h -1 , hydrogen oil The volume ratio is 800:1.
  • DOA comes from a vacuum residue and is mixed with the first heavy component 1 in a mass ratio of 1:10.
  • the properties of the mixed raw materials are shown in Table I-2.
  • the mixed raw material of DOA and the first heavy component 1 is first in the hydrogen dissolving unit (the volume ratio of the amount of hydrogen fed to the mixed raw material of the deoiled asphalt and the first heavy component 1 is 100, and the operating temperature of the hydrogen dissolving unit is 320°C, the pressure is 10MPa) and hydrogen, the obtained mixture material enters the first reaction unit, the operating conditions of the first reaction unit are: the reaction temperature is 360°C, the reaction pressure is 10MPa, the liquid hourly volumetric space velocity is 0.6h -1 , circulating oil: the volume ratio of the feed oil at the inlet of the first reaction unit is 0.5:1.
  • Table I-3 The properties of the mixed raw materials after hydrogenation are shown in Table I-3.
  • the liquid phase product obtained by the fractionation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C in Table I-4.
  • the second light component below 350°C was tested in the second reaction unit to obtain the hydrocracking product.
  • the properties are shown in Table I-5.
  • the aromatic-rich distillate used in this example I is HCO, which comes from Shanghai Petrochemical’s catalytic cracking unit.
  • the HCO hydrogenation operating conditions are: reaction temperature of 330°C, reaction pressure of 6MPa, liquid hourly volumetric space velocity of 1h -1 , hydrogen
  • the oil volume ratio is 800:1.
  • HCO The properties of HCO and the first heavy component 2 are shown in Table I-1.
  • DOA comes from a vacuum residue and is mixed with the first heavy component 2 in a mass ratio of 5:10.
  • the properties of the mixed raw materials are shown in Table I-2.
  • the mixed raw material of DOA and HCO first heavy component 2 after hydrogenation is first in the hydrogen dissolving unit (the volume ratio of the amount of hydrogen fed to the mixed raw material of the deoiled asphalt and the first heavy component 2 is 100, and the hydrogen is dissolved
  • the unit operating temperature is 320°C, the pressure is 10MPa), and the mixture is mixed with hydrogen, and the obtained mixture enters the first reaction unit.
  • the operating conditions in the first reaction unit are: the reaction temperature is 380°C, the reaction pressure is 10MPa, and the liquid-hour volume
  • the space velocity is 0.6h -1 , and the volume ratio of circulating oil: the feedstock oil at the inlet of the first reaction unit is 0.5:1.
  • Table I-3 The properties of the mixed raw materials after hydrogenation are shown in Table I-3.
  • the liquid phase product obtained by the fractionation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C in Table I-4.
  • the second light component below 350°C was tested in the second reaction unit to obtain the hydrocracking product.
  • the properties are shown in Table I-5.
  • Example I- The aromatic-rich distillate used in Example I- is the same LCO as in Example I-1.
  • the LCO hydrogenation operating conditions are: reaction temperature is 320°C, reaction pressure is 6MPa, and liquid hourly volumetric space velocity is 1h -1 , The volume ratio of hydrogen to oil is 800:1.
  • DOA comes from a vacuum residue, mixed with the first heavy component 3 at a mass ratio of 10:10.
  • the properties of the mixed raw materials are shown in Table I-2.
  • the mixed raw material of DOA and the first heavy component 3 is first in the hydrogen dissolving unit (the volume ratio of the amount of hydrogen fed to the mixed raw material of the deoiled asphalt and the first heavy component 3 is 100, and the operating temperature of the hydrogen dissolving unit is 320°C, the pressure is 8MPa) is mixed with hydrogen, the obtained mixture material enters the first reaction unit, the operating conditions in the first reaction unit are: reaction temperature is 370°C, reaction pressure is 8MPa, liquid hourly volumetric space velocity is 0.6 h -1 , circulating oil: the volume ratio of the feed oil at the inlet of the first reaction unit is 0.5:1
  • Table I-3 The properties of the mixed raw materials after hydrogenation are shown in Table I-3.
  • the liquid phase product obtained by the fractionation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C in Table I-4.
  • the second heavy component was subjected to a coking reaction at a reaction temperature of 500° C. and a residence time of 0.5 hours to obtain petroleum coke (yield 32% by mass) with a sulfur content of 2.7% by mass.
  • the second light component below 350°C was tested in the second reaction unit to obtain the hydrocracking product.
  • the properties are shown in Table I-5.
  • Example I- aromatic-rich coal tar distillates coal from a domestic apparatus, coal tar hydrogenation operating conditions: reaction temperature of 300 °C, the reaction pressure is 10 MPa or, when the liquid hourly space velocity of 0.8h - 1.
  • the volume ratio of hydrogen to oil is 800:1.
  • DOA comes from a vacuum residue, mixed with the first heavy component 4 at a mass ratio of 15:10.
  • the properties of the mixed raw materials are shown in Table I-2.
  • the mixed raw material of DOA and the first heavy component 4 is first in the hydrogen dissolving unit (the volume ratio of the amount of hydrogen fed to the mixed raw material of the deoiled asphalt and the first heavy component 4 is 100, and the operating temperature of the hydrogen dissolving unit is 320°C, the pressure is 12MPa) and hydrogen, the obtained mixture material enters the first reaction unit, the operating conditions in the first reaction unit are: the reaction temperature is 350°C, the reaction pressure is 12MPa, the liquid hourly volumetric space velocity is 0.6 h -1 , circulating oil: the volume ratio of the feedstock oil at the inlet of the first reaction unit is 2:1.
  • Table I-3 The properties of the mixed raw materials after hydrogenation are shown in Table I-3.
  • the liquid phase product obtained by the fractionation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C in Table I-4.
  • the second light component below 350°C was tested in the second reaction unit to obtain the hydrocracking product.
  • the properties are shown in Table I-5.
  • Example I- the temperature of the hydroprocessing of the first reaction unit was 395°C.
  • Example I-1 The operating conditions of the raw material, catalyst filling, and heavy oil liquid phase hydrotreating unit are the same as in Example I-1. The difference is:
  • Example I-1 After the same mixed raw material as in Example I-1 was hydrotreated with liquid phase heavy oil, the reaction temperature was increased by 3°C every 30 days, and the hydrogenation test was stopped after a total of 360 days of operation.
  • the rich ore precursor material 1 and rich ore precursor material 2 initially loaded into the reactor become V-rich material 1 and vanadium-rich material 2 after the reaction.
  • the V content is 76% by mass and 71% by mass, respectively.
  • the content is more than 10 times higher than that of natural ore. It is a high-quality material for refining high-value V 2 O 5.
  • the second light component less than 350°C in Example I-3 was subjected to a catalytic cracking test in a small catalytic cracking fixed fluidized bed test device.
  • the catalyst used was the catalytic cracking catalyst MLC-500 produced by the Changling Branch of Sinopec Catalyst Co., Ltd. ,
  • the reaction temperature is 540°C
  • the agent-to-oil ratio is 6, and the residence time is 2s.
  • the quality yield of the obtained product gasoline was 42%, and the gasoline RON octane number was 92.
  • Example I-1 The process is similar to that of Example I-1, except that the second heavy component obtained in Example I- is introduced into the delayed coking unit for reaction to obtain coking gasoline, coking diesel and coking wax oil.
  • the operating conditions of the delayed coking unit are: the reaction temperature is 510°C, and the residence time is 0.6h.
  • the sulfur content of coker diesel oil is 0.26% by mass, the freezing point is -11°C, and the cetane number is 48.
  • the sulfur content of the coking wax oil is 1.12% by mass, and the freezing point is 32°C.
  • the yield of coking gasoline was 14.7%, the sulfur content was 0.10% by mass, and the MON was 61.8.
  • the coker diesel oil and the coker wax oil are recycled back to the third reaction unit and mixed with the LCO for hydrotreating.
  • the reaction process conditions are the same as those in Example I-1.
  • DOA comes from a vacuum residue and is mixed with the first heavy component 8 in a mass ratio of 1:10.
  • the properties of the mixed raw materials are shown in Table I-2.
  • the mixed raw material of DOA and the first heavy component 8 is first in the hydrogen dissolving unit (the volume ratio of the amount of hydrogen fed to the mixed raw material of the deoiled asphalt and the first heavy component 8 is 100, and the operating temperature of the hydrogen dissolving unit is 320°C, the pressure is 8MPa) is mixed with hydrogen, the obtained mixture material enters the first reaction unit, the operating conditions of the first reaction unit are: the reaction temperature is 360°C, the reaction pressure is 8MPa, the liquid hourly volumetric space velocity is 0.3h -1 , circulating oil: the volume ratio of the feed oil at the inlet of the first reaction unit is 0.5:1.
  • Table I-3 The properties of the mixed raw materials after hydrogenation are shown in Table I-3.
  • the second light component below 350°C was tested in the second reaction unit to obtain the hydrocracking product.
  • the properties are shown in Table I-5.
  • Example I-1 The second light component below 350°C obtained in Example I-1 was tested on a hydrocracking unit to obtain a diesel component.
  • the operating conditions are: the reaction temperature is 360°C, the reaction pressure is 10 MPa, the hydrogen-to-oil volume ratio is 1000, and the liquid hourly volumetric space velocity is 1.0 h -1 .
  • the sulfur content of the diesel component is 5ppm, the freezing point is -32°C, and the cetane number is 53.
  • Example I- The process is similar to that of Example I-1, except that the catalyst filling in the first reaction unit in Example I- is as follows:
  • the order of catalyst loading is the hydrogenation protection catalyst, the rich ore precursor material 1, the hydrodemetalization desulfurization catalyst, and the hydrodesulfurization catalyst.
  • the liquid phase product obtained by the fractionation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C in Table I-4.
  • the second light component below 350°C was tested in the second reaction unit to obtain the hydrocracking product.
  • the properties are shown in Table I-5.
  • Example I- The process is similar to that of Example I-1, except that the catalyst filling in the first reaction unit in Example I- is as follows:
  • the order of catalyst loading is the hydrogenation protection catalyst, the rich ore precursor material 2, the rich ore precursor material 1, the hydrodemetalization desulfurization catalyst, and the hydrodesulfurization catalyst.
  • the liquid phase product obtained by the fractionation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C in Table I-4.
  • the second light component below 350°C was tested in the second reaction unit to obtain hydrocracking products.
  • the properties are shown in Table I-5.
  • Example I- The process is similar to that of Example I-1, except that the catalyst filling in the first reaction unit in Example I- is as follows:
  • the order of catalyst loading is: hydrodesulfurization catalyst, hydrodesulfurization catalyst, hydrodesulfurization catalyst.
  • the liquid phase product obtained by the fractionation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C in Table I-4.
  • the second light component below 350°C was tested in the second reaction unit to obtain the hydrocracking product.
  • the properties are shown in Table I-5.
  • Example I- The process is similar to that of Example I-1, except that the catalyst filling in the first reaction unit in Example I- is as follows:
  • the order of catalyst loading is: hydrogenation protection catalyst, rich ore precursor material 3, hydrodemetalization desulfurization catalyst, and hydrodesulfurization catalyst.
  • the liquid phase product obtained by the fractionation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C in Table I-4.
  • the second light component below 350°C was tested in the second reaction unit to obtain the hydrocracking product.
  • the properties are shown in Table I-5.
  • the catalyst and device are similar to those in Example I-1. The difference is:
  • the aromatic-rich distillate QY (aromatic content of 20% by mass) in this comparative example I- does not pass through a partial hydrosaturation treatment device, but is directly mixed with DOA.
  • DOA and QY are mixed at a mass ratio of 1:10.
  • the properties of the mixed raw materials are shown in Table I-2.
  • Example I-1 the mixed raw materials of this comparative example I- were first mixed with hydrogen in the hydrogen dissolving unit, and the obtained mixed material entered the first reaction unit. After the first reaction unit was hydrotreated, the product properties were as shown in Table I-3.
  • the second light component below 350°C was tested on a fixed-bed hydrocracking unit to obtain hydrocracking products.
  • the properties are shown in Table I-5.
  • the catalyst and device are similar to those in Example I-1. The difference is:
  • the aromatic-rich distillate QY in this comparative example I- does not pass through a partial hydrosaturation treatment device, but is directly mixed with DOA.
  • DOA and QY are mixed at a mass ratio of 2:10.
  • the properties of the mixed raw materials are shown in Table I-2.
  • Example I-1 the mixed raw materials of this comparative example I- were first mixed with hydrogen in the hydrogen dissolving unit, and the obtained mixed material entered the first reaction unit. After the first reaction unit was hydrotreated, the product properties were as shown in Table I-3.
  • the second light component below 350°C was tested on a fixed-bed hydrocracking unit to obtain hydrocracking products.
  • the properties are shown in Table I-5.
  • the catalyst and device are similar to those in Example I-1. The difference is:
  • the aromatic-rich distillate QY in this comparative example I- does not pass through a partial hydrosaturation treatment device, but is directly mixed with DOA.
  • DOA and QY are mixed at a mass ratio of 3:10. Because there are a lot of solids in the mixed raw materials (at 100°C), the next test cannot be performed.
  • Example I-1 0.72 >92 ⁇ 10
  • Example I-2 0.72 >92 ⁇ 10
  • Example I-3 0.72 >92 ⁇ 10
  • Example I-4 0.72 >92 ⁇ 10
  • Example I-8 0.72 >92 ⁇ 10
  • Example I-10 0.72 >92 ⁇ 10
  • Example I-11 0.72 >92 ⁇ 10
  • Example I-12 0.72 >92 ⁇ 10
  • Example I-13 0.72 >92 ⁇ 10 Comparative Example I-1 >0.72 ⁇ 92 11 Comparative example I-2 >0.72 ⁇ 92 11
  • the solvent used is a hydrocarbon mixture with a butane content of 70% or more.
  • solvent: vacuum residue 3:1 (mass ratio)
  • Solvent deasphalting is carried out under the conditions, the mass yield of DAO is 70%, and the yield of DOA is 30%.
  • Example II-B The DAO and DOA used in Example II- are all from Example II-B.
  • liquid phase product of DAO after hydrogenation in the fourth reaction unit are shown in Table II-1; the liquid phase product enters the DCC unit for reaction to obtain LCO1 and HCO1.
  • LCO1 is hydrogenated and saturated in the third reaction unit and then fractionated to obtain the first light component 1 and the first heavy component 1.
