WO2023040890A1 - 一种加氢裂化的方法和系统 - Google Patents

一种加氢裂化的方法和系统 Download PDF

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Publication number
WO2023040890A1
WO2023040890A1 PCT/CN2022/118720 CN2022118720W WO2023040890A1 WO 2023040890 A1 WO2023040890 A1 WO 2023040890A1 CN 2022118720 W CN2022118720 W CN 2022118720W WO 2023040890 A1 WO2023040890 A1 WO 2023040890A1
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Prior art keywords
fraction
hydrocracking
weight
oil
unit
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PCT/CN2022/118720
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English (en)
French (fr)
Inventor
胡志海
莫昌艺
任亮
毛以朝
庄立
蔡新恒
赵毅
赵广乐
严张艳
赵阳
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Sinopec Research Institute of Petroleum Processing
China Petroleum and Chemical Corp
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Sinopec Research Institute of Petroleum Processing
China Petroleum and Chemical Corp
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Application filed by Sinopec Research Institute of Petroleum Processing , China Petroleum and Chemical Corp filed Critical Sinopec Research Institute of Petroleum Processing
Priority to US18/692,008 priority Critical patent/US20240384186A1/en
Priority to JP2024516631A priority patent/JP2024531744A/ja
Priority to EP22869259.6A priority patent/EP4403614A4/en
Priority to KR1020247012312A priority patent/KR20240140888A/ko
Publication of WO2023040890A1 publication Critical patent/WO2023040890A1/zh
Anticipated expiration legal-status Critical
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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
    • C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
    • C10G69/04—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one step of catalytic cracking in the absence of hydrogen
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    • C10G65/00—Treatment of hydrocarbon oils by two or more hydrotreatment processes only
    • C10G65/02—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only
    • C10G65/12—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only including cracking steps and other hydrotreatment steps
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    • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
    • C10G45/02—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds to eliminate hetero atoms without changing the skeleton of the hydrocarbon involved and without cracking into lower boiling hydrocarbons; Hydrofinishing
    • C10G45/04—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds to eliminate hetero atoms without changing the skeleton of the hydrocarbon involved and without cracking into lower boiling hydrocarbons; Hydrofinishing characterised by the catalyst used
    • C10G45/06—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds to eliminate hetero atoms without changing the skeleton of the hydrocarbon involved and without cracking into lower boiling hydrocarbons; Hydrofinishing characterised by the catalyst used containing nickel or cobalt metal, or compounds thereof
    • C10G45/08—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds to eliminate hetero atoms without changing the skeleton of the hydrocarbon involved and without cracking into lower boiling hydrocarbons; Hydrofinishing characterised by the catalyst used containing nickel or cobalt metal, or compounds thereof in combination with chromium, molybdenum, or tungsten metals, or compounds thereof
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    • C10G47/04—Oxides
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    • C10G47/02—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions characterised by the catalyst used
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    • C10G47/00—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions
    • C10G47/02—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions characterised by the catalyst used
    • C10G47/10—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions characterised by the catalyst used with catalysts deposited on a carrier
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    • C10G47/02—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions characterised by the catalyst used
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    • C10G47/20—Crystalline alumino-silicate carriers the catalyst containing other metals or compounds thereof
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    • C10G47/00—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions
    • C10G47/36—Controlling or regulating
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    • 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
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
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    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
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    • C10G9/00—Thermal non-catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G9/34—Thermal non-catalytic cracking, in the absence of hydrogen, of hydrocarbon oils by direct contact with inert preheated fluids, e.g. with molten metals or salts
    • C10G9/36—Thermal non-catalytic cracking, in the absence of hydrogen, of hydrocarbon oils by direct contact with inert preheated fluids, e.g. with molten metals or salts with heated gases or vapours
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    • C10G2300/1029—Gas hydrates
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    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/107—Atmospheric residues having a boiling point of at least about 538 °C
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    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1074—Vacuum distillates
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    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/30—Physical properties of feedstocks or products
    • C10G2300/301—Boiling range
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    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/30—Aromatics

Definitions

  • the invention relates to the technical field of processing hydrocarbon raw materials, in particular to a method and system for hydrocracking wax oil raw materials.
  • hydrocracking technology has the advantages of strong adaptability to raw materials, flexible production operations and product plans, and excellent product quality.
  • One of the important processing techniques for quality adjustment and oil-chemical combination is important.
  • the feedstock oil for hydrocracking is usually wax oil, and the wax oil feedstock is composed of paraffin, cycloalkane and aromatic hydrocarbon molecules, and the carbon number range is about 20-40.
  • conventional hydrocracking catalysts mainly use Y-type or ⁇ -type molecular sieves as catalytic materials, and use the acidic function of the catalytic materials to carry out chain scission reactions.
  • CN87105808A discloses an improved method for hydrodewaxing hydrocracking lubricating oil base stock.
  • hydrocracking or solvent dewaxing lubricating oil base oil is passed through a catalyst bed with dewaxing activity and a hydrogenation supplementary refining catalyst bed in sequence. layer, thereby producing a lubricating base oil product with a reduced cloud point.
  • CN102959054A discloses a method for combined hydrocracking and dewaxing of hydrocarbons.
  • raw oil is subjected to hydroprocessing and reaction in the first hydrocracking reaction zone in sequence to obtain the first hydrocracking reaction effluent and enter the first catalytic dewaxing process.
  • the reaction is carried out in the wax reaction zone, and the reaction effluent is separated and fractionated to obtain a naphtha fraction, a first diesel fraction and a bottom product fraction, wherein the bottom product fraction is reacted in the second hydrocracking or second catalytic dewaxing reaction zone, and the reaction
  • the effluent is separated and fractionated to form a second diesel fraction and a lube product fraction.
  • CN102311785A discloses a method for producing lubricating oil base oil by hydrogenation of naphthenic distillate oil.
  • the method takes naphthenic feedstock oil as raw material, adopts hydrogenation catalyst containing ⁇ -type molecular sieve, and hydrogenation reducing agent containing ZSM-5 type molecular sieve.
  • the method of coagulation catalyst and hydrogenation supplementary refining is used to produce rubber extender oil products with reduced pour point.
  • CN102971401B discloses a method for combined hydrocracking and dewaxing of hydrocarbons.
  • raw oil is first subjected to hydrotreating, and the hydrotreated product is separated to obtain a liquid phase residue for catalytic dewaxing and hydrocracking reactions.
  • the reaction flows out Diesel product fractions and lubricating oil base oil product fractions are obtained after separation and fractionation.
  • CN106609803A discloses a catalyst for producing high viscosity index hydrocracking tail oil and its preparation method.
  • the method prepares the catalyst by mixing macroporous alumina, modified USY molecules and modified ZSM-48 molecular sieves, and uses the catalyst The raw material undergoes hydrogenation ring-opening and hydroisomerization reactions to produce lubricating base oil products with low quality alkane content, high isomeric hydrocarbon content and high annual index.
  • conventional hydrocracking technology mainly adopts a hydrocracking catalyst containing Y-type molecular sieves to convert wax oil feedstock oil into
  • the corresponding cracking reaction cannot occur according to the molecular structure composition, and the efficient conversion of the hydrocarbon molecules of the wax oil raw material according to the type of hydrocarbon molecular structure cannot be realized, and the product quality and added value are low.
  • the present invention aims to solve the problems of low added value of products and low utilization efficiency of wax oil raw material molecules caused by indiscriminate transformation of wax oil raw material molecular structure existing in the existing hydrocracking technology.
  • a first aspect of the present invention provides a method for hydrocracking, comprising:
  • Hydroprocessing unit the mixed material of wax oil raw material oil and hydrogen is contacted with hydrogenation protecting agent, optional hydrodemetallization catalyst, hydrorefining catalyst in order to react in hydroprocessing unit, obtain reaction effluent ;
  • step (1) the reaction effluent obtained in step (1) enters the first hydrocracking unit, and reacts with the hydrocracking catalyst 1 in the presence of hydrogen, and the gained reaction effluent after separation, at least Obtain light fraction I and heavy fraction I; Rich in paraffins in the light fraction I, the mass fraction of paraffins in the light fraction I is at least 82%, rich in naphthenes and aromatics in the heavy fraction I, in the heavy fraction I In the hydrocarbon composition of the >350°C fraction in Fraction I, the sum of the mass fractions of naphthenes and aromatics is greater than 82%;
  • step (3) The second hydrocracking unit, the heavy fraction I obtained in step (2) enters the second hydrocracking unit, and reacts in contact with the hydrocracking catalyst II and/or the hydrotreating catalyst in the presence of hydrogen, and the obtained reaction After the effluent is separated, at least a light fraction II and a heavy fraction II are obtained.
  • the wax oil raw material oil has an initial boiling point of 300-350°C, which is selected from normal pressure wax oil, vacuum wax oil, hydrogenated wax oil, coking wax oil, catalytic cracking heavy cycle oil, and deasphalted oil. one or more of.
  • the present invention provides a hydrocracking method based on the structural characteristics of hydrocarbon molecules.
  • the wax oil raw material oil and the hydrogen mixture first pass through the hydroprocessing unit After the contact reaction, the reaction effluent enters the first hydrocracking unit and reacts with the hydrocracking catalyst I to realize the selective conversion of the mid-chain structure of the wax oil raw material oil to obtain light fraction I rich in paraffins and rich in cyclic hydrocarbons
  • the heavy fraction I of (cycloalkanes and aromatics) after being mixed with hydrogen, the heavy fraction I enters the second hydrocracking reaction unit to react with hydrocracking catalyst II and/or hydrotreating catalyst, thereby obtaining rich cycloalkanes and aromatics
  • the present invention realizes the selective and high-efficiency conversion of the wax oil raw material oil according to the hydrocarbon molecular chain structure and the ring structure type on the
  • the reaction effluent has different cutting schemes.
  • the reaction effluent obtained from the first hydrocracking unit is separated to obtain light fraction I and heavy fraction I, light fraction I
  • the initial boiling point is 20°C-30°C
  • the cut point of the light fraction I and the heavy fraction I is 65°C-120°C, preferably 65-105°C
  • the light fraction I is rich in paraffins, preferably in the light fraction I
  • the mass fraction of paraffins in I is at least 85%.
  • the light fraction I rich in paraffins can be used as raw material for high-quality steam cracking ethylene plant.
  • the obtained heavy fraction I is rich in naphthenes and aromatics, and in the hydrocarbon composition of the >350°C fraction in the heavy fraction I, the sum of the mass fractions of naphthenes and aromatics is greater than 82%.
  • the reaction effluent obtained from the first hydrocracking unit is separated to obtain light fraction I, middle fraction I and heavy fraction I, and the initial boiling point of light fraction I is 20°C-30°C , the cut point of the light fraction I and the middle fraction I is 65°C-120°C, preferably 65-105°C, and the cut point of the middle fraction I and the heavy fraction I is 160-180°C.
  • the light fraction I is rich in paraffins, preferably the mass fraction of paraffins in the light fraction I is at least 85%.
  • the middle fraction I can be used as a product alone, or can be sent to the fractionating tower of the second hydrocracking unit for further cutting to obtain part of the light fraction II component and the heavy fraction II component.
  • the obtained heavy fraction I is rich in naphthenes and/or aromatics, and in the hydrocarbon composition of >350° C. cuts in the heavy fraction I, the sum of the mass fractions of naphthenes and aromatics is greater than 82%.
  • the present invention sends the heavy fraction I to the second hydrocracking unit for selective cracking reaction, and the obtained reaction effluent is separated to obtain light fraction II and heavy fraction II.
  • the initial boiling point of the obtained light fraction II is 65°C-100°C
  • the cut point of the light fraction II and the heavy fraction II is 155-180°C, preferably 160-175°C.
  • the sum of the mass fractions of naphthenes and aromatics in the light fraction II is at least 58%, which is a high-quality reformate. According to different product schemes, the resulting heavy fraction II has different cutting schemes.
  • the heavy fraction II can be cut into various naphthenic special oils such as jet fuel fractions with high specific gravity, transformer oil base oils, and refrigeration oils. Taste.
  • the mass fraction of naphthenes in the >350°C distillate in the obtained heavy fraction II is at least 50%.
  • the heavy distillate II rich in naphthenes has good low-temperature fluidity, and the heavy distillate II can be used as various high value-added naphthenic special oil products.
  • the loading volume fractions of the hydroprotecting agent, the optional hydrodemetallization catalyst, and the hydrorefining catalyst are respectively For: 3%-10%; 0%-20%; 70%-90%.
  • the hydroprotectant is a conventional hydroprotectant for heavy hydrocarbon oil processing in the field, not limited to a wax oil hydroprotectant, a residual oil hydroprotectant or a graded combination thereof.
  • the hydrogenation protecting agent contains a carrier and an active metal component loaded on the carrier, the carrier is selected from one or more of alumina, silicon oxide and titanium oxide, and the active metal component is selected from From one or more of Group VIB metals and Group VIII non-noble metals, based on the weight of the hydrogenation protection agent, in terms of oxides, the active metal component is 0.1-15% by weight, the hydrogenation protection agent
  • the particle size is 0.5-50.0mm
  • the bulk density is 0.3-1.2g/cm 3
  • the specific surface area is 50-300m 2 /g.
  • the hydrodemetallization catalyst is a conventional hydrodemetallization catalyst for heavy hydrocarbon oil processing in the field, not limited to a wax oil hydrodemetallization catalyst, a residual oil hydrodemetallization catalyst or a graded combination thereof.
  • the hydrodemetallization catalyst contains a carrier and an active metal component loaded on the carrier, the carrier is selected from one or more of alumina, silicon oxide and titanium oxide, and the active metal component is One or more selected from Group VIB metals and Group VIII non-noble metals, based on the weight of the hydrodemetallization catalyst, in terms of oxides, the active metal component is 3-30% by weight.
  • the particle size of the metal catalyst is 0.2-2.0 mm, the bulk density is 0.3-0.8 g/cm 3 , and the specific surface area is 100-250 m 2 /g.
  • step, catalyst or component is optional, but not essential, that is, the step, catalyst or component may or may not exist.
  • the hydrorefining catalyst is a supported catalyst
  • the carrier is alumina and/or silica-alumina
  • the active metal component is at least one selected from group VIB metals and/or Or at least one metal selected from Group VIII; said Group VIII metal selected from nickel and/or cobalt, said Group VIB metal selected from molybdenum/or tungsten, based on the total weight of the hydrotreating catalyst, in In terms of oxides, the content of Group VIII metals is 1-15% by weight, and the content of Group VIB metals is 5-40% by weight,
  • the active metal component of the hydrorefining catalyst is selected from two or three of nickel, molybdenum and tungsten metals.
  • the reaction conditions of the hydroprocessing unit are: the hydrogen partial pressure is 3.0MPa-20.0MPa, the reaction temperature is 280°C-400°C, and the liquid hourly volume space velocity is 0.5h - 1-6h -1 , the volume ratio of hydrogen to oil is 300-2000.
  • the aromatic hydrocarbon saturation rate of the raw oil is controlled to be less than or equal to 58%.
  • the inventors of the present invention have conducted in-depth research and found that if the saturation ratio of aromatics is too high, when the reaction effluent of the hydroprocessing unit enters the first hydrocracking unit, it will lead to an increase in the ring-opening cracking reaction of naphthenes in the first hydrocracking reaction unit , which adversely affects the effect of the directional conversion reaction of the wax oil raw material oil according to the chain structure and the ring structure.
  • Raw material aromatics saturation rate 100%*(raw material aromatics content-aromatics content in the effluent of the hydrotreating reaction unit)/raw material aromatics content.
  • the reaction conditions of the first hydrocracking unit are: the hydrogen partial pressure is 3.0MPa-20.0MPa, the reaction temperature is 280°C-400°C, and the liquid hourly volume space velocity is 0.5h -1 -6h -1 , the volume ratio of hydrogen to oil is 300-2000.
  • the >350°C distillate conversion rate control range of the first hydrocracking reaction unit for:
  • A is the mass fraction of paraffins in the wax oil feedstock oil
  • B is the sum of the mass fractions of paraffins, cycloalkanes, and single-ring aromatics in the wax oil feedstock oil
  • the conversion rate of the >350°C fraction of the first hydrocracking reaction unit 100%* (mass fraction of >350°C fraction in the wax oil raw material - mass fraction of >350°C fraction in the reaction product of the first hydrocracking reaction unit) /Mass fraction of >350°C distillate in the wax oil raw material oil.
  • the first hydrocracking reaction unit by adjusting the first One or more process condition parameters in the reaction temperature, volume space velocity, hydrogen-to-oil ratio and reaction pressure of the hydrocracking reaction unit make the paraffin conversion rate of the feedstock oil 56%-95%, the sum of naphthenes and aromatics
  • the conversion rate is 10%-65%, preferably through the reaction temperature and volume space velocity to regulate the conversion rate of paraffins, and the conversion rate of naphthenes and aromatics
  • the paraffin conversion rate (paraffin content in the raw material-paraffin content of >350°C fraction in the product of the first hydrocracking reaction unit*mass fraction of >350°C fraction in the product of the first hydrocracking reaction unit)/ Paraffin content in raw materials;
  • the conversion rate of the sum of naphthenes and aromatics (the sum of the contents of naphthenes and aromatics in the raw material - the sum of the contents of naphthenes and aromatics in the first hydrocracking reaction unit product > 350°C * the first hydrocracking reaction unit Mass fraction of fraction >350°C in the product of the hydrocracking reaction unit)/sum of the content of naphthenes and aromatics in the raw material.
  • the saturation rate of raw aromatics the conversion rate of >350°C distillate for the first hydrocracking reaction unit, the conversion rate of paraffins, the conversion rate of naphthenes and aromatics, those skilled in the art know how to set operating parameters such as hydrogen partial pressure or Reaction pressure, reaction temperature, (liquid hour) volume space velocity, hydrogen oil volume ratio to control.
  • reaction temperature and liquid hourly volume space velocity have the most significant effects on saturation/conversion, especially reaction temperature.
  • step (4) with the temperature determined in step (3) as the predetermined temperature of step (4), with the step size smaller than the initial step size of step (3) as the initial step size of step (4), with the step size smaller than that of step (3) Little target difference is used as the target difference of step (4), repeats the process of step (3);
  • step (3) or (4), or repeat step (4) Perform step (3) or (4), or repeat step (4), until the absolute value of the difference between the desired actual saturation rate/conversion rate and target saturation rate/conversion rate is reached, thereby determining the operating temperature and realizing the Saturation/conversion control.
  • step (2) re-determining the operating parameters of step (2), such as increasing or decreasing one or more of hydrogen partial pressure or reaction pressure, liquid hourly volume space velocity, and hydrogen-oil volume ratio up to 10%, 9%, 8% , 7%, 6%, 5%, 4%, 3%, 2%, 1% or larger or smaller values, and repeat step (2).
  • those skilled in the art may first pre-determine a set of operating parameters, determine the actual saturation rate/conversion rate under the predetermined operating parameters, if the actual saturation rate/conversion rate differs from the target saturation rate/conversion rate by more than 20% , increase or decrease the operating temperature in steps of 16°C until the difference between the actual saturation rate/conversion rate and the target saturation rate/conversion rate is less than 20%; if the operating temperature is increased or decreased in steps of 16°C, it will never be achieved If the difference between the actual saturation rate/conversion rate and the target saturation rate/conversion rate is less than 20%, the step size is changed to 8°C, 4°C, 2°C and 1°C.
  • the range of a is 0.10-4.0, and the range of B is 30-300;
  • the range of parameter a of the linear relationship formula is 0.3-3.0, and the range of parameter B is 100-300;
  • the range of parameter a of the linear relationship formula is 0.2-2.0, and the range of parameter B is 40-150;
  • the range of parameter a of the linear relationship formula is 0.25-2.5, and the range of parameter B is 60-250;
  • the operating temperature is determined by the above linear relationship of conversion.
  • the "monocyclic cycloalkane” in the wax oil raw material oil mainly refers to the monocyclic cycloalkane with long side chains
  • the “monocyclic aromatic hydrocarbon” in the wax oil raw material oil mainly refers to the monocyclic cycloalkanes with long side chains.
  • Chain monocyclic aromatic hydrocarbons, the carbon number of the long side chain hydrocarbons is greater than 20.
  • the hydrocracking catalyst I includes a carrier and an active metal component
  • the carrier includes a heat-resistant inorganic oxide and a molecular sieve
  • the heat-resistant inorganic oxide is selected from silicon oxide or One or several types of alumina
  • the active metal components are selected from at least two metal components in Group VIB metals and Group VIII metals; based on the hydrocracking catalyst I as a whole, in terms of oxides, the 10% to 35% by weight of Group VIB metals and 2% to 8% by weight of Group VIII metals;
  • the molecular sieve is 10% by weight-75% by weight, preferably 20% by weight-60% by weight, such as 35% by weight-45% by weight, and the balance is heat-resistant inorganic oxides.
  • the molar ratio is 20-50, and the pore size is 0.4nm-0.58nm.
  • the molecular sieve is selected from one or more of ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-48, ZSM-50, IM-5, MCM-22, EU-1 molecular sieves Several, more preferably ZSM-5 molecular sieve.
  • the reaction conditions of the second hydrocracking unit are: the hydrogen partial pressure is 3.0MPa-20.0MPa, the reaction temperature is 280°C-400°C, and the liquid hourly volume space velocity is 0.5h -1 -6h -1 , the volume ratio of hydrogen to oil is 300-2000.
  • the >350°C fraction conversion control range of the second hydrocracking reaction unit is 5%-80%; in one of the embodiments of the present invention, in order to obtain refrigerating machine oil products, preferably The control range of >350°C fraction conversion of the second hydrocracking reaction unit is 5%-20%; The control range of the fraction conversion rate at °C is 21%-40%. If the conversion rate of the fraction at >350°C is continuously increased, high-aromatic latent reformate with increased yield can be obtained.
  • the >350°C fraction conversion rate of the second hydrocracking reaction unit 100%*(mass fraction of >350°C fraction in heavy fraction I—mass fraction of >350°C fraction in heavy fraction II)/in heavy fraction I Mass fraction of >350°C fraction.
  • the hydrocracking catalyst II includes a carrier and an active metal component
  • the carrier includes a heat-resistant inorganic oxide and a Y-type molecular sieve
  • the heat-resistant inorganic oxide is selected from the group consisting of One or more of silicon, aluminum oxide, and titanium oxide
  • the active metal component is selected from at least two metal components in Group VIB metals and Group VIII metals; based on the hydrocracking catalyst II as a whole, with In terms of oxides, the VIB group metal is 10% by weight to 35% by weight, and the group VIII metal is 2% by weight to 8% by weight;
  • the Y-type molecular sieve is 5%-55% by weight, and the balance is heat-resistant inorganic oxide.
  • the hydroprocessing catalyst is a supported catalyst
  • the carrier is alumina and silica-alumina
  • the active metal component is at least one metal selected from Group VIB and/or at least A metal selected from Group VIII, said Group VIII metal selected from nickel and/or cobalt, said Group VIB metal selected from molybdenum and/or tungsten, based on the total weight of the hydrotreating catalyst, as an oxide
  • the content of Group VIII metals is 1-15% by weight
  • the content of Group VIB metals is 5-40% by weight.
  • the second aspect of the present invention provides a hydrocracking system, comprising: a hydroprocessing unit, a first hydrocracking unit, and a second hydrocracking unit;
  • the hydroprocessing unit is provided with a wax oil raw material oil inlet, a hydrogen gas inlet, and a reaction effluent outlet, and the hydroprocessing unit is sequentially filled with a hydrogenation protecting agent, an optional hydrodemetallization catalyst, and a hydrofining catalyst;
  • the first hydrocracking unit is provided with a first hydrocracking reaction system and a first separation system, a hydrocracking catalyst I is filled in the first hydrocracking reaction system, and a hydroprocessing unit is provided in the first hydrocracking reaction system
  • the reaction effluent inlet of the hydroprocessing unit is connected with the reaction effluent outlet of the hydroprocessing unit
  • the reaction effluent outlet of the first hydrocracking reaction system is connected with the inlet of the first separation system
  • the first separation system is provided with at least the first hydrogen-rich gas outlet, light fraction I outlet and heavy fraction I outlet;
  • the second hydrocracking unit is provided with a second hydrocracking reaction system and a second separation system, and a hydrocracking catalyst II and/or a hydrotreating catalyst are filled in the second hydrocracking reaction system, and the second hydrocracking reaction system
  • the reaction system is provided with a heavy fraction I inlet and communicated with the heavy fraction I outlet of the first separation system, and the reaction effluent outlet of the second hydrocracking reaction system is communicated with the inlet of the second separation system, and the second separation system is provided with at least the second rich fraction I. Hydrogen gas outlet, light fraction II outlet and heavy fraction II outlet.
  • the first separation system and the second separation system are provided with respective gas-liquid separators and fractionation towers, and are not limited to hot high-pressure separators, cold high-pressure separators, hot low-pressure separators, cold Various combinations of the low-pressure separator and the fractionation tower can be used as long as the separation requirements of the present invention are met.
  • the present invention discloses a hydrocracking method and system that can realize the hydrocracking based on the structure characteristics of hydrocarbon molecules. and ring structure type to carry out selective and efficient conversion, thereby obtaining a product fraction rich in paraffins and a product fraction rich in cyclic hydrocarbons, wherein the content of paraffins in the light fraction I rich in paraffins can meet ⁇ 82% by weight, which can be As a raw material for high-quality steam cracking ethylene plant; the light fraction II rich in cyclic hydrocarbons meets the sum of the mass fractions of naphthenes and aromatics ⁇ 58% by weight, and can be used as high-quality reforming material; in addition, the heavy fraction II rich in naphthenes It has good low temperature fluidity and can be used as a high value-added naphthenic special oil.