  • the operating conditions for the hydrogenation of the third reaction unit are: the reaction temperature is 290°C, the reaction pressure is 4MPa, The liquid hourly volumetric space velocity is 1h -1 , and the hydrogen-to-oil volume ratio is 800:1.
  • the properties of LCO1 and the first heavy component 1 are shown in Table II-2.
  • DOA is mixed with the first heavy component 1 in a mass ratio of 1:10.
  • the properties of the mixed raw materials are shown in Table II-3.
  • the obtained mixture material (the hydrogen content is shown in Table II-3) in the first reaction unit operating conditions: reaction temperature of 360 °C, reaction pressure It is 10MPa, the liquid hourly volumetric space velocity is 0.3h -1 , and the circulating oil: the feedstock oil volume ratio at the inlet of the first reaction unit is 0.5:1.
  • the properties of the mixed raw materials after hydrogenation are shown in Table II-4.
  • the liquid phase product obtained by the fractional distillation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C, as shown in Table II-5.
  • the second light component below 350°C was tested in the second reaction unit to obtain hydrocracking products.
  • the properties are shown in Table II-6.
  • Example II-B The DAO and DOA used in Example II- are all from Example II-B.
  • HCO2 is hydrogenated and saturated in the third reaction unit and then fractionally distilled to obtain the first light component 2 and the first heavy component 2.
  • the hydrogenation operation conditions of the third reaction unit are: reaction temperature of 330°C, reaction pressure of 6MPa, liquid
  • the hourly volumetric space velocity is 1h -1
  • the hydrogen-to-oil volume ratio is 800:1.
  • the properties of HCO2 and the first heavy component 2 are shown in Table II-2.
  • DOA and the first heavy component 2 are mixed at a mass ratio of 5:10.
  • the properties of the mixed raw materials are shown in Table II-3.
  • DOA and the first heavy component 2 enter the hydrogen dissolving unit and mix with hydrogen.
  • the obtained mixture (the hydrogen content is shown in Table II-3) in the first reaction unit is operated under the following conditions: reaction temperature is 380°C, reaction pressure It is 8MPa, the liquid hourly volumetric space velocity is 0.3h -1 , and the circulating oil: the feedstock oil volume ratio at the inlet of the first reaction unit is 0.5:1.
  • the properties of the mixed raw materials after hydrogenation are shown in Table II-4.
  • the liquid phase product processed by the first reaction unit is fractionated, and the properties of the second reconstituted fraction at 350°C or higher are shown in Table II-5.
  • the second light component below 350°C was tested in the second reaction unit to obtain hydrocracking products.
  • the properties are shown in Table II-6.
  • Example II-B The DAO and DOA used in Example II- are all from Example II-B.
  • liquid phase products of DAO after hydrogenation in the fourth reaction unit are shown in Table II-1; the liquid phase products enter the DCC unit (operating conditions are the same as those in Example II-1) for reaction to obtain LCO1 and HCO1.
  • LCO1 is subjected to hydrogenation saturation in the third reaction unit and then fractionated to obtain the first light component 3 and the first heavy component 3.
  • the hydrogenation operation conditions of the third reaction unit are: reaction temperature of 320°C, reaction pressure of 6MPa, liquid
  • the hourly volumetric space velocity is 1h -1
  • the hydrogen-to-oil volume ratio is 800:1.
  • the properties of LCO1 and the first heavy component 3 are shown in Table II-2.
  • DOA and the first heavy component 3 are mixed at a mass ratio of 10:10.
  • the properties of the mixed raw materials are shown in Table II-3.
  • DOA and the first heavy component 3 enter the hydrogen dissolving unit and mix with hydrogen, and the obtained mixture material (the hydrogen content is shown in Table II-3) in the first reaction unit operating conditions: reaction temperature is 370 °C, reaction pressure It is 8MPa, the liquid hourly volumetric space velocity is 0.3h -1 , and the circulating oil: the feedstock oil volume ratio at the inlet of the first reaction unit is 0.5:1.
  • Table II-4 The properties of the mixed raw materials after hydrogenation are shown in Table II-4.
  • the liquid phase product obtained by the fractional distillation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C, as shown in Table II-5.
  • the second heavy component was subjected to a coking reaction at a reaction temperature of 500° C. and a residence time of 0.5 hours to obtain petroleum coke (yield of 31% by mass) and a sulfur content of 2.7% by mass.
  • the second light component below 350°C was tested in the second reaction unit to obtain hydrocracking products.
  • the properties are shown in Table II-6.
  • Example II-B The DAO and DOA used in Example II- are all from Example II-B.
  • liquid phase products of DAO after hydrogenation in the fourth reaction unit are shown in Table II-1; the liquid phase products enter the DCC unit (operating conditions are the same as those in Example II-1) for reaction to obtain LCO1 and HCO1.
  • the aromatic-rich distillate used in this example II is coal tar (see Table II-1 for properties) and LCO1 from a domestic coal coking unit.
  • the mass ratio of LCO1 to coal tar is 1:1, and the aromatic-rich distillate is in the first
  • the hydrogenation operation conditions of the third reaction unit are: the reaction temperature is 300° C., the reaction pressure is 10 MPa, and the liquid hour volume is empty.
  • the speed is 0.8h -1 , and the volume ratio of hydrogen to oil is 800:1.
  • the properties of the aromatic-rich distillate and the first heavy component 4 are shown in Table II-2.
  • DOA and the first heavy component 4 are mixed at a mass ratio of 15:10.
  • the properties of the mixed raw materials are shown in Table II-3.
  • DOA and the first heavy component 4 enter the hydrogen dissolving unit and mix with hydrogen, and the obtained mixture material (the hydrogen content is shown in Table II-3) in the first reaction unit operating conditions: reaction temperature of 350 °C, reaction pressure It is 12MPa, the liquid hourly volumetric space velocity is 0.3h -1 , and the circulating oil: the feedstock oil volume ratio at the inlet of the first reaction unit is 0.5:1.
  • Table II-4 The properties of the mixed raw materials after hydrogenation are shown in Table II-4.
  • the liquid phase product obtained by the fractional distillation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C, as shown in Table II-5.
  • the second light component below 350°C was tested in the second reaction unit to obtain hydrocracking products.
  • the properties are shown in Table II-6.
  • the liquid phase product obtained by the fractional distillation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C, as shown in Table II-5.
  • Example II-4 After the same mixed raw materials as in Example II-4 were hydrotreated by the first reaction unit, the reaction temperature was increased by 3°C every 30 days, and the hydrogenation test was stopped after a total of 360 days of operation.
  • the rich ore precursor material 1 and rich ore precursor material 2 initially loaded into the reactor become V-rich material 1 and vanadium-rich material 2 after the reaction. After roasting analysis, the V content is 69% by mass and 60% by mass, respectively. High-quality material of high-value V 2 O 5.
  • the second light component below 350°C in Example II-3 was subjected to a catalytic cracking test in a small catalytic cracking fixed fluidized bed test device.
  • the catalyst used was the catalytic cracking catalyst MLC-500 produced by the Changling Branch of Sinopec Catalyst Co., Ltd. ,
  • the reaction temperature is 540°C
  • the agent-oil ratio is 6, and the residence time is 3s.
  • the product gasoline quality yield was 40%, and the gasoline RON octane number was 93.
  • Example II- The process is similar to that of Example II-1, except that the second heavy component obtained in Example II- is introduced into the delayed coking unit for reaction to obtain coking gasoline, coking diesel and coking wax oil.
  • the operating conditions of the delayed coking unit are: the reaction temperature is 510°C, and the residence time is 0.6h.
  • the sulfur content of coker diesel oil is 0.26% by mass, the freezing point is -11°C, and the cetane number is 48.
  • the sulfur content of the coking wax oil is 1.12% by mass, and the freezing point is 32°C.
  • the yield of coking gasoline was 14.7%, the sulfur content was 0.10% by mass, and the MON was 61.8.
  • the reaction process The conditions are the same as in Example II-1.
  • the properties of the mixed coker diesel, coker wax oil and LCO1 oil and the properties of the first heavy component 8 are shown in Table II-2.
  • DOA comes from Example II-B and is mixed with the first heavy component 8 in a mass ratio of 1:10.
  • the properties of the mixed raw materials are shown in Table II-3.
  • the obtained mixture material (the hydrogen content is shown in Table II-3) in the first reaction unit is operated under the following conditions: reaction temperature is 360°C, reaction pressure It is 8MPa, the liquid hourly volumetric space velocity is 0.3h -1 , and the hydrogen-to-oil volume ratio is 800:1.
  • the properties of the mixed raw materials after hydrogenation are shown in Table II-4.
  • the liquid phase product obtained by the fractional distillation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C, as shown in Table II-5.
  • the second light component below 350°C was tested in the second reaction unit to obtain hydrocracking products.
  • the properties are shown in Table II-6.
  • Example II-1 The second light component below 350°C obtained in Example II-1 was tested on a diesel hydro-upgrading device to obtain a diesel component.
  • the operating conditions of the diesel hydro-upgrading device are as follows: the reaction temperature is 360°C, the reaction pressure is 12MPa, the hydrogen-to-oil volume ratio is 1000, and the liquid hourly volumetric space velocity is 1.0h -1 .
  • the properties of the obtained diesel components are 5ppm sulfur content, -33°C freezing point, and cetane number 53.
  • Example II- The process is similar to that of Example II-1, except that the catalyst filling in the first reaction unit in Example II- is as follows:
  • the order of catalyst loading is the hydrogenation protection catalyst, the rich ore precursor material 1, the hydrodemetalization desulfurization catalyst, and the hydrodesulfurization catalyst.
  • the liquid phase product obtained by the fractional distillation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C, as shown in Table II-5.
  • the second light component below 350°C was tested in the second reaction unit to obtain hydrocracking products.
  • the properties are shown in Table II-6.
  • Example II- The process is similar to that of Example II-1, except that the catalyst filling in the first reaction unit in Example II- is as follows:
  • the order of catalyst loading is the hydrogenation protection catalyst, the rich ore precursor material 2, the rich ore precursor material 1, the hydrodemetalization desulfurization catalyst, and the hydrodesulfurization catalyst.
  • the liquid phase product obtained by the fractional distillation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C, as shown in Table II-5.
  • the second light component below 350°C was tested in the second reaction unit to obtain hydrocracking products.
  • the properties are shown in Table II-6.
  • Example II- The process is similar to that of Example II-1, except that the catalyst filling in the first reaction unit in Example II- is as follows:
  • the order of catalyst loading is: hydrodesulfurization catalyst, hydrodesulfurization catalyst, hydrodesulfurization catalyst.
  • the liquid phase product obtained by the fractional distillation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C, as shown in Table II-5.
  • the second light component below 350°C was tested in the second reaction unit to obtain hydrocracking products.
  • the properties are shown in Table II-6.
  • Example II- The process is similar to that of Example II-1, except that the catalyst filling in the first reaction unit in Example II- is as follows:
  • the order of catalyst loading hydrogenation protection catalyst, rich ore precursor material 3, hydrodemetalization desulfurization catalyst, hydrodesulfurization catalyst.
  • the liquid phase product obtained by the fractionation of the first reaction unit has the properties of the second heavy component greater than or equal to 350°C, as shown in Table II-5.
  • the second light component below 350°C was tested in the second reaction unit to obtain the hydrocracking product.
  • the properties are shown in Table II-6.
  • the catalyst and device are similar to those in Example II-1. The difference is:
  • the aromatic-rich distillate QY (aromatic content of 20% by mass) in this comparative example II- does not pass through a partial hydrosaturation treatment device, but is directly mixed with DOA.
  • DOA and QY are mixed at a mass ratio of 1:10.
  • the properties of the mixed raw materials are shown in Table II-3.
  • the mixed raw material enters the hydrogen dissolving unit and is mixed with hydrogen.
  • the obtained mixed raw material (the hydrogen content is shown in Table II-3) is hydrotreated by the first reaction unit, and the product properties are shown in Table II-4.
  • the second light component below 350°C was tested on the second reaction unit to obtain the hydrocracking product.
  • the properties are shown in Table II-6.
  • the catalyst and device are similar to those in Example II-1. The difference is:
  • the aromatic-rich distillate QY in this comparative example II- does not go through a partial hydrosaturation treatment device, but is directly mixed with DOA.
  • DOA and QY are mixed at a mass ratio of 2:10.
  • the properties of the mixed raw materials are shown in Table II-3.
  • the mixed raw material enters the hydrogen dissolving unit and is mixed with hydrogen.
  • the obtained mixed raw material (the hydrogen content is shown in Table II-3) is hydrotreated by the first reaction unit, and the product properties are shown in Table II-4.
  • the second light component below 350°C was tested on the second reaction unit to obtain the hydrocracking product.
  • the properties are shown in Table II-6.
  • the catalyst and device are similar to those in Example II-1. The difference is:
  • the aromatic-rich distillate QY in Comparative Example II-3 did not pass through a partial hydrosaturation treatment device, but was directly mixed with DOA.
  • DOA and QY are mixed at a mass ratio of 3:10. Because there are a lot of solids in the mixed raw materials (at 100°C), the next test cannot be performed.
  • Table II-1 Properties of DOA, DAO, and liquid phase products after hydroprocessing in the fourth reaction unit
  • Example II-1 0.72 >92 ⁇ 10 Example II-2 0.72 >92 ⁇ 10 Example II-3 0.72 >92 ⁇ 10 Example II-4 0.72 >92 ⁇ 10 Example II-8 0.72 >92 ⁇ 10 Example II-10 0.72 >92 ⁇ 10 Example II-11 0.72 >92 ⁇ 10 Example II-12 0.72 >92 ⁇ 10 Example II-13 0.72 >92 ⁇ 10 Comparative Example II-1 >0.72 ⁇ 92 13 Comparative Example II-2 >0.72 ⁇ 92 12
  • the technology of the present invention can obtain high-quality raw materials for the production of low-sulfur marine fuel or low-sulfur coke products from DOA.
  • the technology of the present invention can obtain high-quality gasoline products that meet the National V standard.

Landscapes

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

Abstract

一种加工富芳馏分油的方法和***，包括：将富芳馏分油（20）引入至第三反应单元（21）中进行加氢饱和后分馏以获得第一轻组分和第一重组分（22）；将脱油沥青（4）和含芳烃物流引入至溶氢单元（23）中与氢气混合，并将混合后的物料引入至第一反应单元（7）中进行加氢反应，所述第一反应单元（7）为液相加氢反应单元；将来自所述第一反应单元（7）的液相产物进行分馏，得到第二轻组分（8）和第二重组分（9）；将第二轻组分（8）引入至第二反应单元（10）中进行反应；以及将第二重组分（9）引入至延迟焦化单元（11）中进行反应；或者将第二重组分（9）作为低硫船用燃料油组分。该处理工艺能够实现高价值利用DOA。

Description

一种加工富芳馏分油的方法和*** 技术领域
本发明涉及烃油加工领域，具体涉及一种加工富芳馏分油的方法和一种加工富芳馏分油的***。
背景技术
渣油高效转化是炼油企业的核心。而固定床渣油加氢是渣油高效转化的关键技术，具有产品质量好、工艺成熟等特点。
但渣油中高含量的沥青质和金属是固定床渣油加氢装置运转周期的制约因素。
为解决这一难题，中石化石油化工科学研究院开发的渣油溶剂脱沥青(脱金属)-加氢处理-催化裂化组合工艺技术(SHF)是从低价值减压渣油中最大限度生产车用清洁燃料并延长运转周期的创新技术，但由于脱油沥青质(DOA)软化点高，难于输送和利用，限制了SHF技术的推广。
向化工转型的渣油加氢-催化裂解(DCC)多产丙烯的新组合工艺，也是受限于渣油中的沥青质和金属的影响，加氢渣油氢含量低，渣油加氢的运转周期短，DCC丙烯收率低，影响组合技术的经济效益。
另外，2020年开始要实行硫质量分数≯0.5％的低硫船燃新标准和硫质量分数≯3.0％的低硫石油焦标准，如何低成本生产低硫船燃(低硫石油焦)技术也是目前急需解决的问题。
因此，将DOA转化成低硫船燃或生产低硫石油焦的原料是迫切需要解决的技术难题。
发明内容
本发明的目的是为了提供一种新的加工富芳馏分油的方法，使得能够在即便是较低的氢分压和较低的氢油比并在较高的空速下进行时，也能获得较好的加氢处理效果以及装置的长周期稳定运行。
为了实现上述目的，本发明的第一方面提供一种加工富芳馏分油的方法，该方法包括：
(1)将富芳馏分油引入至第三反应单元中进行加氢饱和后分馏以 获得第一轻组分和第一重组分，所述第一轻组分和所述第一重组分的切割点为100-250℃，所述第一重组分中的芳烃含量为大于等于20质量％；
(2)将脱油沥青和含有所述第一重组分的含芳烃物流引入至溶氢单元中与氢气混合，并将混合后的物料引入至第一反应单元中进行加氢反应，其中，所述第一反应单元中含有富矿前驱体材料和/或加氢催化剂，所述第一反应单元为液相加氢反应单元，所述富矿前驱体材料为能够吸附选自V、Ni、Fe、Ca和Mg中的至少一种金属的材料，所述脱油沥青和所述含芳烃物流的用量比使得由该脱油沥青和含芳烃物流形成的混合原料在不高于400℃时呈液态；
(3)将来自所述第一反应单元的液相产物进行分馏，得到第二轻组分和第二重组分，其中，所述第二轻组分和所述第二重组分的切割点为240～450℃；
(41)将所述第二轻组分引入至第二反应单元中进行反应以得到选自汽油组分、柴油组分和BTX原料组分中的至少一种产物，其中，所述第二反应单元选自加氢裂化单元、催化裂化单元和柴油加氢提质单元中的至少一种；以及
(42)将所述第二重组分引入至延迟焦化单元中进行反应以得到选自焦化汽油、焦化柴油、焦化蜡油和低硫石油焦中的至少一种产物；或者将所述第二重组分作为低硫船用燃料油组分。
本发明的第二方面提供一种加工富芳馏分油的***，该***中包括：
第三反应单元，该第三反应单元用于将富芳馏分油在其中进行加氢饱和和分馏以得到第一轻组分和第一重组分；
溶氢单元，该溶氢单元与所述第三反应单元保持流体连通，用于将脱油沥青和含有来自所述第三反应单元的第一重组分的含芳烃物流在其中与氢气混合；
第一反应单元，该第一反应单元为液相加氢反应单元且与所述溶氢单元保持流体连通，用于将所述溶氢单元的混合物料在其中进行加氢反应；
分离单元，该分离单元与所述第一反应单元保持流体连通，用于将来自所述第一反应单元的液相产物在其中进行分馏；
第二反应单元，该第二反应单元与所述分离单元保持流体连通，用于将由所述分离单元中获得的第二轻组分在其中进行反应，所述第二反应单元选自加氢裂化单元、催化裂化单元和柴油加氢提质单元中的至少一种；
延迟焦化单元，该延迟焦化单元与所述分离单元保持流体连通，用于将由所述分离单元中获得的第二重组分在其中进行反应以得到选自焦化汽油、焦化柴油、焦化蜡油和低硫石油焦中的至少一种产物；
出口，该出口与所述分离单元保持流体连通，用于将由所述分离单元中获得的第二重组分作为低硫船用燃料油组分引出***。
本发明提供的加工富芳馏分油的方法对渣油进行处理时，即使以较低的氢分压和较低的氢油比并在较高的空速下进行，也能获得较好的加氢处理效果以及装置的长周期稳定运行。
本发明特别适用于常渣与减渣的加氢转化，尤其适用于高金属、高残炭、高稠环物质、高氮含量的劣质渣油加氢转化。
本发明将脱油沥青(DOA)进行加氢处理的工艺方法，使得重油高效转化并能够生产汽油、BTX原料，以及能够灵活生产低硫船燃和低硫石油焦的***和方法。
附图说明
图1是本发明的一种优选的具体实施方式的加工富芳馏分油的工艺流程图。
图2是本发明的第一变体的具体实施方式的加工富芳馏分油的工艺流程图。
附图标记说明
1   重油原料                 2   溶剂脱沥青单元
3   脱沥青油                 4   脱油沥青
5   芳烃化合物               6   混合原料
7   第一反应单元             8   第二轻组分
9   第二重组分               10 第二反应单元
11  延迟焦化单元             12 BTX原料组分
13  汽油组分                 14 柴油组分
15  焦化汽油                 16 焦化柴油
17   焦化蜡油                 18   低硫石油焦
19   分离单元                 20   富芳馏分油
21   第三反应单元             22   第一重组分
23   溶氢单元                 24   第四反应单元
25   DCC单元                  26   丙烯
27   LCO                      28   HCO
29   油浆
具体实施方式
在本文中所披露的范围的端点和任何值都不限于该精确的范围或值，这些范围或值应当理解为包含接近这些范围或值的值。对于数值范围来说，各个范围的端点值之间、各个范围的端点值和单独的点值之间，以及单独的点值之间可以彼此组合而得到一个或多个新的数值范围，这些数值范围应被视为在本文中具体公开。
如前所述，本发明的第一方面提供了一种加工富芳馏分油的方法，该方法包括：
(1)将富芳馏分油引入至第三反应单元中进行加氢饱和后分馏以获得第一轻组分和第一重组分，所述第一轻组分和所述第一重组分的切割点为100-250℃，所述第一重组分中的芳烃含量为大于等于20质量％；
(2)将脱油沥青和含有所述第一重组分的含芳烃物流引入至溶氢单元中与氢气混合，并将混合后的物料引入至第一反应单元中进行加氢反应，其中，所述第一反应单元中含有富矿前驱体材料和/或加氢催化剂，所述第一反应单元为液相加氢反应单元，所述富矿前驱体材料为能够吸附选自V、Ni、Fe、Ca和Mg中的至少一种金属的材料，所述脱油沥青和所述含芳烃物流的用量比使得由该脱油沥青和含芳烃物流形成的混合原料在不高于400℃时呈液态；
(3)将来自所述第一反应单元的液相产物进行分馏，得到第二轻组分和第二重组分，其中，所述第二轻组分和所述第二重组分的切割点为240～450℃；
(41)将所述第二轻组分引入至第二反应单元中进行反应以得到选自汽油组分、柴油组分和BTX原料组分中的至少一种产物，其中， 所述第二反应单元选自加氢裂化单元、催化裂化单元和柴油加氢提质单元中的至少一种；以及
(42)将所述第二重组分引入至延迟焦化单元中进行反应以得到选自焦化汽油、焦化柴油、焦化蜡油和低硫石油焦中的至少一种产物；或者将所述第二重组分作为低硫船用燃料油组分。
优选地，所述脱油沥青和所述含芳烃物流的用量比使得由该脱油沥青和含芳烃物流形成的混合原料在不高于280℃时呈液态；进一步优选所述脱油沥青和所述含芳烃物流的用量比使得由该脱油沥青和含芳烃物流形成的混合原料在不高于100℃时呈液态。
本发明优选所述第三反应单元中进行的加氢饱和反应为部分加氢饱和，特别优选所述第一轻组分和所述第一重组分的切割点为180℃。
优选情况下，在本发明的溶氢单元中的操作条件包括：氢气的送入量与所述脱油沥青和所述含芳烃物流形成的混合原料的体积比(即氢油体积比)为30-200，更优选为50-150，溶氢单元操作温度为300-450℃，压力为2-20MPa。
根据本发明的方法，溶氢单元中与氢气混合后获得的混合后的物料能够以向上流动的方式进入第一反应单元，也可以以向下流动的方式进入第一反应单元。优选地，溶氢单元中与氢气混合后获得的混合后的物料以向上流动的方式进入第一反应单元这样在反应过程中，溶解并分散在油料中的氢气基本不会聚集形成大的气泡并逸出，从而能够为加氢反应提供足够的氢源，获得更好的加氢处理效果，并进一步降低催化剂结焦趋势，使催化剂保持较高的催化活性，进一步延长催化剂的使用寿命和装置的稳定运行周期。
所述第一轻组分优选进入催化裂化单元生产低碳烯烃。本发明对第一轻组分进入催化裂化单元生产低碳烯烃的具体操作条件没有特别的限制。
特别优选地，所述第二轻组分和所述第二重组分的切割点为350℃。
优选地，在步骤(2)中，所述脱油沥青和所述含芳烃物流的用量比使得由该脱油沥青和含芳烃物流形成的混合原料的100℃粘度不大于400mm 2/s，更优选不大于200mm 2/s，进一步优选不大于100mm 2/s。
优选情况下，在步骤(2)中，所述含芳烃物流中还含有芳烃化合 物和/或芳烃油，所述芳烃油选自LCO、HCO、FGO(催化重馏出油)、乙烯焦油、煤焦油、焦化柴油和焦化蜡油中的至少一种。
优选地，所述芳烃化合物选自苯、甲苯、二甲苯、萘、甲基萘、多支链萘及双环以上芳烃中的一种或几种，优选为环数不超过三环的多环芳烃或它们的混合物。特别优选情况下，所述芳烃化合物选自苯、甲苯、二甲苯、萘、由至少一种C 1-6的烷基取代的萘、三环以上芳烃中的至少一种。
更优选地，所述富芳馏分油中的芳烃含量大于等于20质量％，优选大于等于25质量％，优选大于等于40质量％，更优选大于等于60质量％。
优选情况下，在步骤(2)中，所述脱油沥青为由重油原料进入溶剂脱沥青单元中进行溶剂脱沥青处理后得到的脱油沥青。
优选地，在所述溶剂脱沥青单元中，所述脱油沥青的收率质量分数不大于50％，更优选不大于40％，进一步优选不大于30％。
根据一种优选的具体实施方式，在步骤(2)中，所述脱油沥青与所述含芳烃物流的用量质量比为1:10～50:10，更优选为2:10～30:10；进一步优选为3:10～15:10。
优选地，本发明的该方法还包括：将步骤(42)中获得的所述焦化柴油和/或所述焦化蜡油循环回步骤(1)中的所述第一反应单元进行加氢饱和。
优选地，在步骤(1)中，所述第三反应单元为固定床反应器、移动床反应器和沸腾床反应器中的至少一种反应器。
优选情况下，所述第三反应单元中的操作条件包括：反应温度为200-420℃，反应压力为2-18MPa，液时体积空速为0.3-10h -1，氢油体积比50-5000；更优选地，所述第三反应单元中的操作条件包括：反应温度为220-400℃，反应压力为2-15MPa，液时体积空速为0.3-5h -1，氢油体积比为50-4000。
以下提供本发明的第三反应单元中的优选的具体实施方式：
有氢存在的富芳馏分油的部分加氢饱和的条件通常如下：富芳馏分油的部分加氢饱和技术为固定床/沸腾床/移动床加氢处理技术。以目前工业上较成熟固定床柴油或蜡油加氢技术为例，所述反应器或反应床层至少包括一种加氢精制催化剂。富芳馏分油的部分加氢饱和中应 用的加氢精制催化剂优选具有良好且适中的加氢饱和活性，以避免四氢萘类结构进一步饱和为供氢能力较低的十氢萘或环烷烃结构。这些催化剂一般都是以多孔耐熔无机氧化物如氧化铝为载体，第ⅥB族和/或Ⅷ族金属如W、Mo、Co、Ni等的氧化物为活性组分，选择性地加入其它各种助剂如P、Si、F、B等元素的催化剂，例如由石油化工科学研究院研发的RS系列预处理催化剂就属于这类催化剂。RS系列催化剂是一种NiMo催化剂。
本发明特别优选所述第一反应单元为渣油液相加氢反应器。
优选情况下，在步骤(2)中，所述第一反应单元中的操作条件包括：反应温度260～500℃，反应压力为2.0～20.0MPa，循环油与所述第一反应单元入口原料油的体积比例为0.1:1至15:1，液时体积空速为0.1～1.5h -1，液时体积空速为0.1～1.5h -1。液时体积空速和反应压力可以根据待处理物料的特性和要求的转化率及精制深度进行选择的。本发明由脱油沥青和含芳烃物流形成的混合原料与氢气混合后可以从第一反应单元的反应器的顶部进入，自上向下下行穿过催化剂床层；也可以从第一反应单元的反应器的底部进入，自下向上上行穿过催化剂床层。
优选地，在步骤(2)中，所述富矿前驱体材料中含有载体和负载在所述载体上的活性组分元素，所述载体选自氢氧化铝、氧化铝和氧化硅中的至少一种，所述活性组分元素选自第VIB族和VIII族金属元素中的至少一种。更优选地，所述富矿前驱体材料中的活性组分为选自第VIB族和VIII族金属元素的氧化物和/或硫化物。
优选情况下，在步骤(2)中，所述富矿前驱体材料的灼减不低于3质量％，比表面积不低于80m 2/g，吸水率不低于0.9g/g。所述灼减是指富矿前驱体材料在600℃/2h焙烧处理后减少的质量占焙烧前质量的百分比例；所述吸水率是指富矿前驱体材料室温(例如25℃)下浸泡水中半小时增加的质量占浸泡前质量的百分比例。
根据一种优选的具体实施方式，在步骤(2)中，按照反应物流方向，所述第一反应单元中依次装填有第一富矿前驱体材料和第二富矿前驱体材料，且所述第二富矿前驱体材料的灼减大于等于所述第一富矿前驱体材料的灼减。
根据前述优选的具体实施方式，进一步优选地，所述第一富矿前 驱体材料的灼减为3-15质量％，以及所述第二富矿前驱体材料的灼减为不小于15质量％。
根据前述优选的具体实施方式，进一步优选地，所述第一富矿前驱体材料与所述第二富矿前驱体材料的装填体积比为5:95至95:5。
本发明所述的加氢催化剂可以为不同催化剂的级配组合，优选加氢催化剂至少能够催化加氢脱金属反应和加氢脱硫反应。
本发明对能够催化加氢脱金属反应、加氢脱硫反应、加氢脱沥青反应和加氢脱残炭反应的催化剂的具体种类没有特别的限定，可以采用本领域内常规应用的能够催化上述反应的催化剂。
本发明的所述加氢催化剂例如可以为以多孔耐熔无机氧化物为载体，第VIB族和/或VIII族金属的氧化物或硫化物为活性组分，选择性地加入助剂的催化剂。