  • the present invention can realize the separate conversion of chain hydrocarbons and cyclic hydrocarbons (naphthenes and aromatics) in the wax oil feedstock oil as a whole, and enrich them in each product fraction respectively, so that it can directly obtain Light naphtha rich in paraffins, which can be used as chemical raw materials, and special oil products rich in naphthenes with high added value are of great significance for refining and chemical enterprises to realize high-value utilization of wax oil feedstock oil at low cost.
  • Fig. 1 is a schematic diagram of one embodiment of the hydrocracking method provided by the present invention.
  • Fig. 1 is a schematic diagram of one embodiment of the hydrocracking method provided by the present invention.
  • wax oil raw material oil 1 and hydrogen 2 are contacted with hydrogenation protection agent, optional hydrodemetallization catalyst, and hydrorefining catalyst in the hydroprocessing unit to react successively, and the obtained reaction effluent 3 enters
  • the first hydrocracking unit reacts with the hydrocracking catalyst I in the presence of hydrogen, and the resulting reaction effluent 4 enters the separator I for separation, and the hydrogen-rich gas 5, the first liquid phase stream 6 and the heavy fraction are obtained by separation I 10.
  • Gained first liquid stream 6 enters fractionation unit I for fractionation to obtain low-carbon light hydrocarbons 7, light fraction I8, tower bottom oil 9 (middle fraction I), and gained tower bottom oil 9 can enter fractionation unit II for further fractionation.
  • the resulting heavy fraction I 10 is reacted with the hydrogen gas 11 in the second hydrocracking unit in contact with the hydrocracking catalyst II and/or the hydrotreating catalyst, and the resulting reaction effluent 12 enters the separator II for separation to obtain hydrogen-rich gas 13 and the second liquid-phase stream 14, the second liquid-phase stream 14 of gained enters fractionation unit II and carries out fractionation, obtains overhead oil 15, light fraction II 16 and heavy fraction II 17, and gained overhead oil 15 can enter fractionation unit I to carry out Further fractionation.
  • a method for hydrocracking comprising:
  • Hydroprocessing unit the mixed material of wax oil raw material oil and hydrogen is contacted with hydrogenation protecting agent, optional hydrodemetallization catalyst, hydrorefining catalyst in order to react in hydroprocessing unit, obtain reaction effluent ;
  • step (1) the reaction effluent obtained in step (1) enters the first hydrocracking unit, and reacts with the hydrocracking catalyst 1 in the presence of hydrogen, and the gained reaction effluent after separation, at least Obtain light fraction I and heavy fraction I; Rich in paraffins in the light fraction I, the mass fraction of paraffins in the light fraction I is at least 82%, rich in naphthenes and aromatics in the heavy fraction I, in the heavy fraction I In the hydrocarbon composition of the >350°C fraction in Fraction I, the sum of the mass fractions of naphthenes and aromatics is greater than 82%;
  • step (3) The second hydrocracking unit, the heavy fraction I obtained in step (2) enters the second hydrocracking unit, and reacts in contact with the hydrocracking catalyst II and/or the hydrotreating catalyst in the presence of hydrogen, and the obtained reaction After the effluent is separated, at least a light fraction II and a heavy fraction II are obtained.
  • the initial boiling point of the wax oil stock oil is 300-350 DEG C, selected from normal pressure wax oil, vacuum wax oil, hydrogenated wax oil, One or more of coker wax oil, catalytic cracking heavy cycle oil, and deasphalted oil.
  • reaction conditions of the hydroprocessing unit are: the hydrogen partial pressure is 3.0MPa-20.0MPa, such as 8.0MPa-17.0MPa, and the reaction temperature is 280°C -400°C, such as 340-430°C, the liquid hourly volumetric space velocity (calculated as a hydrofining catalyst) is 0.5h -1 -6h -1 , such as 0.5h -1 -2.0h -1 , and the hydrogen-to-oil volume ratio is 300 -2000, such as 600-1000.
  • the hydrogenation protecting agent contains a carrier and an active metal component loaded on the carrier, and the carrier is selected from alumina, silicon oxide and titanium oxide One or more of them, the active metal component is selected from one or more of Group VIB metals and Group VIII non-noble metals, based on the weight of the hydrogenation protection agent, in terms of oxides, the active metal group Divided into 0.1-15% by weight, the particle size of the hydrogenation protecting agent is 0.5-50.0mm, the bulk density is 0.3-1.2g/cm 3 , and the specific surface area is 50-300m 2 /g.
  • the hydrodemetallization catalyst contains a carrier and an active metal component loaded on the carrier, and the carrier is selected from alumina, silicon oxide and oxide One or more of titanium, the active metal component is selected from one or more of Group VIB metals and Group VIII non-noble metals, based on the weight of the hydrodemetallization catalyst, in terms of oxides, the activity
  • the metal component is 3-30% by weight
  • the hydrodemetallization catalyst has a particle size of 0.2-2.0 mm, a bulk density of 0.3-0.8 g/cm 3 and a specific surface area of 100-250 m 2 /g.
  • the hydrorefining catalyst is a supported catalyst
  • the carrier is alumina and/or silica-alumina
  • the active metal component is at least one one selected from Group VIB metals and/or at least one selected from Group VIII metals
  • said Group VIII metals selected from nickel and/or cobalt, said Group VIB metals selected from molybdenum/or tungsten, to hydrogenate
  • the content of Group VIII metal is 1-15% by weight, and the content of Group VIB metal is 5-40% by weight.
  • the active metal component of the hydrotreating catalyst is selected from two or three of nickel, molybdenum and tungsten metals.
  • reaction conditions of the first hydrocracking unit are: the hydrogen partial pressure is 3.0MPa-20.0MPa, such as 8.0MPa-17.0MPa, and the reaction temperature is 280°C-430°C, such as 280°C-400°C, or 340-430°C, the liquid hourly volume space velocity is 0.5h -1 -6h -1 , such as 0.7h -1 -3.0h -1 , the hydrogen-oil volume ratio is 300-2000, such as 800-1500.
  • A is the mass fraction of paraffins in the wax oil feedstock oil
  • B is the sum of the mass fractions of paraffins, cycloalkanes, and single-ring aromatics in the wax oil feedstock oil
  • the conversion rate of the >350°C fraction of the first hydrocracking reaction unit 100%* (mass fraction of >350°C fraction in the wax oil raw material - mass fraction of >350°C fraction in the reaction product of the first hydrocracking reaction unit) /Mass fraction of >350°C distillate in the wax oil raw material oil.
  • the hydrocracking catalyst I includes a carrier and an active metal component
  • the carrier includes a heat-resistant inorganic oxide and a molecular sieve, and the resistant
  • the thermal inorganic oxide is selected from one or more of silicon oxide or aluminum oxide
  • the active metal component is selected from at least two metal components in Group VIB metals and Group VIII metals;
  • the Group VIB metal is 10% by weight to 35% by weight
  • the Group VIII metal is 2% by weight to 8% by weight;
  • the molecular sieve is 10% by weight to 75% by weight, preferably 20% by weight to 60% by weight, for example 35% by weight to 45% by weight, and the balance is heat-resistant inorganic oxide;
  • the silicon-aluminum molar ratio of the molecular sieve is 20-50, the pore diameter is 0.4nm-0.58nm, and preferably, the specific surface area is 200m 2 /g-400m 2 /g.
  • reaction conditions of the second hydrocracking unit are: the hydrogen partial pressure is 3.0MPa-20.0MPa, such as 8.0MPa-17.0MPa, and the reaction temperature is 280°C-430°C, for example, 280-400°C, the liquid hourly volume space velocity is 0.5h -1 -6h -1 , for example 0.7h -1 -3.0h -1 , the hydrogen-oil volume ratio is 300-2000, for example 800 -1800.
  • the >350°C fraction conversion rate of the second hydrocracking reaction unit 100%*(mass fraction of >350°C fraction in heavy fraction I—mass fraction of >350°C fraction in heavy fraction II)/in heavy fraction I Mass fraction of >350°C fraction.
  • the hydrocracking catalyst II includes a carrier and an active metal component
  • the carrier includes a heat-resistant inorganic oxide and a Y-type molecular sieve
  • the heat-resistant inorganic oxide is selected from one or more of silicon oxide, aluminum oxide, and titanium oxide
  • the active metal component is selected from at least two metal components in Group VIB metals and Group VIII metals
  • the hydrocracking catalyst II as a whole is based on the oxide, the VIB group metal is 10% by weight to 35% by weight, and the group VIII metal is 2% by weight to 8% by weight;
  • the Y-type molecular sieve is 5% by weight to 55% by weight, and the balance is heat-resistant inorganic oxide;
  • the reaction temperature of the second hydrocracking reaction unit is 0-30° C. higher than the temperature of the first hydrocracking reaction unit.
  • the hydrotreating catalyst is a supported catalyst
  • the carrier is alumina and silica-alumina
  • the active metal component is at least one selected From a group VIB metal and/or at least one selected from a group VIII metal selected from nickel and/or cobalt, the group VIB metal selected from molybdenum and/or tungsten, as a hydrotreating catalyst
  • the content of Group VIII metal is 1-15% by weight
  • the content of Group VIB metal is 5-40% by weight;
  • the reaction temperature of the second hydrocracking reaction unit is 0-35° C. lower than the temperature of the first hydrocracking reaction unit.
  • the sum of the mass fractions of naphthenes and aromatics in the light fraction II is at least 58%, and the mass fraction of naphthenes in the heavy fraction II>350°C fraction is at least 50%.
  • the conversion ratio of the paraffins in the feedstock oil is 56%-95%, and the sum of naphthenes and aromatics is The conversion rate is 10%-65%.
  • the hydrocracking catalyst I includes a carrier and an active metal component
  • the carrier includes a heat-resistant Inorganic oxides and molecular sieves, based on the carrier, the molecular sieves are 10% by weight to 75% by weight, preferably 20% by weight to 60% by weight, such as 35% by weight to 45% by weight, and the balance is heat-resistant inorganic oxides;
  • the silicon-aluminum molar ratio of the molecular sieve is 20-50, the pore diameter is 0.4nm-0.58nm, and preferably, the specific surface area is 200m 2 /g-400m 2 /g.
  • fraction cutting is carried out at 65°C-120°C, preferably at 65-105°C, and optionally at Fraction cutting was performed at 160-180°C.
  • the paraffin conversion rate of the raw oil is 56%-95%
  • the conversion rate of the sum of naphthenes and aromatics is 10%-65%, preferably through the reaction temperature and volume space velocity to regulate the conversion rate of paraffins, and the conversion rate of naphthenes and aromatics
  • the paraffin conversion rate (paraffin content in the raw material-paraffin content of >350°C fraction in the product of the first hydrocracking reaction unit*mass fraction of >350°C fraction in the product of the first hydrocracking reaction unit)/ Paraffin content in raw materials;
  • the conversion rate of the sum of naphthenes and aromatics (the sum of the contents of naphthenes and aromatics in the raw material - the sum of the contents of naphthenes and aromatics in the first hydrocracking reaction unit product > 350°C * the first hydrocracking reaction unit Mass fraction of fraction >350°C in the product of the hydrocracking reaction unit)/sum of the content of naphthenes and aromatics in the raw material.
  • step (4) with the temperature determined in step (3) as the predetermined temperature of step (4), with the step size smaller than the initial step size of step (3) as the initial step size of step (4), with the step size smaller than that of step (3) Little target difference is used as the target difference of step (4), repeats the process of step (3);
  • step (3) or (4), or repeat step (4) Perform step (3) or (4), or repeat step (4), until the absolute value of the difference between the desired actual saturation rate/conversion rate and the target saturation rate/conversion rate is reached, thereby determining the operating temperature and realizing the Saturation/conversion control.
  • step (2) re-determining the operating parameters of step (2), such as increasing or decreasing one or more of hydrogen partial pressure or reaction pressure, liquid hourly volume space velocity, and hydrogen-oil volume ratio up to 10%, 9%, 8% , 7%, 6%, 5%, 4%, 3%, 2%, 1% or larger or smaller values, and repeat step (2).
  • the range of a is 0.10-4.0, and the range of B is 30-300;
  • the range of parameter a of the linear relationship formula is 0.3-3.0, and the range of parameter B is 100-300;
  • the range of parameter a of the linear relationship formula is 0.2-2.0, and the range of parameter B is 40-150;
  • the range of parameter a of the linear relationship formula is 0.25-2.5, and the range of parameter B is 60-250;
  • the operating temperature is determined by the above linear relationship of conversion.