优选情况下，本发明的第一反应单元在长周期运行后，富矿前驱体材料能够转变为富钒材料，富钒材料中的钒含量不小于10质量％；特别优选情况下，所述富矿前驱体材料转变为V含量20质量％以上的富钒材料，能够直接炼制高价值的V 2O 5。
以下提供本发明的第一反应单元中的优选的具体实施方式：
本发明的所述第一反应单元中涉及的原料加氢处理技术为液相加氢处理技术，所述反应器或反应床层至少包括一种富矿前驱体材料和/或一种加氢催化剂，富矿前驱体材料主要由两部分组成：一是吸附油中含钒有机化合物能力强的载体，二是具有加氢活性功能的活性组分。所述载体主要由氧化硅、氢氧化铝或氢氧化铝/氧化铝混合物挤条成型、干燥得到，表面富含丰富的-OH，对油中含钒有机化合物有强的吸附能力，600℃焙烧2h，其灼减不低于5质量％。活性组分主要采用第VIB族和/或VIII族金属如W、Mo、Co、Ni等的氧化物或硫化物。
前述优选的具体实施方式中涉及的加氢催化剂一般为重渣油加氢催化剂，重渣油加氢催化剂是指具有重、渣油加氢脱金属、加氢脱硫和加氢脱残炭等功能的组合催化剂。这些催化剂一般都是以多孔耐熔无机氧化物如氧化铝为载体，第VIB族和/或VIII族金属如W、Mo、Co、Ni等的氧化物或硫化物为活性组分，选择性地加入其它各种助剂如P、Si、F、B等元素的催化剂，例如由石油化工科学研究院研发的RDM、RCS系列重、渣油加氢脱金属催化剂和脱硫催化剂。目前在液 相加氢处理技术中，经常是多种催化剂配套使用。本发明中优选有富矿前驱体材料、加氢脱金属脱硫催化剂、加氢脱硫催化剂，装填顺序一般是使原料依次与富矿前驱体材料、加氢脱金属脱硫、加氢脱硫催化剂接触，有时也可根据情况，少装一种或两种催化剂，例如只装填富矿前驱体材料和加氢脱硫催化剂，而不装加氢脱金属脱硫催化剂。当然也有将这几种催化剂混合装填的技术。
根据一种优选的具体实施方式，在步骤(41)中，所述第二反应单元为加氢裂化单元，且所述加氢裂化单元中的操作条件包括：反应温度为360～420℃，反应压力为10.0～18.0MPa，氢油体积比为600～2000，液时体积空速为1.0～3.0h -1。
优选地，所述加氢裂化单元中装填有至少一种加氢处理催化剂和至少一种加氢裂化催化剂。
优选地，所述加氢裂化单元为固定床加氢裂化单元。
当所述第二反应单元为加氢裂化单元时，以下提供本发明的第二反应单元中的优选的具体实施方式：
在步骤(41)中，将所述第二轻组分引入至第二反应单元中进行反应，采用的加氢裂化技术为固定床加氢裂化技术。以目前工业上较成熟固定床蜡油加氢裂化技术为例，所述反应器或反应床层至少包括两种加氢裂化催化剂，一种是预处理催化剂，一种是加氢裂化催化剂。由于经液相加氢处理技术后又经分馏得到的物料中金属含量、硫、氮含量及残炭值都较高，因此预处理催化剂优选具有很强的脱金属活性和很好的脱硫、脱氮活性，以保证后面的加氢裂化催化剂的活性。加氢裂化催化剂优选具有很好的加氢裂化活性和高VGO转化与HDS活性。这些催化剂一般都是以多孔耐熔无机氧化物如氧化铝或分子筛为载体，第ⅥB族和/或Ⅷ族金属如W、Mo、Co、Ni等的氧化物为活性组分，选择性地加入其它各种助剂如P、Si、F、B等元素的催化剂，例如由石油化工科学研究院研发的RS系列预处理催化剂和RHC系列加氢裂化催化剂就属于这类催化剂。RS系列催化剂是一种NiW催化剂，RHC系列催化剂是一种NiMo分子筛催化剂。
根据另一种优选的具体实施方式，在步骤(41)中，所述第二反应单元为催化裂化单元，且所述催化裂化单元为流化催化裂化(FCC)单元。
根据另一种优选的具体实施方式，催化裂化单元中采用的第二轻组分催化裂化技术为流化床催化裂化(FCC)技术，优选采用石油化工科学研究院开发的LTAG技术，主要生产汽油馏分和液化气。
优选地，所述流化催化裂化单元中的操作条件包括：反应温度为500～600℃，剂油比为3～12，停留时间为0.6～6s。
在没有特别说明的情况下，本发明所述剂油比均表示剂油质量比。
根据另一种优选的具体实施方式，在步骤(41)中，所述第二反应单元为柴油加氢提质单元，且所述柴油加氢提质单元中的操作条件包括：反应温度为330～420℃，反应压力为5.0～18.0MPa，氢油体积比为500～2000，液时体积空速为0.3～3.0h -1。
优选地，所述柴油加氢提质单元中装填有至少一种柴油加氢提质催化剂。
本发明的所述柴油加氢提质催化剂例如可以为石油化工科学研究院研发的RS系列预处理催化剂和RHC-100系列柴油加氢裂化催化剂。
根据一种优选的具体实施方式，在步骤(42)中，将所述第二重组分引入至延迟焦化单元中进行反应以得到选自焦化汽油、焦化柴油、焦化蜡油和低硫石油焦中的至少一种产物，且所述至延迟焦化单元中的操作条件包括：反应温度为440～520℃，停留时间为0.1～4h。
根据另一种优选的具体实施方式，在步骤(42)中，所述第二重组分的硫含量不大于1.8质量％，将所述第二重组分引入至延迟焦化单元中进行反应以得到低硫石油焦，更优选所述低硫石油焦的硫含量不大于3质量％。
优选情况下，在步骤(42)中，将所述第二重组分作为低硫船用燃料油组分，且控制条件使得所述低硫船用燃料油组分中的硫含量不大于0.5质量％。
本发明对所述溶剂脱沥青处理的具体操作没有特别的限制，可以采用本领域内常规的溶剂脱沥青工艺进行。本发明的实例中示例性地列举了溶剂脱沥青工艺的操作参数，本领域技术人员不应理解为对本发明的限制。
本发明适用于常渣与减渣的加氢转化，尤其适用于高金属(Ni+V>150μg/g，尤其Ni+V>200μg/g)、高残炭(残炭质量分数>17％，尤其残炭质量分数>20％)、高稠环物质的劣质渣油加氢转化。
如前所述，本发明的第二方面提供了一种加工富芳馏分油的***，该***中包括：
第三反应单元，该第三反应单元用于将富芳馏分油在其中进行加氢饱和和分馏以得到第一轻组分和第一重组分；
溶氢单元，该溶氢单元与所述第三反应单元保持流体连通，用于将脱油沥青和含有来自所述第三反应单元的第一重组分的含芳烃物流在其中与氢气混合；
第一反应单元，该第一反应单元为液相加氢反应单元且与所述溶氢单元保持流体连通，用于将所述溶氢单元的混合物料在其中进行加氢反应；
分离单元，该分离单元与所述第一反应单元保持流体连通，用于将来自所述第一反应单元的液相产物在其中进行分馏；
第二反应单元，该第二反应单元与所述分离单元保持流体连通，用于将由所述分离单元中获得的第二轻组分在其中进行反应，所述第二反应单元选自加氢裂化单元、催化裂化单元和柴油加氢提质单元中的至少一种；
延迟焦化单元，该延迟焦化单元与所述分离单元保持流体连通，用于将由所述分离单元中获得的第二重组分在其中进行反应以得到选自焦化汽油、焦化柴油、焦化蜡油和低硫石油焦中的至少一种产物；
出口，该出口与所述分离单元保持流体连通，用于将由所述分离单元中获得的第二重组分作为低硫船用燃料油组分引出***。
优选地，所述延迟焦化单元与所述溶氢单元保持流体连通，用于将所述延迟焦化单元中获得的所述焦化柴油和/或所述焦化蜡油循环回所述第一反应单元中。
优选情况下，该***中还包括溶剂脱沥青单元，该溶剂脱沥青单元与所述溶氢单元保持流体连通，用于将重油原料在其中进行溶剂脱沥青处理，并将所述溶剂脱沥青处理后得到的脱油沥青引入至所述溶氢单元中。
根据一种优选的具体实施方式，在本发明的***中，所述第二反应单元为加氢裂化单元。
根据另一种优选的具体实施方式，在本发明的***中，所述第二反应单元为催化裂化单元，且所述催化裂化单元为流化催化裂化单元。
根据另一种优选的具体实施方式，在本发明的***中，所述第二反应单元为柴油加氢提质单元。
本发明还提供了所述方法的第一变体，在该第一变体中，还包括：
(11)将重质原料油引入至溶剂脱沥青单元中进行溶剂脱沥青处理，得到脱油沥青和脱沥青油；
(12)将所述脱沥青油引入至第四加氢单元反应单元中进行加氢反应，并将所述第四加氢单元反应单元中获得的液相流出物引入至DCC单元进行反应，得到丙烯、LCO、HCO和油浆，其中，所述第四加氢单元反应单元为固定床加氢单元反应单元；
(1)将含有来自所述DCC单元的LCO和/或HCO的富芳馏分油用作所述步骤(1)中富芳馏分油。
在该第一变体中，优选地，本发明的该方法还包括：将步骤(42)中获得的所述焦化柴油和/或所述焦化蜡油循环回所述第三反应单元中进行加氢饱和。
优选情况下，在步骤(12)中，所述第四反应单元的操作条件包括：反应温度为280～400℃，反应压力为6.0～14.0MPa，氢油体积比为600～1200，液时体积空速为0.3～2.0h -1。
优选地，在步骤(12)中，所述第四反应单元中装填有至少两种加氢催化剂。更优选地，在步骤(12)中，所述加氢催化剂为能够催化选自加氢脱金属反应、加氢脱硫反应和加氢脱残炭反应中的至少一种反应的催化剂。所述加氢催化剂一般都是以多孔耐熔无机氧化物如氧化铝为载体；特别优选情况下，在步骤(12)中，所述加氢催化剂中含有作为载体的氧化铝和作为活性组分元素的第VIB族和/或VIII族金属元素，且该加氢催化剂中任选还含有选自P、Si、F和B中的至少一种助剂元素。在所述加氢催化剂中，所述第VIB族和VIII族金属元素例如可以为W、Mo、Co、Ni等。并且，在所述加氢催化剂中，所述活性组分可以为上述活性组分元素的氧化物和/或硫化物。
以下提供本发明的第四反应单元中的优选的具体实施方式：
有氢存在的脱沥青油(DAO)的第四反应单元的条件通常如下：DAO的加氢处理技术为固定床加氢处理技术。以目前工业上较成熟固定床重、渣油加氢技术为例，所述反应器或反应床层至少包括两种加氢催化剂，采用的重渣油加氢催化剂是指具有重、渣油加氢脱金属、 加氢脱硫、加氢脱氮和加氢脱残炭等功能的组合催化剂。这些催化剂一般都是以多孔耐熔无机氧化物如氧化铝为载体，第ⅥB族和/或Ⅷ族金属如W、Mo、Co、Ni等的氧化物或硫化物为活性组分，选择性地加入其它各种助剂如P、Si、F、B等元素的催化剂，例如由石油化工科学研究院研发的RDM、RCS系列重、渣油加氢脱金属催化剂和脱硫催化剂。目前在固定床渣油加氢技术中，经常是多种催化剂配套使用，其中有加氢脱金属催化剂、加氢脱硫催化剂、加氢脱氮催化剂，装填顺序一般是使原料油依次与加氢脱金属、加氢脱硫、加氢脱氮催化剂接触，有时也可根据情况，少装一种或两种催化剂，例如只装填加氢脱金属催化剂和加氢脱硫催化剂，而不装加氢脱氮催化剂。当然也有将这几种催化剂混合装填的技术。
以下结合图1和2对本发明的加工富芳馏分油的方法进行进一步详细说明。
如图1所示，将富芳馏分油20引入至第三反应单元21中进行加氢饱和后分馏以获得第一轻组分和第一重组分22；以及重油原料1进入溶剂脱沥青单元2中进行溶剂脱沥青处理后得到的脱油沥青4和脱沥青油3；脱油沥青4与含有所述第一重组分22的含芳烃物流一起形成混合原料6并进入溶氢单元23中与氢气混合，由此获得的混合物料进入第一反应单元7中进行加氢反应，所述含芳烃物流中优选还含有来自外界的芳烃化合物5，其中，所述第一反应单元中含有富矿前驱体材料和能够催化选自加氢脱金属反应、加氢脱硫反应、加氢脱沥青反应和加氢脱残炭反应中的至少一种反应的加氢催化剂，所述第一反应单元为液相加氢反应单元；来自所述第一反应单元7的液相产物进入分离单元19中进行分馏，得到第二轻组分8和第二重组分9，其中，所述第二轻组分和所述第二重组分的切割点为240～450℃；将所述第二轻组分8引入至第二反应单元10中进行反应以得到选自汽油组分13、BTX原料组分12、柴油组分14中的至少一种产物，其中，所述第二反应单元选自加氢裂化单元、催化裂化单元和柴油加氢提质单元中的至少一种；以及将所述第二重组分9引入至延迟焦化单元11中进行反应以得到选自焦化汽油15、焦化柴油16、焦化蜡油17和低硫石油焦18中的至少一种产物；或者将所述第二重组分9作为低硫船用燃料油组分。
如图2所示，重油原料1进入溶剂脱沥青单元2中进行溶剂脱沥青处理后得到脱油沥青4和脱沥青油3；将所述脱沥青油3引入至第四反应单元24中进行加氢反应，并将所述第四反应单元24中获得的液相流出物引入至DCC单元25中进行反应，得到丙烯26、LCO27、HCO28和油浆29；将含有来自所述DCC单元25的LCO27和/或HCO28的富芳馏分油20引入至第三反应单元21中进行加氢饱和后分馏以获得第一重组分22和第一轻组分；将脱油沥青4和含有所述第一重组分22的含芳烃物流一起形成混合原料6并引入至溶氢单元29中与氢气混合，并将混合后的物料引入至第一反应单元7中进行加氢反应，所述含芳烃物流中优选还含有来自外界的芳烃化合物5，其中，所述第一反应单元7中含有富矿前驱体材料和能够催化选自加氢脱金属反应、加氢脱硫反应、加氢脱沥青反应和加氢脱残炭反应中的至少一种反应的加氢催化剂；来自所述第一反应单元7的液相产物进入分离单元19中进行分馏，得到第二轻组分8和第二重组分9；将所述第二轻组分8引入至第二反应单元10中进行反应以得到选自汽油组分13、BTX原料组分12、柴油组分14中的至少一种产物，或者将所述第二轻组分8循环回所述DCC单元25中；以及将所述第二重组分9引入至延迟焦化单元11中进行反应以得到选自焦化汽油15、焦化柴油16、焦化蜡油17和低硫石油焦18中的至少一种产物；或者将所述第二重组分9作为低硫船用燃料油组分。
本发明的技术使得重油高效转化并能够生产汽油、BTX原料，以及能够灵活生产低硫船燃和低硫石油焦的***和方法。
与现有技术相比，优选情况下，本发明由于采用了渣油加氢、加氢裂化或催化裂化等工艺的有机联合，不但使低价值的DOA转化成符合环保要求的低硫船燃组分和低硫石油焦原料，而且实现了重质石油资源的高效、环保和综合利用。
另外，本发明提供的技术能够使得DOA在渣油液相加氢反应器中高效转化并能够生产汽油馏分、BTX原料，以及能够提供生产低硫船燃和低硫焦产品的原料。
以下将通过实例对本发明进行详细描述。在没有特别说明的情况下，以下实例均采用图1所示的工艺流程进行。以及，在没有特别说明的情况下，以下实例具有如下共同特征：
在没有特别说明的情况下，以下实例中的表I-3和表II-4的结果为装置持续运行100h中，每25h取样检测获得的结果的平均值。
富芳馏分油部分加氢饱和实验在中型固定床柴油加氢处理装置上进行试验，反应器总体积为200mL。在以下实例中，富芳馏分油部分加氢饱和使用的加氢催化剂和材料是由石油化工科学研究院研发的RS-2100系列加氢催化剂。
将部分加氢饱和得到的液相物流进行分馏，得到切割点为180℃的第一轻组分和第一重组分，第一重组分和DOA形成混合原料。混合原料的加氢反应在中型重油液相加氢处理装置上进行试验，反应器总体积为200mL。在以下实例中。第一反应单元中使用的加氢催化剂和材料是由石油化工科学研究院研发的RG-30B保护催化剂、富矿前驱体材料1、富矿前驱体材料2、RDM-33B渣油脱金属脱硫过渡催化剂、RCS-31脱硫催化剂。按照物流方向，催化剂装填的顺序为加氢保护催化剂、富矿前驱体材料1、富矿前驱体材料2、加氢脱金属脱硫催化剂、加氢脱硫催化剂。第一反应单元中，各催化剂之间的装填比为：RG-30B:富矿前驱体材料1：富矿前驱体材料2：RDM-33B:RCS-31＝6：30：30：14：20(V/V)。