  • the hydroprocessing unit is provided with a wax oil raw material oil inlet, a hydrogen gas inlet, and a reaction effluent outlet, and the hydroprocessing unit is sequentially filled with a hydrogenation protecting agent, an optional hydrodemetallization catalyst, and a hydrofining catalyst;
  • the first hydrocracking unit is provided with a first hydrocracking reaction system and a first separation system, a hydrocracking catalyst I is filled in the first hydrocracking reaction system, and a hydroprocessing unit is provided in the first hydrocracking reaction system
  • the reaction effluent inlet of the hydroprocessing unit is connected with the reaction effluent outlet of the hydroprocessing unit
  • the reaction effluent outlet of the first hydrocracking reaction system is connected with the inlet of the first separation system
  • the first separation system is provided with at least the first hydrogen-rich gas outlet, light fraction I outlet and heavy fraction I outlet;
  • the second hydrocracking unit is provided with a second hydrocracking reaction system and a second separation system, and a hydrocracking catalyst II and/or a hydrotreating catalyst are filled in the second hydrocracking reaction system, and the second hydrocracking reaction system
  • the reaction system is provided with a heavy fraction I inlet and communicated with the heavy fraction I outlet of the first separation system, and the reaction effluent outlet of the second hydrocracking reaction system is communicated with the inlet of the second separation system, and the second separation system is provided with at least the second rich fraction I. Hydrogen gas outlet, light fraction II outlet and heavy fraction II outlet.
  • the hydrocracking catalyst I includes a carrier and an active metal component
  • the carrier includes a heat-resistant inorganic oxide and a molecular sieve, based on the carrier, the molecular sieve is 10% by weight to 75% by weight %, preferably, 20% by weight-60% by weight, such as 35% by weight-45% by weight, and the balance is heat-resistant inorganic oxide
  • the silicon-aluminum molar ratio of the molecular sieve is 20-50, and the pore size is 0.4nm-0.58 nm;
  • a control device for controlling fraction cutting at 65°C-120°C, preferably 65-105°C, and optionally a device for controlling fraction cutting at 160-180°C control device In the first hydrocracking unit, there is provided a control device for controlling fraction cutting at 65°C-120°C, preferably 65-105°C, and optionally a device for controlling fraction cutting at 160-180°C control device.
  • the hydrocarbon composition data of the wax oil raw material oil was obtained by SH/T 0659 "Methods for Determination of Hydrocarbons of Saturated Hydrocarbon Fractions in Gas Oil (Mass Spectrometry)".
  • hydrocarbon composition data of light fraction I and light fraction II were obtained by SH/T 0714 "Determination of Monomer Hydrocarbon Composition in Naphtha (Capillary Gas Chromatography)".
  • hydrocarbon composition data of heavy fraction I > 350°C and heavy fraction II > 350°C were obtained through SH/T 0659 "Methods for Determination of Hydrocarbons in Saturated Hydrocarbon Fractions in Gas Oil (Mass Spectrometry)".
  • Table 1 Listed in Table 1 are the properties of the wax oil raw material oil used in the present invention.
  • Table 2 and Table 3 have listed the physical and chemical properties of each catalyst used in the examples and comparative examples of the present invention.
  • the catalysts with trade marks are all produced by Sinopec Catalyst Branch, and the catalysts without trade marks are all conventional fixed-bed The preparation method of supported hydrogenation catalyst is obtained.
  • the mass fraction (A) of paraffins in the wax oil stock oil used in the present invention is 20.4, and the mass fraction sum (B) of paraffins, cycloalkanes, and monocyclic aromatics in the wax stock oil is 49.3.
  • the >350°C fraction conversion control range of the first hydrocracking reaction unit is:
  • A is the mass fraction of paraffins in the wax oil raw material oil
  • B is the sum of the mass fractions of paraffins, monocycloparaffins and monocyclic aromatics in the wax oil raw material oil.
  • control range of the >350°C fraction conversion of the first hydrocracking reaction unit is 22.7-54.7%.
  • the yield of low-carbon light hydrocarbons, the yield of light fraction I, the yield of light fraction II, and the yield of heavy fraction II are all calculated based on wax oil feedstock oil.
  • the mass fraction of the heavy fraction I>350°C is based on the quality of the heavy fraction I; the mass fraction of the heavy fraction II (280-370°C) is based on the mass fraction of the heavy fraction II The mass of the heavy fraction II>350 ° C fraction is based on the quality of the heavy fraction II.
  • the wax oil raw material oil is contacted with the hydrogenation protecting agent (protecting agent), hydrodemetallization catalyst (demetallization agent) and hydrorefining catalyst (refining agent) successively in the hydroprocessing unit for reaction, and the obtained reaction effluent enters the first
  • the hydrocracking unit is contacted with the hydrocracking catalyst I (cracking agent 1) containing ZSM-5 molecular sieve for reaction, and the obtained reaction effluent is separated to obtain light fraction I and heavy fraction I; the obtained heavy fraction I enters the second
  • the hydrocracking unit is in contact with a hydrotreating catalyst (treating agent) for reaction, and the obtained reaction effluent is separated to obtain light fraction II and heavy fraction II.
  • the specific reaction conditions and product properties are shown in Table 4.
  • the aromatics saturation rate of the hydrotreating reaction unit is controlled to 50%
  • the >350°C fraction conversion rate of the first hydrocracking reaction unit is 49.4%
  • the >350°C fraction conversion rate of the second hydrocracking reaction unit is The 350°C fraction conversion was 20%.
  • the obtained light fraction I has a paraffin content of 92.7% by weight and can be used as a high-quality steam cracking ethylene plant raw material;
  • the obtained light fraction II has a naphthene+aromatic content of 62.0% by weight and can be used as a high-quality reforming material;
  • the naphthene+aromatic content in the resulting heavy fraction I>350°C fraction is 82.8% by weight;
  • the freezing point of the gained heavy fraction II (280-370°C) fraction is ⁇ -50°C, and the kinematic viscosity at 40°C is 6.944mm /s, the content of polycyclic aromatics (PCA) is less than 3.0%, and it can be used as transformer oil;
  • the content of naphthenes + aromatics in the obtained heavy fraction II>350°C fraction is 77.8% by weight, and the freezing point is -38°C, which can be used as high-quality ring oil Alkane specialty oils, such as refrigeration oil.
  • Comparative example 1 and comparative example 2 adopt the technical process identical with embodiment 1, and the difference with embodiment 1 is, in comparative example 1, the hydrocracking catalyst (cracking agent) that contains Y-type molecular sieve is filled in the first hydrocracking reaction unit 2); Comparative Example 2 The hydrocracking catalyst (cracking agent 3) containing ⁇ -type molecular sieve is loaded in the first hydrocracking reaction unit.
  • the reaction is carried out under the similar conditions of controlling the saturation rate of aromatics in the hydroprocessing unit, the conversion rate of the first hydrocracking unit >350°C and the conversion rate of the second hydrocracking unit >350°C.
  • the specific reaction conditions and product properties are shown in Table 4.
  • the light fraction I paraffin content of the products of Comparative Example 1 and Comparative Example 2 are 54.9% by weight and 47.9% by weight respectively;
  • the content of cycloalkane+aromatics in the heavy fraction I>350°C of the product is 59.0% by weight and 72.4% by weight respectively, and the content of naphthene+aromatics in the heavy fraction II>350°C of the product is 54.0% by weight and 68.2% by weight respectively, and the freezing point is respectively +28°C and +8°C.
  • the wax oil raw material oil is contacted with the hydrogenation protecting agent (protecting agent), hydrodemetallization catalyst (demetallization agent) and hydrorefining catalyst (refining agent) successively in the hydroprocessing unit for reaction, and the obtained reaction effluent enters the first
  • the hydrocracking unit is contacted with the hydrocracking catalyst I (cracking agent 1) containing ZSM-5 molecular sieve for reaction, and the obtained reaction effluent is separated to obtain light fraction I, middle fraction I and heavy fraction I; the obtained middle fraction I enters the fractionating tower of the second hydrocracking unit for fractionation; the heavy fraction I obtained enters the second hydrocracking unit, and reacts with a hydrotreating catalyst (treating agent), and the resulting reaction effluent is separated to obtain a light fraction II and heavy fraction II.
  • Concrete reaction condition and product property are as shown in table 5.
  • the saturation rate of aromatics in the hydrotreating reaction unit is controlled to be 38.6%
  • the conversion rate of >350°C distillate in the first hydrocracking reaction unit is 47.1%
  • the >350°C fraction conversion rate of the second hydrocracking reaction unit is The 350°C fraction conversion was 5%.
  • the obtained light fraction I has a paraffin content of 91.74% by weight and can be used as a high-quality steam cracking ethylene plant raw material;
  • the obtained light fraction II has a naphthene+aromatic content of 61.55% by weight and can be used as a high-quality reforming material;
  • the naphthene+aromatic content in the resulting heavy fraction I>350°C fraction is 84.7% by weight;
  • the freezing point of the gained heavy fraction II (280-370°C) fraction is ⁇ -50°C, and the kinematic viscosity at 40°C is 7.790mm /s, the content of polycyclic aromatics (PCA) is less than 3.0%, and it can be used as transformer oil;
  • the content of naphthenes + aromatics in the obtained heavy fraction II > 350 ° C is 81.7% by weight, and the freezing point is -38 ° C, which can be used as high-quality ring oil Alkane specialty oils, such as refrigeration oil.
  • the aromatic hydrocarbon saturation rate of the hydrotreating reaction unit is controlled to be 54.6%
  • the >350°C fraction conversion rate of the first hydrocracking reaction unit is 44.4%
  • the >350°C fraction conversion rate of the second hydrocracking reaction unit is The 350°C fraction conversion was 5%.
  • the obtained light fraction I has a paraffin content of 90.18% by weight and can be used as a high-quality steam cracking ethylene plant raw material;
  • the obtained light fraction II has a naphthene+aromatic content of 60.82% by weight and can be used as a high-quality reforming material;
  • the cycloalkane+aromatic content in the resulting heavy fraction I>350°C fraction is 83.5% by weight;
  • the freezing point of the resulting heavy fraction II (280-370°C) fraction is ⁇ -50°C, and the kinematic viscosity at 40°C is 7.065mm /s, the content of polycyclic aromatics (PCA) is less than 3.0%, and it can be used as transformer oil;
  • the content of naphthenes + aromatics in the obtained heavy fraction II>350°C fraction is 80.5% by weight, and the freezing point is -38°C, which can be used as high-quality ring oil Alkane specialty oils, such as refrigeration oil.
  • Example 2 The same process flow as in Example 2 was adopted, but the difference from Example 2 was that the saturation rate of aromatics in the hydrotreating reaction unit was controlled to 59.2%. Concrete reaction condition and product property are as shown in table 5.
  • the paraffin content of the obtained light fraction I is 86.08% by weight, and the naphthene+aromatic content of the gained light fraction II is 56.42% by weight; is 81.3% by weight, the naphthene+aromatic content in the obtained heavy fraction II>350°C fraction is 79.8% by weight, and the freezing point is -38°C.
  • This comparative example does not adopt the preferred range of the present invention, and increasing the saturation rate of aromatic hydrocarbons in the hydrotreating reaction unit will lead to an increase in the ring-opening cracking reaction of naphthenes in the first hydrocracking reaction unit, and the wax oil feedstock oil according to Chain structure and ring structure directional conversion reaction effect will be adversely affected.
  • the wax oil raw material oil is contacted with the hydrogenation protecting agent (protecting agent), hydrodemetallization catalyst (demetallization agent) and hydrorefining catalyst (refining agent) successively in the hydroprocessing unit for reaction, and the obtained reaction effluent enters the first
  • the hydrocracking unit is contacted with the hydrocracking catalyst I (cracking agent 1) containing ZSM-5 molecular sieve for reaction, and the obtained reaction effluent is separated to obtain light fraction I and heavy fraction I; the obtained heavy fraction I enters the second
  • the hydrocracking unit is in contact with the hydrocracking catalyst II (cracking agent 4) for reaction, and the obtained reaction effluent is separated to obtain light fraction II and heavy fraction II.
  • Concrete reaction condition and product property are as shown in table 6.
  • the saturation rate of aromatics in the hydrotreating reaction unit is controlled to be 38.6%
  • the conversion rate of >350°C distillate in the first hydrocracking reaction unit is 47.1%
  • the >350°C fraction conversion rate of the second hydrocracking reaction unit is The 350°C fraction conversion rate was 56.25%.
  • the obtained light fraction I has a paraffin content of 91.74% by weight and can be used as a high-quality steam cracking ethylene plant raw material;
  • the obtained light fraction II has a naphthene+aromatic content of 64.37% by weight and can be used as a high-quality reforming material;
  • the naphthene + aromatic content in the heavy fraction I>350°C cut is 84.7% by weight;
  • the content of polycyclic aromatics (PCA) is less than 3.0%, and it can be used as transformer oil;
  • the content of naphthenes + aromatics in the obtained heavy fraction II>350°C fraction is 65.1% by weight, and the freezing point is -38°C, which can be used as high-quality ring oil Alkane specialty oils, such as refrigeration oil.
  • the saturation rate of aromatics in the hydrotreating reaction unit is controlled to be 38.6%
  • the conversion rate of >350°C distillate in the first hydrocracking reaction unit is 47.1%
  • the >350°C fraction conversion rate of the second hydrocracking reaction unit is The 350°C fraction conversion was 72.4%.
  • the obtained light fraction I has a paraffin content of 91.74% by weight and can be used as a high-quality steam cracking ethylene plant raw material;
  • the obtained light fraction II has a naphthene+aromatic content of 59.64% by weight and can be used as a high-quality reforming material;
  • the cycloalkane+aromatic content in the resulting heavy fraction I>350°C fraction is 84.7% by weight;
  • the freezing point of the resulting heavy fraction II (280-370°C) fraction is ⁇ -50°C, and the kinematic viscosity at 40°C is 6.725mm /s, the content of polycyclic aromatics (PCA) is less than 3.0%, and it can be used as transformer oil;
  • the content of naphthenes + aromatics in the obtained heavy fraction II>350°C fraction is 63.0% by weight, and the freezing point is -35°C, which can be used as high-quality ring oil Alkane specialty oils, such as refrigeration oil.
  • Example 4 Using the same process flow as in Example 4, the difference from Example 4 is that the second hydrocracking reaction unit uses a higher fraction conversion rate of >350°C, and the fraction conversion rate of >350°C is 88.5%. Concrete reaction condition and product property are as shown in table 6.
  • Example 4 Using the same process flow as in Example 4, the difference from Example 4 is that the first hydrocracking reaction unit uses a higher fraction conversion rate of >350°C, and the fraction conversion rate of >350°C is 65%. Concrete reaction condition and product property are as shown in table 6.
  • the product light fraction I paraffin content is 88.25% by weight; the product light fraction II cycloalkane+aromatic content is 61.36% by weight, and the naphthene in the gained heavy fraction I>350 DEG C cut
  • the content of +aromatics is 80.4% by weight, and the content of cycloalkanes+aromatics in the obtained heavy fraction II > 350°C is 59.9% by weight, and the freezing point is -45°C.
  • Embodiment 6 comparative examples 6 and 7
  • the wax oil raw material oil is contacted with the hydrogenation protecting agent (protecting agent), hydrodemetallization catalyst (demetallization agent) and hydrorefining catalyst (refining agent) successively in the hydroprocessing unit for reaction, and the obtained reaction effluent enters the first
  • the hydrocracking unit is contacted with the hydrocracking catalyst I (cracking agent 5-7) containing ZSM-5 molecular sieve to react, and the obtained reaction effluent is separated to obtain light fraction I and heavy fraction I; the obtained heavy fraction I enters
  • the second hydrocracking unit is in contact with a hydrotreating catalyst (treating agent) for reaction, and the obtained reaction effluent is separated to obtain light fraction II and heavy fraction II.
  • the specific reaction conditions and product properties are shown in Table 7.
  • Example 6 Using the same process flow as in Example 6, the difference from Example 6 is that other molecular sieves such as IM-5 and ZSM-48 are used instead of ZSM-5 to obtain qualified products.
  • IM-5 and ZSM-48 are used instead of ZSM-5 to obtain qualified products.
  • Table 8 lists the product quality indicators of transformer oil and refrigerating machine oil.
  • Protective agent Demetallizer Refined preparation treatment agent brand name RG-30A/B RAM-100 RJW-3 RN-32V Metal Ni/Mo Ni/Mo Ni/Mo/W Ni/Mo/W NiO, wt% 0.5-1.5 ⁇ 1 ⁇ 3 ⁇ 2.4 MoO3, wt% 2-6 ⁇ 6 ⁇ 1 ⁇ 2.3 WO3, weight % / / ⁇ 26 ⁇ 23.0 carrier Aluminum oxide Aluminum oxide Aluminum oxide Aluminum oxide Alumina and silica-alumina

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Abstract

一种加氢裂化方法和系统，蜡油原料油和氢气混合物料先经过加氢处理单元接触反应后，反应流出物进入第一加氢裂化单元与加氢裂化催化剂I进行反应，得到富含链烷烃的轻馏分I和富含环状烃的重馏分I，重馏分I与氢气混合后进入第二加氢裂化反应单元进行反应，从而得到富含环状烃的重馏分II。

Description

一种加氢裂化的方法和系统 技术领域
本发明涉及烃类原料处理技术领域，具体涉及一种蜡油原料的加氢裂化的方法和系统。
背景技术
在原油二次加工技术中，加氢裂化技术具有原料适应性强、生产操作和产品方案灵活、产品质量优等优点，能将原料油转化为清洁燃料和化工原料，为炼化企业产品分布和产品质量调整及油化结合的重要加工工艺之一。
加氢裂化的原料油通常为蜡油，蜡油原料由链烷烃、环烷烃和芳烃分子组成，碳数范围约为20-40。现有技术中，常规的加氢裂化催化剂主要以Y型或β型分子筛为催化材料，利用其催化材料的酸性功能进行断链反应，因此，常规加氢裂化技术转化蜡油原料过程中，除发生环烷烃的开环裂化反应外，链烷烃、芳烃或环烷烃分子的长侧链上也会发生断链反应，使得各产品馏分中同时存在链烷烃、有侧链的环烷烃和有侧链的芳烃，导致加氢裂化产品中蒸汽裂解制乙烯原料(尾油、轻石脑油)难以实现链烷烃的高效富集，同时其产品中重整料(重石脑油)难以实现环烷烃和芳烃的高效富集。
CN87105808A公开了一种加氢裂化润滑油基础油料加氢脱蜡改进方法，该方法将加氢裂化或溶剂脱蜡润滑油基础油依次通过具有脱蜡活性的催化剂床层和加氢补充精制催化剂床层，从而生产得到浊点降低的润滑油基础油产品。
CN102959054A公开了一种烃的联合加氢裂化和脱蜡方法，该方法将原料油依次经过加氢处理和第一加氢裂化反应区反应，得到第一加氢裂化反应流出物进入第一催化脱蜡反应区进行反应，反应流出物经分离和分馏得到石脑油馏分、第一柴油馏分和底部产物馏分，其中底部产物馏分在第二加氢裂化或第二催化脱蜡反应区进行反应，反应流出物经分离和分馏形成第二柴油馏分和润滑油产物馏分。
CN102311785A公开了环烷基馏分油加氢生产润滑油基础油的方法，该方法以环烷基原料油为原料，采用含β型分子筛加氢处理催化剂、和含ZSM-5型分子筛的加氢降凝催化剂和加氢补充精制的方法生产倾点降低的橡胶填充油产品。
CN102971401B公开了一种烃类的联合加氢裂化和脱蜡方法，该方法先将原料油进行加氢处理，加氢处理产物分离得到液相残留物进行催化脱蜡和加氢裂化反应，反应流出物经过分离分馏后得到柴油产品馏分和润滑油基础油产品馏分。
CN106609803A公开了一种生产高粘度指数加氢裂化尾油的催化剂及其制备方法，该方法通过将大孔氧化铝、改性USY分子和改性ZSM-48分子筛混合制备得到催化剂，并采用该催化剂使原料发生加氢开环和加氢异构反应生产质量烷烃含量低、异构烃类含量高及年度指数高的润滑油基础油产品。
从上述列举的现有技术可知，常规的加氢裂化技术主要问题为：其一，常规加氢裂化技术，主要采用含有Y型分子筛的加氢裂化催化剂，将蜡油原料油转化为馏程降低的产品馏分，但是不能按照分子结构组成发生相应的裂化反应，未能实现蜡油原料烃分子的按照烃分子结构类型的高效转化，产品质量和附加值较低。其二，现有加氢裂化技术在生产高附加值的环烷基特种产品时，或受限于仅能采用环烷基蜡油原料，或通过将正构链烷烃转化为含支链的异构烷的催化脱蜡反应单元来实现产品低温流动性的改善，其工艺流程复杂，装置投资和操作费用高。
因此，开发出可满足蜡油原料分子按照链结构和环结构分别转化的碳链梯级转化加氢裂化技术，对实现蜡油原料“宜芳则芳、宜烯则烯”的高效利用有重要现实意义。