第二反应单元为固定床加氢裂化装置，所用的催化剂为石油化工科学研究院研发的RS-2100精制催化剂、RHC-131加氢裂化催化剂。各催化剂之间的装填比为：RS-2100:RHC-131＝40：60(V/V)。固定床加氢裂化装置的操作条件为：精制段反应温度为370℃，裂化段反应温度为385℃，反应压力为10MPa，液时体积空速为2.0h -1，氢油体积比为：1200：1。
实施例A
富矿前驱体材料1制备：选取中石化催化剂有限公司长岭分公司生产的RPB110拟薄水铝石2000g，其中1000g在550℃下处理2h，得到约700g氧化铝，将约700g氧化铝和另外1000g拟薄水铝石充分混合，之后加入40g田菁粉和20g柠檬酸，并加入2200g去离子水，混捏并挤条成型，在300℃下干燥3h，得到约1730g载体，加入2100mL含Mo和Ni的溶液进行饱和浸渍，溶液中Mo含量以MoO 3质量计为5.5％，Ni含量以NiO质量计为1.5％，浸渍半小时，之后在180℃下处 理4h，得到富矿前驱体材料1，性质如表I-6中所示。
富矿前驱体材料2制备：选取中石化催化剂有限公司长岭分公司生产的RPB110拟薄水铝石2000g，加入30g田菁粉和30g柠檬酸，并加入2400g去离子水，混捏并挤条成型，在120℃下干燥5h，得到约2040g载体，加入2200mL含Mo和Ni的溶液进行饱和浸渍，溶液中Mo含量以MoO 3质量计为7.5％，Ni含量以NiO质量计为1.7％，浸渍半小时，之后在200℃下处理3h，得到富矿前驱体材料2，性质如表I-6中所示。
富矿前驱体材料3制备：选取市售的氧化硅2000g，加入30g田菁粉和30g氢氧化钠，并加入2400g去离子水，混捏并挤条成型，在120℃下干燥5h，得到载体，加入2200mL含Mo和Ni的溶液进行饱和浸渍，溶液中Mo含量以MoO 3重量计为4.5％，Ni含量以NiO重量计为1.0％，浸渍半小时，之后在200℃下处理3h，得到富矿前驱体材料3，性质如表I-6中所示。
实施例I-1
本实施例I-采用的富芳馏分油为LCO，来自上海石化RLG装置，LCO加氢操作条件为：反应温度为290℃，反应压力为4MPa，液时体积空速为1h -1，氢油体积比为800：1。
LCO性质和第一重组分1的性质如表I-1所示。
DOA来自一种减压渣油，与第一重组分1按照质量比1:10混合，混合原料的性质见表I-2。
DOA和第一重组分1的混合原料先在溶氢单元(氢气的送入量与所述脱油沥青和所述第一重组分1的混合原料的体积比为100，溶氢单元操作温度为320℃，压力为10MPa)中与氢气混合，获得的混合物料进入第一反应单元，第一反应单元的操作条件为：反应温度为360℃，反应压力为10MPa，液时体积空速为0.6h -1，循环油：第一反应单元入口原料油体积比例0.5:1。混合原料加氢后产品性质见表I-3。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表I-4。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表I-5。
实施例I-2
本实施例I-采用的富芳馏分油为HCO，来自上海石化催化裂化装置，HCO加氢操作条件为：反应温度为330℃，反应压力为6MPa，液时体积空速为1h -1，氢油体积比为800：1。
HCO性质和第一重组分2性质如表I-1所示。
DOA来自一种减压渣油，与第一重组分2按照质量比5:10混合，混合原料的性质见表I-2。
DOA和加氢后HCO第一重组分2的混合原料先在溶氢单元(氢气的送入量与所述脱油沥青和所述第一重组分2的混合原料的体积比为100，溶氢单元操作温度为320℃，压力为10MPa)中与氢气混合，获得的混合物料进入第一反应单元，在第一反应单元的操作条件为：反应温度为380℃，反应压力为10MPa，液时体积空速为0.6h -1，循环油：第一反应单元入口原料油体积比例0.5:1。混合原料加氢后产品性质见表I-3。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表I-4。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表I-5。
实施例I-3
本实施例I-采用的富芳馏分油为与实施例I-1中相同的LCO，LCO加氢操作条件为：反应温度为320℃，反应压力为6MPa，液时体积空速为1h -1，氢油体积比为800：1。
LCO性质和第一重组分3性质如表I-1所示。
DOA来自一种减压渣油，与第一重组分3按照质量比10:10混合，混合原料的性质见表I-2。
DOA和第一重组分3的混合原料先在溶氢单元(氢气的送入量与所述脱油沥青和所述第一重组分3的混合原料的体积比为100，溶氢单元操作温度为320℃，压力为8MPa)中与氢气混合，获得的混合物料进入第一反应单元，在第一反应单元的操作条件为：反应温度为370℃，反应压力为8MPa，液时体积空速为0.6h -1，循环油：第一反应单元入口原料油体积比例0.5:1。混合原料加氢后产品性质见表I-3。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表I-4。
将第二重组分在反应温度为500℃停留时间为0.5小时下进行焦化反应，得到石油焦(收率32质量％)，硫含量为2.7质量％。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表I-5。
实施例I-4
本实施例I-采用的富芳馏分油为煤焦油，来自国内某煤焦化装置，煤焦油加氢操作条件为：反应温度为300℃，反应压力为10MPa，液时体积空速为0.8h -1，氢油体积比为800：1。
煤焦油性质和第一重组分4性质如表I-1所示。
DOA来自一种减压渣油，与第一重组分4按照质量比15:10混合，混合原料的性质见表I-2。
DOA和第一重组分4的混合原料先在溶氢单元(氢气的送入量与所述脱油沥青和所述第一重组分4的混合原料的体积比为100，溶氢单元操作温度为320℃，压力为12MPa)中与氢气混合，获得的混合物料进入第一反应单元，在第一反应单元的操作条件为：反应温度为350℃，反应压力为12MPa，液时体积空速为0.6h -1，循环油：第一反应单元入口原料油体积比例2:1。混合原料加氢后产品性质见表I-3。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表I-4。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表I-5。
实施例I-5
采用与实施例I-3相似的方法进行，所不同的是：
本实施例I-中，第一反应单元的加氢处理的温度为395℃。
其余条件与实施例I-3中相同。
混合原料加氢后产品性质见表I-3。
所得>350℃第二重组分主要物化性质见表I-3。
实施例I-6
原料、催化剂装填和重油液相加氢处理装置的操作条件等均同实施例I-1。所不同的是：
与实施例I-1相同的混合原料经液相重油加氢处理后，每过30天，反应温度提3℃，加氢试验共计运行360天后停止运转。
初始装到反应器的富矿前驱体材料1和富矿前驱体材料2，反应后变成富V材料1和富钒材料2，经焙烧分析其V含量分别为76质量％和71质量％，其钒含量比自然矿石高10倍以上，是提炼高价值V 2O 5的高品质材料。
实施例I-7
将实施例I-3中的小于350℃第二轻组分在小型催化裂化固定流化床试验装置进行催化裂化试验，所用催化剂为中石化催化剂有限公司长岭分公司生产的催化裂化催化剂MLC-500，反应温度为540℃，剂油比为6，停留时间为2s。
结果，所得产品汽油质量收率为42％，汽油RON辛烷值为92。
实施例I-8
采用与实施例I-1相似的工艺，不同之处在于，本实施例I-中将所得第二重组分引入至延迟焦化单元中进行反应，得到焦化汽油、焦化柴油和焦化蜡油。
延迟焦化单元的操作条件为：反应温度为510℃，停留时间为0.6h。
焦化柴油的硫含量0.26质量％，凝点-11℃，十六烷值48。
焦化蜡油的硫含量1.12质量％，凝点32℃。
焦化汽油的收率为14.7％，硫含量0.10质量％，MON为61.8。
并将焦化柴油和焦化蜡油循环回第三反应单元和所述LCO混合，以进行加氢处理，反应工艺条件同实施例I-1。
混合焦化柴油、焦化蜡油以及LCO的性质和第一重组分8的性质如表I-1所示。
DOA来自一种减压渣油，与第一重组分8按照质量比1:10混合，混合原料的性质见表I-2。
DOA和第一重组分8的混合原料先在溶氢单元(氢气的送入量与所述脱油沥青和所述第一重组分8的混合原料的体积比为100，溶氢单元操作温度为320℃，压力为8MPa)中与氢气混合，获得的混合物料进入第一反应单元，第一反应单元的操作条件为：反应温度为360℃，反应压力为8MPa，液时体积空速为0.3h -1，循环油：第一反应单元入口原料油体积比例0.5:1。混合原料加氢后产品性质见表I-3。
分馏第一反应单元得到的液相产品，大于等于350℃第二重组分性质见表I-4。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表I-5。
实施例I-9
将实施例I-1所得小于350℃第二轻组分在加氢裂化装置上进行试验，得到柴油组分。
操作条件为：反应温度为360℃，反应压力为10MPa，氢油体积比为1000，液时体积空速为1.0h -1。
结果：柴油组分硫含量5ppm，凝点-32℃，十六烷值53。
实施例I-10
采用与实施例I-1相似的工艺进行，所不同的是，本实施例I-中的第一反应单元中的催化剂装填情况如下：
按照物流方向，催化剂装填的顺序为加氢保护催化剂、富矿前驱体材料1、加氢脱金属脱硫催化剂、加氢脱硫催化剂。第一反应单元中，各催化剂之间的装填比为：RG-30B:富矿前驱体材料1：RDM-33B:RCS-31＝6：60：14：20(V/V)。
混合原料加氢后产品性质见表I-3。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表I-4。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表I-5。
实施例I-11
采用与实施例I-1相似的工艺进行，所不同的是，本实施例I-中的第一反应单元中的催化剂装填情况如下：
按照物流方向，催化剂装填的顺序为加氢保护催化剂、富矿前驱体材料2、富矿前驱体材料1、加氢脱金属脱硫催化剂、加氢脱硫催化剂。第一反应单元中，各催化剂之间的装填比为：RG-30B:富矿前驱体材料2：富矿前驱体材料1：RDM-33B:RCS-31＝6：30：30：14：20(V/V)。
混合原料加氢后产品性质见表I-3。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表I-4。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产 品，性质见表I-5。
实施例I-12
采用与实施例I-1相似的工艺进行，所不同的是，本实施例I-中的第一反应单元中的催化剂装填情况如下：
按照物流方向，催化剂装填的顺序为：加氢保护催化剂、加氢脱金属脱硫催化剂、加氢脱硫催化剂。第一反应单元中，各催化剂之间的装填比为：RG-30B：RDM-33B:RCS-31＝15:35:50(V/V)。
混合原料加氢后产品性质见表I-3。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表I-4。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表I-5。
实施例I-13
采用与实施例I-1相似的工艺进行，所不同的是，本实施例I-中的第一反应单元中的催化剂装填情况如下：
按照物流方向，催化剂装填的顺序为：加氢保护催化剂、富矿前驱体材料3、加氢脱金属脱硫催化剂、加氢脱硫催化剂。第一反应单元中，各催化剂之间的装填比为：RG-30B:富矿前驱体材料3：RDM-33B：RCS-31＝10：40：20：30(V/V)。
混合原料加氢后产品性质见表I-3。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表I-4。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表I-5。
对比例I-1
催化剂与装置与实施例I-1相似。所不同的是：
本对比例I-中富芳馏分油QY(芳烃含量为20质量％)不经过部分加氢饱和处理装置，而直接与DOA混合。DOA与QY以质量比1:10混合，混合原料的性质见表I-2。
与实施例I-1中相同，本对比例I-的混合原料先在溶氢单元中与氢气混合，获得的混合物料进入第一反应单元，经第一反应单元加氢处理后，产品性质见表I-3。
分馏第一反应单元加氢处理得到的液相产品，大于等于350℃第二重组分性质见表I-4。
小于350℃第二轻组分在固定床加氢裂化装置上进行试验，得到加氢裂化产品，性质见表I-5。
对比例I-2
催化剂与装置与实施例I-1相似。所不同的是：
本对比例I-中富芳馏分油QY不经过部分加氢饱和处理装置，而直接与DOA混合。DOA与QY以质量比2:10混合，混合原料的性质见表I-2。
与实施例I-1中相同，本对比例I-的混合原料先在溶氢单元中与氢气混合，获得的混合物料进入第一反应单元，经第一反应单元加氢处理后，产品性质见表I-3。
分馏第一反应单元加氢处理得到的液相产品，大于等于350℃第二重组分性质见表I-4。
小于350℃第二轻组分在固定床加氢裂化装置上进行试验，得到加氢裂化产品，性质见表I-5。
对比例I-3
催化剂与装置与实施例I-1相似。所不同的是：
本对比例I-中富芳馏分油QY不经过部分加氢饱和处理装置，而直接与DOA混合。DOA与QY以质量比3:10混合，因混合原料中有大量固体(100℃下)，故无法进行下一步试验。
表I-1：富芳馏分油加氢前后性质
Figure PCTCN2020125068-appb-000001
表I-2：混合原料性质
Figure PCTCN2020125068-appb-000002
表I-2(续表)：混合原料性质
Figure PCTCN2020125068-appb-000003
表I-3：液相加氢处理后产品性质
Figure PCTCN2020125068-appb-000004
表I-4：第二重组分性质
Figure PCTCN2020125068-appb-000005
表I-5：加氢裂化产品性质
项目 密度(20℃)，g/cm 3 RON 硫含量，μg/g
实施例I-1 0.72 ＞92 ＜10
实施例I-2 0.72 ＞92 ＜10
实施例I-3 0.72 ＞92 ＜10
实施例I-4 0.72 ＞92 ＜10
实施例I-8 0.72 ＞92 ＜10
实施例I-10 0.72 ＞92 ＜10
实施例I-11 0.72 ＞92 ＜10
实施例I-12 0.72 ＞92 ＜10
实施例I-13 0.72 ＞92 ＜10
对比例I-1 >0.72 <92 11
对比例I-2 >0.72 <92 11
表I-6：富矿前驱体材料性质
Figure PCTCN2020125068-appb-000006
实施例B