发明内容
本发明是为了解决现有加氢裂化技术中所存在的蜡油原料分子结构无差别转化所带来的产品附加值低和蜡油原料分子利用效益低的问题。
本发明的第一方面提供一种加氢裂化的方法，包括：
(1)加氢处理单元，蜡油原料油与氢气的混合物料在加氢处理单元依次与加氢保护剂、任选的加氢脱金属催化剂、加氢精制催化剂接触进行反应，得到反应流出物；
(2)第一加氢裂化单元，步骤(1)所得反应流出物进入第一加氢裂化单元，在氢气的存在下与加氢裂化催化剂I接触进行反应，所得反应流出物经分离后，至少得到轻馏分I和重馏分I；所述轻馏分I中富含链烷烃，在轻馏分I中链烷烃的质量分数至少为82％，所述重馏分I中富含环烷烃和芳烃，在重馏分I中>350℃馏分的烃类组成中，环烷烃和芳烃的质量分 数之和大于82％；
(3)第二加氢裂化单元，步骤(2)所得重馏分I进入第二加氢裂化单元，在氢气的存在下与加氢裂化催化剂II和/或加氢处理催化剂接触进行反应，所得反应流出物经分离后，至少得到轻馏分II和重馏分II。
在本发明中，蜡油原料油的初馏点为300-350℃，选自常压蜡油、减压蜡油、加氢蜡油、焦化蜡油、催化裂化重循环油、脱沥青油中的一种或几种。
本发明为了提高蜡油原料油中烃分子的利用价值，提供了一种可实现基于烃分子结构特征的加氢裂化方法，在本发明中蜡油原料油和氢气混合物料先经过加氢处理单元接触反应后，反应流出物进入第一加氢裂化单元与加氢裂化催化剂I进行反应，实现蜡油原料油中链结构的选择性转化得到富含链烷烃的轻馏分I和富含环状烃(环烷烃和芳烃)的重馏分I，重馏分I与氢气混合后进入第二加氢裂化反应单元与加氢裂化催化剂II和/或加氢处理催化剂进行反应，从而得到富含环烷烃和芳烃的轻馏分II及低温流动性好的富含环状烃的重馏分II。本发明从整体上实现了蜡油原料油按照烃分子链结构和环结构类型进行选择性高效转化，分别得到富含链烷烃和富含环状烃的产品馏分。
根据分离方式不同，反应流出物有不同的切割方案，在本发明的其中一种实施方式中，第一加氢裂化单元所得反应流出物经分离后得到轻馏分I和重馏分I，轻馏分I的初馏点为20℃-30℃，轻馏分I和重馏分I的切割点为65℃-120℃，优选为65-105℃；所述轻馏分I中富含链烷烃，优选在轻馏分I中链烷烃的质量分数至少为85％。富含链烷烃的轻馏分I可作为优质蒸汽裂解制乙烯装置原料。所得重馏分I中富含环烷烃和芳烃，在重馏分I中>350℃馏分的烃类组成中，环烷烃和芳烃的质量分数之和大于82％。
在本发明的另一种实施方式中，第一加氢裂化单元所得反应流出物经分离后得到轻馏分I、中馏分I和重馏分I，轻馏分I的初馏点为20℃-30℃，轻馏分I和中馏分I的切割点为65℃-120℃，优选为65-105℃，中馏分I和重馏分I的切割点为160-180℃。所述轻馏分I中富含链烷烃，优选在轻馏分I中链烷烃的质量分数至少为85％。所述中馏分I即可单独作为产品，也可送至第二加氢裂化单元的分馏塔，进行再次切割后，得到部分轻馏分II组分和重馏分II组分。所得重馏分I中富含环烷烃和/或芳 烃，在重馏分I中>350℃馏分的烃类组成中，环烷烃和芳烃的质量分数之和大于82％。
为了进一步提高重馏分I中烃分子的利用价值，本发明将重馏分I送至第二加氢裂化单元进行选择性裂化反应，所得反应流出物经分离后得到轻馏分II和重馏分II。在本发明其中一种实施方式中，所得轻馏分II的初馏点为65℃-100℃，轻馏分II和重馏分II的切割点为155-180℃，优选为160-175℃。在轻馏分II中环烷烃和芳烃的质量分数之和至少为58％，是优质的重整料。根据不同的产品方案，所得重馏分II具有不同的切割方案，在不同的切割方案中，重馏分II可以切割为大比重航煤馏分、变压器油基础油和冷冻机油等多种环烷基特种油品。在本发明其中一种实施方式中，所得重馏分II中>350℃馏分环烷烃质量分数至少为50％。富含环烷烃的重馏分II低温流动性好，重馏分II可作为各种高附加值的环烷基特种油品。
在本发明的一种实施方式中，在加氢处理单元中，以加氢处理单元整体催化剂为基准，加氢保护剂、任选的加氢脱金属催化剂、加氢精制催化剂的装填体积分数分别为：3％-10％；0％-20％；70％-90％。
所述加氢保护剂为本领域常规的重质烃类油品加工的加氢保护剂，不限于蜡油加氢保护剂、渣油加氢保护剂或其级配组合。
在优选的情况下，所述的加氢保护剂含有载体和负载在载体上的活性金属组分，载体选自氧化铝、氧化硅和氧化钛中的一种或几种，活性金属组分选自第VIB族金属、第VIII族非贵金属中的一种或几种，以加氢保护剂的重量为基准，以氧化物计，活性金属组分为0.1-15重量％，加氢保护剂的粒径为0.5-50.0mm，堆密度为0.3-1.2g/cm 3，比表面积为50-300m 2/g。
所述加氢脱金属催化剂为本领域常规的重质烃类油品加工的加氢脱金属催化剂，不限于蜡油加氢脱金属催化剂、渣油加氢脱金属催化剂或其级配组合。
在优选的情况下，所述的加氢脱金属催化剂含有载体和负载在载体上的活性金属组分，载体选自氧化铝、氧化硅和氧化钛中的一种或几种，活性金属组分选自第VIB族金属、第VIII族非贵金属中的一种或几种，以加氢脱金属催化剂的重量为基准，以氧化物计，活性金属组分为3-30重量％，加氢脱金属催化剂的粒径为0.2-2.0mm，堆密度为0.3-0.8g/cm 3， 比表面积为100-250m 2/g。
在本发明中，“任选”的意思是指相应的步骤、催化剂或者成分是可选择的，而非必需的，亦即该步骤、催化剂或者成分可存在也可以不存在。
在本发明的一种实施方式中，所述加氢精制催化剂是负载型催化剂，载体为氧化铝和/或氧化硅-氧化铝，活性金属组分为至少一种选自第VIB族金属和/或至少一种选自第VIII族金属；所述第VIII族金属选自镍和/或钴，所述第VIB族金属选自钼/或钨，以加氢精制催化剂的总重量为基准，以氧化物计，第VIII族金属的含量为1-15重量％，第VIB族金属的含量为5-40重量％，
在优选的情况下，所述加氢精制催化剂的活性金属组分选自镍、钼和钨金属中的两种或三种。
在本发明的一种实施方式中，加氢处理单元的反应条件为：氢分压为3.0MPa-20.0MPa，反应温度为280℃-400℃，液时体积空速为0.5h -1-6h -1，氢油体积比为300-2000。
在优选的情况下，在加氢处理单元，控制原料油的芳烃饱和率小于等于58％。本发明的发明人深入研究发现，如果芳烃饱和率过高，当加氢处理单元的反应流出物进入第一加氢裂化单元后，会导致第一加氢裂化反应单元中环烷烃开环裂化反应增加，对蜡油原料油按照链结构和环结构定向转化反应效果产生不利影响。
原料芳烃饱和率＝100％*(原料芳烃含量-加氢处理反应单元流出物芳烃含量)/原料芳烃含量。
在本文中，如果没有特别指出的话，含量是基于重量的含量。
在本发明的一种实施方式中，第一加氢裂化单元的反应条件为：氢分压为3.0MPa-20.0MPa，反应温度为280℃-400℃，液时体积空速为0.5h -1-6h -1，氢油体积比为300-2000。
为了更好地实现蜡油原料油按照烃分子链结构和环结构类型进行选择性高效转化，在本发明的一种实施方式中，第一加氢裂化反应单元的>350℃馏分转化率控制范围为：
由100*(A重量％/蜡油原料油中>350℃馏分质量分数)至100*(B重量％/蜡油原料油中>350℃馏分质量分数)，
其中，A为蜡油原料油中链烷烃质量分数，B为蜡油原料油中链烷烃、一环烷烃、单环芳烃的质量分数总和，
其中，第一加氢裂化反应单元的>350℃馏分转化率＝100％*(蜡油原料油中>350℃馏分质量分数—第一加氢裂化反应单元反应产物中>350℃馏分质量分数)/蜡油原料油中>350℃馏分质量分数。
同样，为了更好地实现蜡油原料油按照烃分子链结构和环结构类型进行选择性高效转化，在本发明的一种实施方式中，在第一加氢裂化反应单元中，通过调整第一加氢裂化反应单元的反应温度、体积空速、氢油比和反应压力中的一个或多个工艺条件参数使得原料油的链烷烃转化率为56％-95％，环烷烃和芳烃之和的转化率为10％-65％，优选通过反应温度和体积空速调控链烷烃转化率，以及环烷烃和芳烃的转化率，
其中
所述的链烷烃转化率为＝(原料中链烷烃含量-第一加氢裂化反应单元产物中>350℃馏分链烷烃含量*第一加氢裂化反应单元产物中>350℃馏分质量分数)/原料中链烷烃含量；
所述的环烷烃和芳烃之和的转化率为＝(原料中环烷烃和芳烃的含量之和-第一加氢裂化反应单元产物中>350℃馏分环烷烃和芳烃的含量之和*第一加氢裂化反应单元产物中>350℃馏分质量分数)/原料中环烷烃和芳烃的含量之和。
对于原料芳烃饱和率、对于第一加氢裂化反应单元的>350℃馏分转化率、链烷烃转化率、环烷烃和芳烃转化率，本领域技术人员知晓如何通过合理设置操作参数如氢分压或反应压力、反应温度、(液时)体积空速、氢油体积比来进行控制。例如，反应温度和液时体积空速对饱和率/转化率的影响最显著，特别是反应温度。
对于原料芳烃饱和率、对于第一加氢裂化反应单元的>350℃馏分转化率、链烷烃转化率、环烷烃和芳烃转化率，通过下述方式确定操作参数来控制饱和率和/或转化率：
(1)设定实际饱和率/转化率与目标饱和率/转化率的目标差值，如20％，
(2)预定一组操作参数，包括氢分压或反应压力、反应温度、液时体积空速、和氢油体积比，确定该预定操作参数下的实际饱和率/转化率，
(3)当该实际饱和率/转化率与目标饱和率/转化率的差值绝对值大于目标差值时，以一定的步长作为初始步长增加或降低操作温度，直至到达实际饱和率/转化率与目标饱和率/转化率的差值绝对值小于目标差值；
如果以该步长增加或降低操作温度，在操作温度范围内始终无法实现实际饱和率/转化率与目标饱和率/转化率的差值绝对值小于目标差值，则减小步长，从预定温度开始，增加或降低操作温度，直至到达实际饱和率/转化率与目标饱和率/转化率的差值绝对值小于目标差值；
(4)以步骤(3)确定的温度作为步骤(4)的预定温度，以比步骤(3)的初始步长小的步长作为步骤(4)的初始步长，以比步骤(3)小的目标差值作为步骤(4)的目标差值，重复步骤(3)的过程；
(5)进行步骤(3)或(4)，或重复步骤(4)，直至达到期望的实际饱和率/转化率与目标饱和率/转化率的差值绝对值，从而确定操作温度和实现对饱和率/转化率的控制。
如果始终无法达到实际饱和率/转化率与目标饱和率/转化率的差值绝对值小于目标差值，或者如果始终无法达到期望的实际饱和率/转化率与目标饱和率/转化率的差值绝对值，重新预定步骤(2)的操作参数，例如增加或减少氢分压或反应压力、液时体积空速、和氢油体积比中的一个或多个达10％，9％，8％，7％，6％，5％，4％，3％，2％，1％或更大或更小的值，并重复步骤(2)。
举例来说，本领域技术人员可以首先预定一组操作参数，确定该预定操作参数下的实际饱和率/转化率，如果该实际饱和率/转化率与目标饱和率/转化率相差达20％以上，以16℃的步长增加或降低操作温度，直至到达实际饱和率/转化率与目标饱和率/转化率相差达小于20％；如果以16℃的步长增加或降低操作温度，始终无法实现实际饱和率/转化率与目标饱和率/转化率相差达小于20％，则改变步长为8℃、4℃、2℃和1℃。当实际饱和率/转化率与目标饱和率/转化率相差达小于20％，先后根据需要以8℃、4℃、2℃和1℃的步长改变温度直至实现期望的饱和率/转化率。如果始终不能实现期望的饱和率/转化率，则重新预定一组操作参数并重复上述过程。
本领域技术人员知晓如何在给定操作条件的范围内选择特定的操作条件来实现期望的饱和率/转化率。例如，在反应温度、空速、氢油比和氢分压中，对于本技术方案而言，采用固定床加氢裂化工艺，通常在确定的原料处理量下，空速、氢油比和氢分压的调整幅度较小，本领域技术人员主要通过裂化反应温度的调整影响第一加氢裂化反应单元>350℃馏分转化率、链烷烃转化率、环烷烃和芳烃转化率。因此，还可以通过 下述方式确定操作参数来控制转化率：
确定第一加氢裂化反应单元的反应温度和第一加氢裂化反应单元>350℃馏分转化率、链烷烃转化率、环烷烃和芳烃转化率的线性关系，其满足：
y 转化率＝a*反应温度数值-B，
其中a范围为0.10-4.0，B的范围为30-300；
对于第一加氢裂化反应单元>350℃馏分转化率，线性关系公式参数a范围为0.3-3.0，参数B范围为100-300；
对于链烷烃转化率，线性关系公式参数a范围为0.2-2.0，参数B范围为40-150；
对于环烷烃和芳烃转化率，线性关系公式参数a范围为0.25-2.5，参数B范围为60-250；
通过上述转化率的线性关系式，确定操作温度。
在本发明中，所述蜡油原料油中“一环烷烃”主要是指带有长侧链的单环环烷烃，所述蜡油原料油中“单环芳烃”主要是指带有长侧链的单环芳香烃，所述长侧链烃的碳数为大于20。