以一种减压渣油为原料进行溶剂脱沥青，所用溶剂为丁烷含量为70量％以上的烃类混合物，在120℃下，溶剂：减压渣油＝3：1(质量比)的条件下进行溶剂脱沥青，DAO质量收率70％，DOA质量收率30％。
所得DAO和DOA的性质见表II-1。
所得DAO和DOA的性质见表II-1。
实施例II-1
本实施例II-采用的DAO和DOA均来自实施例II-B。
DAO经第四反应单元中进行加氢反应后的液相产品性质见表II-1； 液相产品进入DCC单元进行反应，得到LCO1和HCO1。
LCO1在第三反应单元中进行加氢饱和后分馏以获得第一轻组分1和第一重组分1，第三反应单元加氢的操作条件为：反应温度为290℃，反应压力为4MPa，液时体积空速为1h -1，氢油体积比为800：1。LCO1和第一重组分1性质如表II-2所示。
DOA与第一重组分1按照质量比1:10混合，混合原料的性质见表II-3。
DOA和第一重组分1进入溶氢单元中与氢气混合，获得的混合物料(其中的氢含量见表II-3中)在第一反应单元的操作条件为：反应温度为360℃，反应压力为10MPa，液时体积空速为0.3h -1，循环油：第一反应单元入口原料油体积比例0.5:1。混合原料加氢后产品性质见表II-4。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表II-5。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表II-6。
实施例II-2
本实施例II-采用的DAO和DOA均来自实施例II-B。
DAO经第四反应单元中进行加氢反应后的液相产品性质见表II-1；液相产品进入DCC单元进行反应，得到LCO2和HCO2。
HCO2在第三反应单元中进行加氢饱和后分馏以获得第一轻组分2和第一重组分2，第三反应单元加氢操作条件为：反应温度为330℃，反应压力为6MPa，液时体积空速为1h -1，氢油体积比为800：1。HCO2和第一重组分2性质如表II-2所示。
DOA与第一重组分2按照质量比5:10混合，混合原料的性质见表II-3。
DOA和第一重组分2进入溶氢单元中与氢气混合，获得的混合物料(其中的氢含量见表II-3中)在第一反应单元的操作条件为：反应温度为380℃，反应压力为8MPa，液时体积空速为0.3h -1，循环油：第一反应单元入口原料油体积比例0.5:1。混合原料加氢后产品性质见表II-4。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组 分性质见表II-5。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表II-6。
实施例II-3
本实施例II-采用的DAO和DOA均来自实施例II-B。
DAO经第四反应单元中进行加氢反应后的液相产品性质见表II-1；液相产品进入DCC单元(操作条件同实施例II-1中)进行反应，得到LCO1和HCO1。
LCO1在第三反应单元中进行加氢饱和后分馏以获得第一轻组分3和第一重组分3，第三反应单元加氢操作条件为：反应温度为320℃，反应压力为6MPa，液时体积空速为1h -1，氢油体积比为800：1。LCO1和第一重组分3性质如表II-2所示。
DOA与第一重组分3按照质量比10:10混合，混合原料的性质见表II-3。
DOA和第一重组分3进入溶氢单元中与氢气混合，获得的混合物料(其中的氢含量见表II-3中)在第一反应单元的操作条件为：反应温度为370℃，反应压力为8MPa，液时体积空速为0.3h -1，循环油：第一反应单元入口原料油体积比例0.5:1。混合原料加氢后产品性质见表II-4。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表II-5。
将该第二重组分在反应温度为500℃停留时间为0.5小时下进行焦化反应，得到石油焦(收率31质量％)，硫含量为2.7质量％。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表II-6。
实施例II-4
本实施例II-采用的DAO和DOA均来自实施例II-B。
DAO经第四反应单元中进行加氢反应后的液相产品性质见表II-1；液相产品进入DCC单元(操作条件同实施例II-1中)进行反应，得到LCO1和HCO1。
本实施例II-采用的富芳馏分油为来自国内某煤焦化装置的煤焦油(性质见表II-1)和LCO1，LCO1与煤焦油的质量比为1:1，富芳馏分 油在第三反应单元中进行加氢饱和后分馏以获得第一轻组分4和第一重组分4，第三反应单元加氢操作条件为：反应温度为300℃，反应压力为10MPa，液时体积空速为0.8h -1，氢油体积比为800：1。富芳馏分油和第一重组分4的性质如表II-2所示。
DOA与第一重组分4按照质量比15:10混合，混合原料的性质见表II-3。
DOA和第一重组分4进入溶氢单元中与氢气混合，获得的混合物料(其中的氢含量见表II-3中)在第一反应单元的操作条件为：反应温度为350℃，反应压力为12MPa，液时体积空速为0.3h -1，循环油：第一反应单元入口原料油体积比例0.5:1。混合原料加氢后产品性质见表II-4。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表II-5。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表II-6。
实施例II-5
采用与实施例II-3相似的方法进行，所不同的是：
本实施例II-中，第一反应单元的加氢处理的温度为395℃。
其余条件与实施例II-3中相同。
混合原料加氢后产品性质见表II-4。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表II-5。
实施例II-6
催化剂装填和加氢处理的操作条件同实施例II-4。
与实施例II-4相同的混合原料经第一反应单元加氢处理后，每过30天，反应温度提3℃，加氢试验共计运行360天后停止运转。
初始装到反应器的富矿前驱体材料1和富矿前驱体材料2，反应后变成富V材料1和富钒材料2，经焙烧分析其V含量分别为69质量％和60质量％，是提炼高价值V 2O 5的高品质材料。
实施例II-7
将实施例II-3中的小于350℃第二轻组分在小型催化裂化固定流化床试验装置进行催化裂化试验，所用催化剂为中石化催化剂有限公司 长岭分公司生产的催化裂化催化剂MLC-500，反应温度为540℃，剂油比为6，停留时间为3s。
结果，产品汽油质量收率为40％，汽油RON辛烷值为93。
实施例II-8
采用与实施例II-1相似的工艺，不同之处在于，本实施例II-中将所得第二重组分引入至延迟焦化单元中进行反应，得到焦化汽油、焦化柴油和焦化蜡油。
延迟焦化单元的操作条件为：反应温度为510℃，停留时间为0.6h。
焦化柴油的硫含量0.26质量％，凝点-11℃，十六烷值48。
焦化蜡油的硫含量1.12质量％，凝点32℃。
焦化汽油的收率为14.7％，硫含量0.10质量％，MON为61.8。
并将焦化柴油和焦化蜡油循环回第三反应单元和所述LCO1混合，以进行加氢饱和后分馏以获得切割点为180℃的第一轻组分8和第一重组分8，反应工艺条件同实施例II-1。混合焦化柴油、焦化蜡油以及LCO1的油料的性质和第一重组分8的性质如表II-2所示。
DOA来自实施例II-B，与第一重组分8按照质量比1:10混合，混合原料的性质见表II-3。
DOA和第一重组分8进入溶氢单元中与氢气混合，获得的混合物料(其中的氢含量见表II-3中)在第一反应单元的操作条件为：反应温度为360℃，反应压力为8MPa，液时体积空速为0.3h -1，氢油体积比为800：1。混合原料加氢后产品性质见表II-4。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表II-5。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表II-6。
实施例II-9
将实施例II-1所得小于350℃第二轻组分在柴油加氢改质装置上进行试验，得到柴油组分。
柴油加氢改质装置的操作条件为：反应温度为360℃，反应压力为12MPa，氢油体积比为1000，液时体积空速为1.0h -1。
结果：所得柴油组分性质为硫含量5ppm，凝点-33℃，十六烷值53。
实施例II-10
采用与实施例II-1相似的工艺进行，所不同的是，本实施例II-中的第一反应单元中的催化剂装填情况如下：
按照物流方向，催化剂装填的顺序为加氢保护催化剂、富矿前驱体材料1、加氢脱金属脱硫催化剂、加氢脱硫催化剂。第一反应单元中，各催化剂之间的装填比为：RG-30B:富矿前驱体材料1：RDM-33B:RCS-31＝6：60：14：20(V/V)。
混合原料加氢后产品性质见表II-4。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表II-5。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表II-6。
实施例II-11
采用与实施例II-1相似的工艺进行，所不同的是，本实施例II-中的第一反应单元中的催化剂装填情况如下：
按照物流方向，催化剂装填的顺序为加氢保护催化剂、富矿前驱体材料2、富矿前驱体材料1、加氢脱金属脱硫催化剂、加氢脱硫催化剂。第一反应单元中，各催化剂之间的装填比为：RG-30B:富矿前驱体材料2：富矿前驱体材料1：RDM-33B：RCS-31＝6：30：30：14：20(V/V)。
混合原料加氢后产品性质见表II-4。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表II-5。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表II-6。
实施例II-12
采用与实施例II-1相似的工艺进行，所不同的是，本实施例II-中的第一反应单元中的催化剂装填情况如下：
按照物流方向，催化剂装填的顺序为：加氢保护催化剂、加氢脱金属脱硫催化剂、加氢脱硫催化剂。各催化剂之间的装填比为：RG-30B：RDM-33B:RCS-31＝15:40:45(V/V)。
混合原料加氢后产品性质见表II-4。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表II-5。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表II-6。
实施例II-13
采用与实施例II-1相似的工艺进行，所不同的是，本实施例II-中的第一反应单元中的催化剂装填情况如下：
按照物流方向，催化剂装填的顺序：加氢保护催化剂、富矿前驱体材料3、加氢脱金属脱硫催化剂、加氢脱硫催化剂。各催化剂之间的装填比为：RG-30B:富矿前驱体材料3：RDM-33B:RCS-31＝10：40：25：35(V/V)。
混合原料加氢后产品性质见表II-4。
分馏第一反应单元处理得到的液相产品，大于等于350℃第二重组分性质见表II-5。
小于350℃第二轻组分在第二反应单元进行试验，得到加氢裂化产品，性质见表II-6。
对比例II-1
催化剂与装置与实施例II-1相似。所不同的是：
本对比例II-中富芳馏分油QY(芳烃含量为20质量％)不经过部分加氢饱和处理装置，而直接与DOA混合。DOA与QY以质量比1:10混合，混合原料的性质见表II-3。
混合原料进入溶氢单元中与氢气混合，获得的混合原料(其中的氢含量见表II-3中)经第一反应单元加氢处理后，产品性质见表II-4。
分馏第一反应单元加氢处理得到的液相产品，大于等于350℃第二重组分性质见表II-5。
小于350℃第二轻组分在第二反应单元上进行试验，得到加氢裂化产品，性质见表II-6。
对比例II-2
催化剂与装置与实施例II-1相似。所不同的是：
本对比例II-中富芳馏分油QY不经过部分加氢饱和处理装置，而直接与DOA混合。DOA与QY以质量比2:10混合，混合原料的性质见表II-3。
混合原料进入溶氢单元中与氢气混合，获得的混合原料(其中的氢含量见表II-3中)经第一反应单元加氢处理后，产品性质见表II-4。
分馏第一反应单元加氢处理得到的液相产品，大于等于350℃第二重组分性质见表II-5。
小于350℃第二轻组分在第二反应单元上进行试验，得到加氢裂化产品，性质见表II-6。
对比例II-3
催化剂与装置与实施例II-1相似。所不同的是：
对比例II-3中富芳馏分油QY不经过部分加氢饱和处理装置，而直接与DOA混合。DOA与QY以质量比3:10混合，因混合原料中有大量固体(100℃下)，故无法进行下一步试验。
表II-1：DOA、DAO及第四反应单元加氢处理后液相产品等的性质
项目 DOA DAO 第四反应单元加氢处理后液相产品
密度(20℃)，g/cm 3 1153.6 981.1 939.1
残炭，质量％ 53.5 10.4 5.1
硫含量，质量％ 7.2 4.5 0.39
氮含量，质量％ 0.68 0.39 0.21
(Ni+V)，μg/g 390 39 6.8
表II-2：富芳馏分油加氢前后性质
Figure PCTCN2020125068-appb-000007
表II-3：混合原料性质
Figure PCTCN2020125068-appb-000008
表II-3(续表)：混合原料性质
Figure PCTCN2020125068-appb-000009
表II-4：混合原料加氢处理后产品性质
Figure PCTCN2020125068-appb-000010
表II-5：第二重组分性质
Figure PCTCN2020125068-appb-000011
表II-6：加氢裂化产品性质
项目 密度(20℃)，g/cm 3 RON 硫含量，μg/g
实施例II-1 0.72 ＞92 ＜10
实施例II-2 0.72 ＞92 ＜10
实施例II-3 0.72 ＞92 ＜10
实施例II-4 0.72 ＞92 ＜10
实施例II-8 0.72 ＞92 ＜10
实施例II-10 0.72 ＞92 ＜10
实施例II-11 0.72 ＞92 ＜10
实施例II-12 0.72 ＞92 ＜10
实施例II-13 0.72 ＞92 ＜10
对比例II-1 >0.72 <92 13
对比例II-2 >0.72 <92 12
表II-7：富矿前驱体材料性质
  灼减，质量％ 比表II-面积，m 2/g 吸水率，g/g
富矿前驱体材料1 13.5 263 1.08
富矿前驱体材料2 29.9 279 1.22
富矿前驱体材料3 20.5 99 1.05
由上述结果可以看出，本发明的技术能够从DOA得到优质的生产低硫船燃或低硫焦产品原料。
并且，本发明的技术能够得到优质的和符合国V标准的汽油产品。
以上详细描述了本发明的优选实施方式，但是，本发明并不限于此。在本发明的技术构思范围内，可以对本发明的技术方案进行多种简单变型，包括各个技术特征以任何其它的合适方式进行组合，这些简单变型和组合同样应当视为本发明所公开的内容，均属于本发明的保护范围。