在本发明的一种实施方式中，所述的加氢裂化催化剂I包括载体和活性金属组分，所述载体包括耐热无机氧化物和分子筛，所述耐热无机氧化物选自氧化硅或氧化铝的一种或几种，所述活性金属组分选自第VIB族金属和第VIII族金属中至少两种金属组分；以加氢裂化催化剂I整体为基准，以氧化物计，第VIB族金属为10重量％-35重量％，第VIII族金属为2重量％-8重量％；
以载体为基准，分子筛为10重量％-75重量％，优选地，20重量％-60重量％，例如35重量％-45重量％，余量为耐热无机氧化物，所述分子筛的硅铝摩尔比为20-50，孔径为0.4nm-0.58nm。
优选所述分子筛选自ZSM-5、ZSM-11、ZSM-12、ZSM-22、ZSM-23、ZSM-48、ZSM-50、IM-5、MCM-22、EU-1分子筛的一种或几种，进一步优选为ZSM-5分子筛。
在本发明的一种实施方式中，第二加氢裂化单元的反应条件为：氢分压为3.0MPa-20.0MPa，反应温度为280℃-400℃，液时体积空速为0.5h -1-6h -1，氢油体积比为300-2000。
在本发明的一种实施方式中，第二加氢裂化反应单元的>350℃馏分 转化率控制范围为5％-80％；在本发明其中一种实施方式中，为了获得冷冻机油产品，优选第二加氢裂化反应单元的>350℃馏分转化率控制范围为5％-20％；在本发明其中一种实施方式中，为了获得变压器油产品，优选第二加氢裂化反应单元的>350℃馏分转化率控制范围为21％-40％，继续提高>350℃馏分转化率，则可得到收率增加的高芳潜重整料。
如果第二加氢裂化反应单元控制过高的>350℃转化率，不仅会降低轻馏分II环烷烃和芳烃含量，而且会导致产品重馏分II馏分质量指标无法满足环烷基特种油品质量要求。
其中，第二加氢裂化反应单元的>350℃馏分转化率＝100％*(重馏分I中>350℃馏分的质量分数—重馏分II中>350℃馏分的质量分数)/重馏分I中>350℃馏分的质量分数。
在本发明的一种实施方式中，所述的加氢裂化催化剂II包括载体和活性金属组分，所述载体包括耐热无机氧化物和Y型分子筛，所述耐热无机氧化物选自氧化硅、氧化铝、氧化钛的一种或几种，所述活性金属组分选自第VIB族金属和第VIII族金属中至少两种金属组分；以加氢裂化催化剂II整体为基准，以氧化物计，第VIB族金属为10重量％-35重量％，第VIII族金属为2重量％-8重量％；
以载体为基准，Y型分子筛为5重量％-55重量％，余量为耐热无机氧化物。
在本发明的一种实施方式中，所述加氢处理催化剂是负载型催化剂，载体为氧化铝和氧化硅-氧化铝，活性金属组分为至少一种选自第VIB族金属和/或至少一种选自第VIII族金属，所述第VIII族金属选自镍和/或钴，所述第VIB族金属选自钼/或钨，以加氢处理催化剂的总重量为基准，以氧化物计，第VIII族金属的含量为1-15重量％，第VIB族金属的含量为5-40重量％。
本发明的第二方面提供一种加氢裂化的系统，包括：加氢处理单元、第一加氢裂化单元、第二加氢裂化单元；
所述加氢处理单元设置蜡油原料油入口、氢气入口、反应流出物出口，在加氢处理单元内依次装填加氢保护剂、任选的加氢脱金属催化剂、加氢精制催化剂；
所述第一加氢裂化单元设置第一加氢裂化反应系统和第一分离系统，在第一加氢裂化反应系统内装填加氢裂化催化剂I，第一加氢裂化反应系 统设置加氢处理单元的反应流出物入口并与加氢处理单元的反应流出物出口连通，第一加氢裂化反应系统的反应流出物出口与第一分离系统的入口连通，第一分离系统至少设置第一富氢气体出口、轻馏分I出口和重馏分I出口；
所述第二加氢裂化单元设置第二加氢裂化反应系统和第二分离系统，在第二加氢裂化反应系统内装填加氢裂化催化剂II和/或加氢处理催化剂，第二加氢裂化反应系统设置重馏分I入口并与第一分离系统的重馏分I出口连通，第二加氢裂化反应系统的反应流出物出口与第二分离系统的入口连通，第二分离系统至少设置第二富氢气体出口、轻馏分II出口和重馏分II出口。
在本发明的一种实施方式中，第一分离系统、第二分离系统分别设置各自的气液分离器和分馏塔，并且不限于热高压分离器、冷高压分离器、热低压分离器、冷低压分离器与分馏塔的各种组合，只要满足本发明的分离需求即可。
本发明为了提高蜡油原料油中烃分子的利用价值，公开了一种可实现基于烃分子结构特征的加氢裂化方法和系统，本发明的特点在于可实现蜡油原料油按照烃分子链结构和环结构类型进行选择性高效转化，从而分别得到富含链烷烃产品馏分和富含环状烃的产品馏分，其中富含链烷烃的轻馏分I中链烷烃含量可满足≮82重量％，可作为优质蒸汽裂解制乙烯装置原料；富含环状烃的轻馏分II满足环烷烃和芳烃质量分数之和≮58重量％，可作为优质重整料；此外、富含环烷烃的产品重馏分II的低温流动性好，可作为高附加值的环烷基特种油品。
本发明能够从整体上实现蜡油原料油中链状烃和环状烃(环烷烃和芳烃)的分别转化，并分别富集在各产品馏分中，从而不需要额外加工过程，就能够直接得到可作为化工原料的富含链烷烃的轻石脑油，以及具有高附加值的富含环烷烃特种油品，对炼化企业低成本实现蜡油原料油的高价值利用具有重要意义。
附图说明
图1是本发明提供的加氢裂化方法的其中一个实施方式的示意图。
具体实施方式
下面结合附图对本发明进行进一步的说明，但并不因此而限制本发明。
图1是本发明提供的加氢裂化方法其中一个实施方式示意图。由图1所示，蜡油原料油1与氢气2一起在加氢处理单元依次与加氢保护剂、任选的加氢脱金属催化剂、加氢精制催化剂接触进行反应，所得反应流出物3进入第一加氢裂化单元，在氢气的存在下与加氢裂化催化剂I接触进行反应，所得反应流出物4进入分离器I进行分离，分离得到富氢气体5、第一液相物流6和重馏分I 10。所得第一液相物流6进入分馏单元I进行分馏，得到低碳轻烃7、轻馏分I 8、塔底油9(中馏分I)，所得塔底油9可进入分馏单元II进行进一步分馏。所得重馏分I 10与氢气11一起在第二加氢裂化单元与加氢裂化催化剂II和/或加氢处理催化剂接触进行反应，所得反应流出物12进入分离器II进行分离，分离得到富氢气体13和第二液相物流14，所得第二液相物流14进入分馏单元II进行分馏，得到塔顶油15、轻馏分II 16和重馏分II 17，所得塔顶油15可进入分馏单元I进行进一步分馏。
本发明提供了下述技术方案以及其任何组合：
1.一种加氢裂化的方法，包括：
(1)加氢处理单元，蜡油原料油与氢气的混合物料在加氢处理单元依次与加氢保护剂、任选的加氢脱金属催化剂、加氢精制催化剂接触进行反应，得到反应流出物；
(2)第一加氢裂化单元，步骤(1)所得反应流出物进入第一加氢裂化单元，在氢气的存在下与加氢裂化催化剂I接触进行反应，所得反应流出物经分离后，至少得到轻馏分I和重馏分I；所述轻馏分I中富含链烷烃，在轻馏分I中链烷烃的质量分数至少为82％，所述重馏分I中富含环烷烃和芳烃，在重馏分I中>350℃馏分的烃类组成中，环烷烃和芳烃的质量分数之和大于82％；
(3)第二加氢裂化单元，步骤(2)所得重馏分I进入第二加氢裂化单元，在氢气的存在下与加氢裂化催化剂II和/或加氢处理催化剂接触进行反应，所得反应流出物经分离后，至少得到轻馏分II和重馏分II。
2.根据前述技术方案中任一项所述的方法，其特征在于，蜡油原料油的初馏点为300-350℃，选自常压蜡油、减压蜡油、加氢蜡油、焦化蜡油、催化裂化重循环油、脱沥青油中的一种或几种。
3.根据前述技术方案中任一项所述的方法，其特征在于，加氢处理单元中，以加氢处理单元整体催化剂为基准，加氢保护剂、任选的加氢 脱金属催化剂、加氢精制催化剂的装填体积分数分别为：3％-10％；0％-20％；70％-90％。
4.根据前述技术方案中任一项所述的方法，其特征在于，加氢处理单元的反应条件为：氢分压为3.0MPa-20.0MPa，例如8.0MPa-17.0MPa，反应温度为280℃-400℃，例如340-430℃，液时体积空速(以加氢精制催化剂计)为0.5h -1-6h -1，例如0.5h -1-2.0h -1，氢油体积比为300-2000，例如600-1000。
5.根据前述技术方案中任一项所述的方法，其特征在于，所述的加氢保护剂含有载体和负载在载体上的活性金属组分，载体选自氧化铝、氧化硅和氧化钛中的一种或几种，活性金属组分选自第VIB族金属、第VIII族非贵金属中的一种或几种，以加氢保护剂的重量为基准，以氧化物计，活性金属组分为0.1-15重量％，加氢保护剂的粒径为0.5-50.0mm，堆密度为0.3-1.2g/cm 3，比表面积为50-300m 2/g。
6.根据前述技术方案中任一项所述的方法，其特征在于，所述的加氢脱金属催化剂含有载体和负载在载体上的活性金属组分，载体选自氧化铝、氧化硅和氧化钛中的一种或几种，活性金属组分选自第VIB族金属、第VIII族非贵金属中的一种或几种，以加氢脱金属催化剂的重量为基准，以氧化物计，活性金属组分为3-30重量％，加氢脱金属催化剂的粒径为0.2-2.0mm，堆密度为0.3-0.8g/cm 3，比表面积为100-250m 2/g。
7.根据前述技术方案中任一项所述的方法，其特征在于，所述加氢精制催化剂是负载型催化剂，载体为氧化铝和/或氧化硅-氧化铝，活性金属组分为至少一种选自第VIB族金属和/或至少一种选自第VIII族金属；所述第VIII族金属选自镍和/或钴，所述第VIB族金属选自钼/或钨，以加氢精制催化剂的总重量为基准，以氧化物计，第VIII族金属的含量为1-15重量％，第VIB族金属的含量为5-40重量％。
8.根据技术方案7所述的方法，其特征在于，加氢精制催化剂的活性金属组分选自镍、钼和钨金属中的两种或三种。
9.根据前述技术方案中任一项所述的方法，其特征在于，在加氢处理单元，控制原料油的芳烃饱和率小于等于58％；任选地，所述原料油的芳烃饱和率＝100％*(原料油中芳烃含量-加氢处理单元反应流出物芳烃含量)/原料油芳烃含量。
10.根据前述技术方案中任一项所述的方法，其特征在于，第一加 氢裂化单元的反应条件为：氢分压为3.0MPa-20.0MPa，例如8.0MPa-17.0MPa，反应温度为280℃-430℃，例如280℃-400℃，或340-430℃，液时体积空速为0.5h -1-6h -1，例如0.7h -1-3.0h -1，氢油体积比为300-2000，例如800-1500。
11.根据前述技术方案中任一项所述的方法，其特征在于，第一加氢裂化反应单元的>350℃馏分转化率控制范围为：
由100*(A重量％/蜡油原料油中>350℃馏分质量分数)至100*(B重量％/蜡油原料油中>350℃馏分质量分数)，
其中，A为蜡油原料油中链烷烃质量分数，B为蜡油原料油中链烷烃、一环烷烃、单环芳烃的质量分数总和，
其中，第一加氢裂化反应单元的>350℃馏分转化率＝100％*(蜡油原料油中>350℃馏分质量分数—第一加氢裂化反应单元反应产物中>350℃馏分质量分数)/蜡油原料油中>350℃馏分质量分数。
12.根据前述技术方案中任一项所述的方法，其特征在于，所述的加氢裂化催化剂I包括载体和活性金属组分，所述载体包括耐热无机氧化物和分子筛，所述耐热无机氧化物选自氧化硅或氧化铝的一种或几种，所述活性金属组分选自第VIB族金属和第VIII族金属中至少两种金属组分；以加氢裂化催化剂I整体为基准，以氧化物计，第VIB族金属为10重量％-35重量％，第VIII族金属为2重量％-8重量％；
以载体为基准，分子筛为10重量％-75重量％，优选地，20重量％-60重量％，例如35重量％-45重量％，余量为耐热无机氧化物；
所述分子筛的硅铝摩尔比为20-50，孔径为0.4nm-0.58nm，优选地，比表面积为200m 2/g-400m 2/g。
13.根据技术方案12所述的方法，其特征在于，所述分子筛选自ZSM-5、ZSM-11、ZSM-12、ZSM-22、ZSM-23、ZSM-48、ZSM-50、IM-5、MCM-22、EU-1分子筛的一种或几种，优选为ZSM-5分子筛。
14.根据前述技术方案中任一项所述的方法，其特征在于，第二加氢裂化单元的反应条件为：氢分压为3.0MPa-20.0MPa，例如8.0MPa-17.0MPa，反应温度为280℃-430℃，例如，280-400℃，液时体积空速为0.5h -1-6h -1，例如0.7h -1-3.0h -1，氢油体积比为300-2000，例如800-1800。
15.根据前述技术方案中任一项所述的方法，其特征在于，第二加 氢裂化反应单元的>350℃馏分转化率控制范围为5％-80％，
其中，第二加氢裂化反应单元的>350℃馏分转化率＝100％*(重馏分I中>350℃馏分的质量分数—重馏分II中>350℃馏分的质量分数)/重馏分I中>350℃馏分的质量分数。
16.根据前述技术方案中任一项所述的方法，其特征在于，所述的加氢裂化催化剂II包括载体和活性金属组分，所述载体包括耐热无机氧化物和Y型分子筛，所述耐热无机氧化物选自氧化硅、氧化铝、氧化钛的一种或几种，所述活性金属组分选自第VIB族金属和第VIII族金属中至少两种金属组分；以加氢裂化催化剂II整体为基准，以氧化物计，第VIB族金属为10重量％-35重量％，第VIII族金属为2重量％-8重量％；