Claims (26)

  1. 一种加工富芳馏分油的方法，其特征在于，该方法包括：
    (2)将脱油沥青和含有第一重组分的含芳烃物流引入至溶氢单元中与氢气混合，并将混合后的物料引入至第一反应单元中进行加氢反应，所述脱油沥青和所述含芳烃物流的用量比使得由该脱油沥青和含芳烃物流形成的混合原料在不高于400℃时呈液态；
    (3)将来自所述第一反应单元的液相产物进行分馏，得到第二轻组分和第二重组分，其中，所述第二轻组分和所述第二重组分的切割点为240～450℃；
    (41)将所述第二轻组分引入至第二反应单元中进行反应以得到选自汽油组分、柴油组分和BTX原料组分中的至少一种产物，其中，所述第二反应单元选自加氢裂化单元、催化裂化单元和柴油加氢提质单元中的至少一种；以及
    (42)将所述第二重组分引入至延迟焦化单元中进行反应以得到选自焦化汽油、焦化柴油、焦化蜡油和低硫石油焦中的至少一种产物；或者将所述第二重组分作为低硫船用燃料油组分。
  2. 根据权利要求1所述的方法，该方法还包括：
    (1)将富芳馏分油引入至第三反应单元中进行加氢饱和后分馏以获得第一轻组分和第一重组分，所述第一轻组分和所述第一重组分的切割点为100-250℃，所述第一重组分中的芳烃含量为大于等于20质量％；其中，所述第一重组分用作步骤(2)中所述含芳烃物流所含有的第一重组分。
  3. 根据权利要求1所述的方法，其中，在步骤(2)中，所述第一反应单元中含有富矿前驱体材料和/或加氢催化剂，所述第一反应单元为液相加氢反应单元，所述富矿前驱体材料为能够吸附选自V、Ni、Fe、Ca和Mg中的至少一种金属的材料。
  4. 根据权利要求1所述的方法，其中，在步骤(2)中，所述脱油沥青和所述含芳烃物流的用量比使得由该脱油沥青和含芳烃物流形成的混合原料的100℃粘度不大于400mm 2/s，优选不大于200mm 2/s，更优选不大于100mm 2/s。
  5. 根据权利要求1所述的方法，其中，在步骤(2)中，所述含芳 烃物流中还含有芳烃化合物和/或芳烃油，所述芳烃油选自LCO、HCO、FGO、乙烯焦油、煤焦油、焦化柴油和焦化蜡油中的至少一种；
    优选地，所述芳烃化合物选自苯、甲苯、二甲苯、萘、由至少一种C 1-6的烷基取代的萘、三环以上芳烃中的至少一种。
  6. 根据权利要求1所述的方法，其中，所述富芳馏分油中的芳烃含量大于等于20质量％，优选大于等于25质量％，更优选大于等于40质量％。
  7. 根据权利要求1所述的方法，其中，在步骤(2)中，所述脱油沥青为由重油原料进入溶剂脱沥青单元中进行溶剂脱沥青处理后得到的脱油沥青；
    优选地，在所述溶剂脱沥青单元中，所述脱油沥青的收率质量分数不大于50％，优选不大于40％，更优选不大于30％。
  8. 根据权利要求1所述的方法，其中，在步骤(2)中，所述脱油沥青与所述含芳烃物流的用量质量比为1:10～50:10，优选为2:10～30:10；更优选为3:10～15:10。
  9. 根据权利要求1所述的方法，其中，该方法还包括：将步骤(42)中获得的所述焦化柴油和/或所述焦化蜡油循环回步骤(1)中的所述第一反应单元进行加氢饱和。
  10. 根据权利要求1所述的方法，其中，在步骤(1)中，所述第三反应单元为固定床反应器、移动床反应器和沸腾床反应器中的至少一种反应器；
    优选地，所述第三反应单元中的操作条件包括：反应温度为200-420℃，反应压力为2-18MPa，液时体积空速为0.3-10h -1，氢油体积比50-5000；
    优选地，所述第三反应单元中的操作条件包括：反应温度为220-400℃，反应压力为2-15MPa，液时体积空速为0.3-5h -1，氢油体积比为50-4000。
  11. 根据权利要求1所述的方法，其中，在步骤(2)中，所述第一反应单元中的操作条件包括：反应温度260～500℃，反应压力为2.0～20.0MPa，循环油与所述第一反应单元入口原料油的体积比例为0.1:1至15:1，液时体积空速为0.1～1.5h -1。
  12. 根据权利要求1所述的方法，其中，在步骤(2)中，所述富 矿前驱体材料的灼减不低于3质量％，比表面积不低于80m 2/g，吸水率不低于0.9g/g。
  13. 根据权利要求12所述的方法，其中，在步骤(2)中，所述富矿前驱体材料中含有载体和负载在所述载体上的活性组分元素，所述载体选自氢氧化铝、氧化铝和氧化硅中的至少一种，所述活性组分元素选自第VIB族和VIII族金属元素中的至少一种。
  14. 根据权利要求13所述的方法，其中，在步骤(2)中，按照反应物流方向，所述第一反应单元中依次装填有第一富矿前驱体材料和第二富矿前驱体材料，且所述第二富矿前驱体材料的灼减大于等于所述第一富矿前驱体材料的灼减；
    优选地，所述第一富矿前驱体材料的灼减为3-15质量％，以及所述第二富矿前驱体材料的灼减为不小于15质量％；
    优选地，所述第一富矿前驱体材料与所述第二富矿前驱体材料的装填体积比为5:95～95:5。
  15. 根据权利要求1所述的方法，其中，在步骤(41)中，所述第二反应单元为加氢裂化单元，且所述加氢裂化单元中的操作条件包括：反应温度为360～420℃，反应压力为10.0～18.0MPa，氢油体积比为600～2000，液时体积空速为1.0～3.0h -1；
    优选地，所述加氢裂化单元中装填有至少一种加氢处理催化剂和至少一种加氢裂化催化剂。
  16. 根据权利要求1所述的方法，其中，在步骤(41)中，所述第二反应单元为催化裂化单元，且所述催化裂化单元为流化催化裂化单元；
    优选地，所述流化催化裂化单元中的操作条件包括：反应温度为500～600℃，剂油比为3～12，停留时间为0.6～6s。
  17. 根据权利要求1所述的方法，其中，在步骤(41)中，所述第二反应单元为柴油加氢提质单元，且所述柴油加氢提质单元中的操作条件包括：反应温度为330～420℃，反应压力为5.0～18.0MPa，氢油体积比为500～2000，液时体积空速为0.3～3.0h -1；
    优选地，所述柴油加氢提质单元中装填有至少一种柴油加氢提质催化剂。
  18. 根据权利要求1所述的方法，其中，在步骤(42)中，将所述 第二重组分引入至延迟焦化单元中进行反应以得到选自焦化汽油、焦化柴油、焦化蜡油和低硫石油焦中的至少一种产物，且所述至延迟焦化单元中的操作条件包括：反应温度为440～520℃，停留时间为0.1～4h；
    优选地，在步骤(42)中，所述第二重组分的硫含量不大于1.8质量％，将所述第二重组分引入至延迟焦化单元中进行反应以得到低硫石油焦，优选所述低硫石油焦的硫含量不大于3质量％。
  19. 根据权利要求1所述的方法，其中，在步骤(42)中，将所述第二重组分作为低硫船用燃料油组分，且控制条件使得所述低硫船用燃料油组分中的硫含量不大于0.5质量％。
  20. 根据权利要求2所述的方法，该方法还包括：
    (11)将重质原料油引入至溶剂脱沥青单元中进行溶剂脱沥青处理，得到脱油沥青和脱沥青油；
    (12)将所述脱沥青油引入至第四反应单元中进行加氢反应，并将所述第四反应单元中获得的液相流出物引入至DCC单元进行反应，得到丙烯、LCO、HCO和油浆，其中，所述第四反应单元为固定床反应单元；
    将含有来自所述DCC单元的LCO和/或HCO的富芳馏分油用作所述步骤(1)中富芳馏分油。
  21. 根据权利要求20所述的方法，其中，该方法还包括：将步骤(42)中获得的所述焦化柴油和/或所述焦化蜡油循环回所述第三反应单元中进行加氢饱和。
  22. 根据权利要求20所述的方法，其中，在步骤(12)中，所述第四反应单元的操作条件包括：反应温度为280～400℃，反应压力为6.0～14.0MPa，氢油体积比为600～1200，液时体积空速为0.3～2.0h -1；
    优选地，在步骤(12)中，所述第四反应单元中装填有至少两种加氢催化剂；
    优选地，在步骤(12)中，所述加氢催化剂为能够催化选自加氢脱金属反应、加氢脱硫反应和加氢脱残炭反应中的至少一种反应的催化剂；
    优选地，在步骤(12)中，所述加氢催化剂中含有作为载体的氧化铝和作为活性组分元素的第VIB族和/或VIII族金属元素，且该加氢催化剂中任选还含有选自P、Si、F和B中的至少一种助剂元素。
  23. 一种加工富芳馏分油的***，其特征在于，该***中包括：
    第三反应单元，该第三反应单元用于将富芳馏分油在其中进行加氢饱和和分馏以得到第一轻组分和第一重组分；
    溶氢单元，该溶氢单元与所述第三反应单元保持流体连通，用于将脱油沥青和含有来自所述第三反应单元的第一重组分的含芳烃物流在其中与氢气混合；
    第一反应单元，该第一反应单元为液相加氢反应单元且与所述溶氢单元保持流体连通，用于将所述溶氢单元的混合物料在其中进行加氢反应；
    分离单元，该分离单元与所述第一反应单元保持流体连通，用于将来自所述第一反应单元的液相产物在其中进行分馏；
    第二反应单元，该第二反应单元与所述分离单元保持流体连通，用于将由所述分离单元中获得的第二轻组分在其中进行反应，所述第二反应单元选自加氢裂化单元、催化裂化单元和柴油加氢提质单元中的至少一种；
    延迟焦化单元，该延迟焦化单元与所述分离单元保持流体连通，用于将由所述分离单元中获得的第二重组分在其中进行反应以得到选自焦化汽油、焦化柴油、焦化蜡油和低硫石油焦中的至少一种产物；
    出口，该出口与所述分离单元保持流体连通，用于将由所述分离单元中获得的第二重组分作为低硫船用燃料油组分引出***。
  24. 根据权利要求23所述的***，其中，所述延迟焦化单元与所述溶氢单元保持流体连通，用于将所述延迟焦化单元中获得的所述焦化柴油和/或所述焦化蜡油循环回所述第一反应单元中。
  25. 根据权利要求23所述的***，其中，该***中还包括溶剂脱沥青单元，该溶剂脱沥青单元与所述溶氢单元保持流体连通，用于将重油原料在其中进行溶剂脱沥青处理，并将所述溶剂脱沥青处理后得到的脱油沥青引入至所述溶氢单元中。
  26. 根据权利要求23所述的***，该***还包括：
    溶剂脱沥青单元，该溶剂脱沥青单元用于将重质原料油在其中进行溶剂脱沥青处理，得到脱油沥青和脱沥青油；
    第四反应单元，该第四反应单元与所述溶剂脱沥青单元保持流体连通，且该第四反应单元为固定床反应单元，用于将来自所述溶剂脱 沥青单元的脱沥青油在其中进行加氢反应；
    DCC单元，该DCC单元与所述第四反应单元保持流体连通，用于将所述第四反应单元中获得的液相流出物在其中进行反应以得到丙烯、LCO、HCO和油浆；
    其中该DCC单元与所述第三反应单元保持流体连通，以将含有来自所述DCC单元的LCO和/或HCO的富芳馏分油输送至所述第三反应单元中用作所述富芳馏分油。
PCT/CN2020/125068 2019-10-31 2020-10-30 一种加工富芳馏分油的方法和*** WO2021083302A1 (zh)

Priority Applications (3)

Application Number Priority Date Filing Date Title
KR1020227017073A KR20220091510A (ko) 2019-10-31 2020-10-30 방향족이 풍부한 증류유 가공 방법 및 시스템
US17/772,317 US20220403263A1 (en) 2019-10-31 2020-10-30 Process and system for processing aromatics-rich fraction oil
JP2022525049A JP2023501181A (ja) 2019-10-31 2020-10-30 芳香族リッチ留分油を加工するための方法およびシステム

Applications Claiming Priority (4)

Application Number Priority Date Filing Date Title
CN201911054674.9 2019-10-31
CN201911054674.9A CN112745952B (zh) 2019-10-31 2019-10-31 一种加工富芳馏分油的方法和***
CN201911053864.9 2019-10-31
CN201911053864.9A CN112745949B (zh) 2019-10-31 2019-10-31 一种联合加工脱油沥青和富芳馏分油的方法和***

Publications (1)

Publication Number Publication Date
WO2021083302A1 true WO2021083302A1 (zh) 2021-05-06

Family

ID=75714901

Family Applications (1)

Application Number Title Priority Date Filing Date
PCT/CN2020/125068 WO2021083302A1 (zh) 2019-10-31 2020-10-30 一种加工富芳馏分油的方法和***

Country Status (5)

Country Link
US (1) US20220403263A1 (zh)
JP (1) JP2023501181A (zh)
KR (1) KR20220091510A (zh)
TW (1) TW202136482A (zh)
WO (1) WO2021083302A1 (zh)

Cited By (1)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN116024014A (zh) * 2021-10-27 2023-04-28 中国石油化工股份有限公司 两个加氢裂化***联合的方法

Families Citing this family (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN114958419B (zh) * 2021-02-18 2023-08-08 中国石油化工股份有限公司 一种加工催化柴油的方法
CN117844525A (zh) * 2024-03-07 2024-04-09 陕西煤业化工集团神木天元化工有限公司 一种中温煤焦油制化学品和特种燃料的方法

Citations (8)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4213846A (en) * 1978-07-17 1980-07-22 Conoco, Inc. Delayed coking process with hydrotreated recycle
CN1654603A (zh) * 2004-02-13 2005-08-17 中国石油化工股份有限公司 一种劣质重、渣油的转化方法
CN1766059A (zh) * 2004-10-29 2006-05-03 中国石油化工股份有限公司 一种劣质重、渣油的处理方法
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
CN104093818A (zh) * 2012-01-27 2014-10-08 沙特***石油公司 用于直接加工原油的整合的溶剂脱沥青、加氢处理以及水蒸气热解方法
CN104232158A (zh) * 2014-08-22 2014-12-24 中国石油大学 沥青质轻质化方法
CN105567316A (zh) * 2015-12-23 2016-05-11 上海新佑能源科技有限公司 劣质重油加工处理方法
CN105623725A (zh) * 2014-10-27 2016-06-01 中国石油化工股份有限公司 一种重/渣油加工的组合工艺

Patent Citations (8)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4213846A (en) * 1978-07-17 1980-07-22 Conoco, Inc. Delayed coking process with hydrotreated recycle
CN1654603A (zh) * 2004-02-13 2005-08-17 中国石油化工股份有限公司 一种劣质重、渣油的转化方法
CN1766059A (zh) * 2004-10-29 2006-05-03 中国石油化工股份有限公司 一种劣质重、渣油的处理方法
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
CN104093818A (zh) * 2012-01-27 2014-10-08 沙特***石油公司 用于直接加工原油的整合的溶剂脱沥青、加氢处理以及水蒸气热解方法
CN104232158A (zh) * 2014-08-22 2014-12-24 中国石油大学 沥青质轻质化方法
CN105623725A (zh) * 2014-10-27 2016-06-01 中国石油化工股份有限公司 一种重/渣油加工的组合工艺
CN105567316A (zh) * 2015-12-23 2016-05-11 上海新佑能源科技有限公司 劣质重油加工处理方法

Cited By (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN116024014A (zh) * 2021-10-27 2023-04-28 中国石油化工股份有限公司 两个加氢裂化***联合的方法
CN116024014B (zh) * 2021-10-27 2024-06-11 中国石油化工股份有限公司 两个加氢裂化***联合的方法

Also Published As

Publication number Publication date
US20220403263A1 (en) 2022-12-22
KR20220091510A (ko) 2022-06-30
TW202136482A (zh) 2021-10-01
JP2023501181A (ja) 2023-01-18

Similar Documents

Publication Publication Date Title
CN101210200B (zh) 一种渣油加氢处理与催化裂化组合工艺方法
WO2021083302A1 (zh) 一种加工富芳馏分油的方法和***
CN101418222A (zh) 一种处理劣质渣油的组合工艺
CA2652227C (en) Improved hydrocracker post-treat catalyst for production of low sulfur fuels
CN103102944A (zh) 一种渣油加氢处理及轻质化的组合工艺方法
CN112538384B (zh) 一种多产异丁烷和轻质芳烃的加氢处理-催化裂化组合工艺方法
CN112745952B (zh) 一种加工富芳馏分油的方法和***
CN114437786B (zh) 一种劣质原料油的加氢裂化方法
CN112745949B (zh) 一种联合加工脱油沥青和富芳馏分油的方法和***
CN112745948B (zh) 一种加工重质原料油和富芳馏分油的方法和***
CN112745951B (zh) 一种加工富芳馏分油的方法和***
WO2021083305A1 (zh) 一种加氢处理脱油沥青的方法和***
CN112745947B (zh) 一种加工重质原料油的方法和***
CN112745946B (zh) 一种加工重质原料油的方法和***
CN112745950B (zh) 一种加氢处理脱油沥青的方法和***
CN114437795B (zh) 一种加工重油的方法和***
CN114437808B (zh) 一种加工重油的方法和***
CN112745953B (zh) 一种加氢处理脱油沥青的方法和***
CN114426887B (zh) 一种加工富芳馏分油的方法
CN110408430B (zh) 一种组合工艺处理重烃的方法
CN103102985B (zh) 一种渣油加氢处理与催化裂化组合工艺方法
US4210525A (en) Hydrodenitrogenation of demetallized residual oil

Legal Events

Date Code Title Description
121 Ep: the epo has been informed by wipo that ep was designated in this application

Ref document number: 20881495

Country of ref document: EP

Kind code of ref document: A1

ENP Entry into the national phase

Ref document number: 2022525049

Country of ref document: JP

Kind code of ref document: A

ENP Entry into the national phase

Ref document number: 20227017073

Country of ref document: KR

Kind code of ref document: A

NENP Non-entry into the national phase

Ref country code: DE

122 Ep: pct application non-entry in european phase

Ref document number: 20881495

Country of ref document: EP

Kind code of ref document: A1

122 Ep: pct application non-entry in european phase

Ref document number: 20881495

Country of ref document: EP

Kind code of ref document: A1

WWE Wipo information: entry into national phase

Ref document number: 522432414

Country of ref document: SA