以载体为基准，Y型分子筛为5重量％-55重量％，余量为耐热无机氧化物；
任选地，所述的第二加氢裂化反应单元装填加氢裂化催化剂时，第二加氢裂化反应单元反应温度较第一加氢裂化反应单元温度高0-30℃。
17.根据前述技术方案中任一项所述的方法，其特征在于，所述加氢处理催化剂是负载型催化剂，载体为氧化铝和氧化硅-氧化铝，活性金属组分为至少一种选自第VIB族金属和/或至少一种选自第VIII族金属，所述第VIII族金属选自镍和/或钴，所述第VIB族金属选自钼/或钨，以加氢处理催化剂的总重量为基准，以氧化物计，第VIII族金属的含量为1-15重量％，第VIB族金属的含量为5-40重量％；
任选地，所述的第二加氢裂化反应单元装填加氢处理催化剂时，第二加氢裂化反应单元反应温度较第一加氢裂化反应单元温度低0-35℃。
18.根据前述技术方案中任一项所述的方法，其特征在于，第一加氢裂化单元所得反应流出物经分离后得到轻馏分I和重馏分I，轻馏分I的初馏点为20℃-30℃，轻馏分I和重馏分I的切割点为65℃-120℃，优选为65-105℃；在轻馏分I中链烷烃的质量分数至少为85％。
19.根据前述技术方案中任一项所述的方法，其特征在于，第一加氢裂化单元所得反应流出物经分离后得到轻馏分I和重馏分I，轻馏分I的初馏点为20℃-30℃，轻馏分I和中馏分I的切割点为65℃-120℃，优选为65-105℃，中馏分I和重馏分I的切割点为160-180℃。所述轻馏分I中富含链烷烃，优选在轻馏分I中链烷烃的质量分数至少为85％。
20.根据前述技术方案中任一项所述的方法，其特征在于，轻馏分 II的初馏点为65℃-100℃，轻馏分II和重馏分II的切割点为155-180℃；
在轻馏分II中环烷烃和芳烃的质量分数之和至少为58％，重馏分II>350℃馏分中环烷烃质量分数至少为50％。
21.根据前述技术方案中任一项所述的方法，其特征在于，蜡油原料油的烃类组成中芳烃+环烷烃质量含量为大于70％，例如70％-90％，75％-90％，80％-90％，85-90％，75％-85％，80％-85％。
22.根据前述技术方案中任一项所述的方法，其特征在于，
通过调整第一加氢裂化反应单元的反应温度、体积空速、氢油比和反应压力工艺条件参数控制使得原料油中链烷烃的转化率为56％-95％，环烷烃和芳烃之和的转化率为10％-65％。
23.根据前述技术方案中任一项所述的方法，其特征在于，被输入第一加氢裂化单元进行处理的物流，其芳烃质量含量为10wt％-40wt％，并且以芳烃含量为100wt％计，单环芳烃含量为60wt％-85wt％。
24.根据前述技术方案中任一项所述的方法，其特征在于，被输入第二加氢裂化单元进行处理的物流，其环烷烃和芳烃的质量含量之和为75wt％-90wt％。
25.根据前述技术方案中任一项所述的方法，其特征在于，在第一加氢裂化单元中，所述的加氢裂化催化剂I包括载体和活性金属组分，所述载体包括耐热无机氧化物和分子筛，以载体为基准，分子筛为10重量％-75重量％，优选地，20重量％-60重量％，例如35重量％-45重量％，余量为耐热无机氧化物；所述分子筛的硅铝摩尔比为20-50，孔径为0.4nm-0.58nm，优选地，比表面积为200m 2/g-400m 2/g。
26.根据前述技术方案中任一项所述的方法，其特征在于，在第一加氢裂化单元中，在65℃-120℃，优选为65-105℃进行馏分切割，并且任选地在160-180℃进行馏分切割。
27.根据前述技术方案中任一项所述的方法，其特征在于，蜡油原料油的初馏点为300-350℃，终馏点为520-650℃，20℃密度为0.890g/cm 3-0.940g/cm 3，蜡油原料油的烃类组成中芳烃+环烷烃质量含量为大于70％，例如70％-90％，75％-90％，80％-90％，85-90％，75％-85％，80％-85％，选自常压蜡油、减压蜡油、加氢蜡油、焦化蜡油、催化裂化重循环油、脱沥青油中的一种或几种。
28.根据前述技术方案中任一项所述的方法，其特征在于，
在第一加氢裂化反应单元中，通过调整第一加氢裂化反应单元的反应温度、体积空速、氢油比和反应压力中的一个或多个工艺条件参数使得原料油的链烷烃转化率为56％-95％，环烷烃和芳烃之和的转化率为10％-65％，优选通过反应温度和体积空速调控链烷烃转化率，以及环烷烃和芳烃的转化率，
其中
所述的链烷烃转化率为＝(原料中链烷烃含量-第一加氢裂化反应单元产物中>350℃馏分链烷烃含量*第一加氢裂化反应单元产物中>350℃馏分质量分数)/原料中链烷烃含量；
所述的环烷烃和芳烃之和的转化率为＝(原料中环烷烃和芳烃的含量之和-第一加氢裂化反应单元产物中>350℃馏分环烷烃和芳烃的含量之和*第一加氢裂化反应单元产物中>350℃馏分质量分数)/原料中环烷烃和芳烃的含量之和。
29.根据前述技术方案中任一项所述的方法，其特征在于，对于原料油的芳烃饱和率，和对于第一加氢裂化反应单元的>350℃馏分转化率、链烷烃转化率、环烷烃和芳烃转化率，通过下述方式确定操作参数来控制饱和率/转化率：
(1)设定实际饱和率/转化率与目标饱和率/转化率的目标差值，如20％，
(2)预定一组操作参数，包括氢分压或反应压力、反应温度、液时体积空速、和氢油体积比，确定该预定操作参数下的实际饱和率/转化率，
(3)当该实际饱和率/转化率与目标饱和率/转化率的差值绝对值大于目标差值时，以一定的步长作为初始步长增加或降低操作温度，直至到达实际饱和率/转化率与目标饱和率/转化率的差值绝对值小于目标差值；
如果以该步长增加或降低操作温度，在操作温度范围内始终无法实现实际饱和率/转化率与目标饱和率/转化率的差值绝对值小于目标差值，则减小步长，从预定温度开始，增加或降低操作温度，直至到达实际饱和率/转化率与目标饱和率/转化率的差值绝对值小于目标差值；
(4)以步骤(3)确定的温度作为步骤(4)的预定温度，以比步骤(3)的初始步长小的步长作为步骤(4)的初始步长，以比步骤(3)小的目标差值作为步骤(4)的目标差值，重复步骤(3)的过程；
(5)进行步骤(3)或(4)，或重复步骤(4)，直至达到期望的实际饱和率/ 转化率与目标饱和率/转化率的差值绝对值，从而确定操作温度和实现对饱和率/转化率的控制。
如果始终无法达到实际饱和率/转化率与目标饱和率/转化率的差值绝对值小于目标差值，或者如果始终无法达到期望的实际饱和率/转化率与目标饱和率/转化率的差值绝对值，重新预定步骤(2)的操作参数，例如增加或减少氢分压或反应压力、液时体积空速、和氢油体积比中的一个或多个达10％，9％，8％，7％，6％，5％，4％，3％，2％，1％或更大或更小的值，并重复步骤(2)。
30.根据前述技术方案中任一项所述的方法，其特征在于，对于第一加氢裂化反应单元的>350℃馏分转化率、链烷烃转化率、环烷烃和芳烃转化率，通过下述方式确定操作参数来控制转化率：
确定第一加氢裂化反应单元的反应温度和第一加氢裂化反应单元>350℃馏分转化率、链烷烃转化率、环烷烃和芳烃转化率的线性关系，其满足：
y 转化率＝a*反应温度数值-B，
其中a范围为0.10-4.0，B的范围为30-300；
对于第一加氢裂化反应单元>350℃馏分转化率，线性关系公式参数a范围为0.3-3.0，参数B范围为100-300；
对于链烷烃转化率，线性关系公式参数a范围为0.2-2.0，参数B范围为40-150；
对于环烷烃和芳烃转化率，线性关系公式参数a范围为0.25-2.5，参数B范围为60-250；
通过上述转化率的线性关系式，确定操作温度。
31.用于前述技术方案中任一方法的系统，包括加氢处理单元、第一加氢裂化单元、第二加氢裂化单元；
所述加氢处理单元设置蜡油原料油入口、氢气入口、反应流出物出口，在加氢处理单元内依次装填加氢保护剂、任选的加氢脱金属催化剂、加氢精制催化剂；
所述第一加氢裂化单元设置第一加氢裂化反应系统和第一分离系统，在第一加氢裂化反应系统内装填加氢裂化催化剂I，第一加氢裂化反应系统设置加氢处理单元的反应流出物入口并与加氢处理单元的反应流出物出口连通，第一加氢裂化反应系统的反应流出物出口与第一分离系统的 入口连通，第一分离系统至少设置第一富氢气体出口、轻馏分I出口和重馏分I出口；
所述第二加氢裂化单元设置第二加氢裂化反应系统和第二分离系统，在第二加氢裂化反应系统内装填加氢裂化催化剂II和/或加氢处理催化剂，第二加氢裂化反应系统设置重馏分I入口并与第一分离系统的重馏分I出口连通，第二加氢裂化反应系统的反应流出物出口与第二分离系统的入口连通，第二分离系统至少设置第二富氢气体出口、轻馏分II出口和重馏分II出口。
32.前述技术方案中任一项的系统，其中
在第一加氢裂化单元中，所述的加氢裂化催化剂I包括载体和活性金属组分，所述载体包括耐热无机氧化物和分子筛，以载体为基准，分子筛为10重量％-75重量％，优选地，20重量％-60重量％，例如35重量％-45重量％，余量为耐热无机氧化物；所述分子筛的硅铝摩尔比为20-50，孔径为0.4nm-0.58nm；
在第一加氢裂化单元中，设置有用于控制在65℃-120℃，优选为65-105℃进行馏分切割的控制装置，并且任选地设置有用于控制在160-180℃进行馏分切割的控制装置。
下面结合实施例对本发明作进一步的说明，但并不因此而使本发明受到任何限制。
在实施例和对比例中，蜡油原料油的烃类组成数据通过SH/T 0659《瓦斯油中饱和烃馏分的烃类测定方法(质谱法)》得到。
轻馏分I、轻馏分II的烃类组成数据通过SH/T 0714《石脑油中单体烃组成测定法(毛细管气象色谱法)》得到。
重馏分I>350℃馏分、重馏分II>350℃馏分的烃类组成数据通过SH/T 0659《瓦斯油中饱和烃馏分的烃类测定方法(质谱法)》得到。
表1中列出了本发明所采取的蜡油原料油的性质。
表2、表3中列出了本发明中实施例和对比例所采用各催化剂的理化性质，具有商品牌号的催化剂均为中国石化催化剂分公司生产，无商品牌号的催化剂均为采用常规固定床负载型加氢催化剂制备方法所得。
从表1可以看出，本发明所用蜡油原料油中链烷烃质量分数(A)为20.4，蜡油原料油中链烷烃、一环烷烃、单环芳烃的质量分数总和(B)为49.3。
根据本发明第一加氢裂化反应单元的>350℃馏分转化率控制范围为：
由100*(A重量％/蜡油原料油中>350℃馏分质量分数)至100*(B重量％/蜡油原料油中>350℃馏分质量分数)，
其中，A为蜡油原料油中链烷烃质量分数，B为蜡油原料油中链烷烃、一环烷烃、单环芳烃的质量分数总和。
那么，所述第一加氢裂化反应单元的>350℃馏分转化率控制范围为22.7-54.7％。
在本发明实施例和对比例中，所述低碳轻烃收率、轻馏分I收率、轻馏分II收率、重馏分II收率均是以蜡油原料油为基准计算的。
在本发明实施例和对比例中，所述重馏分I>350℃馏分质量分数，是以重馏分I的质量为基准；重馏分II(280-370℃)馏分质量分数，是以重馏分II的质量为基准；所述重馏分II>350℃馏分质量分数，是以重馏分II的质量为基准。
实施例1
蜡油原料油在加氢处理单元依次与加氢保护剂(保护剂)、加氢脱金属催化剂(脱金属剂)、加氢精制催化剂(精制剂)接触进行反应，所得反应流出物进入第一加氢裂化单元，与含有ZSM-5分子筛的加氢裂化催化剂I(裂化剂1)接触进行反应，所得反应流出物经分离后，得到轻馏分I和重馏分I；所得重馏分I进入第二加氢裂化单元，与加氢处理催化剂(处理剂)接触进行反应，所得反应流出物经分离后，得到轻馏分II和重馏分II。具体反应条件和产品性质如表4所示。
在本实施例的反应过程中，控制加氢处理反应单元的芳烃饱和率为50％，第一加氢裂化反应单元的>350℃馏分转化率为49.4％，第二加氢裂化反应单元的>350℃馏分转化率为20％。
由表4可以看出，所得轻馏分I的链烷烃含量为92.7重量％，可作为优质蒸汽裂解制乙烯装置原料；所得轻馏分II的环烷烃+芳烃含量为62.0重量％，可作为优质重整料；所得重馏分I>350℃馏分中的环烷烃+芳烃含量为82.8重量％；所得重馏分II(280-370℃)馏分的凝点为<-50℃，40℃运动粘度为6.944mm 2/s，稠环芳烃(PCA)含量小于3.0％，可作为变压器油；所得重馏分II>350℃馏分中的环烷烃+芳烃含量为77.8重量％，凝点为-38℃，可作为优质环烷烃特种油品，如冷冻机油。
对比例1和对比例2
对比例1和对比例2采用与实施例1相同的工艺流程，与实施例1不同的是，对比例1中在第一加氢裂化反应单元装填含有Y型分子筛的加氢裂化催化剂(裂化剂2)；对比例2在第一加氢裂化反应单元装填含有β型分子筛的加氢裂化催化剂(裂化剂3)。控制加氢处理单元芳烃饱和率、第一加氢裂化单元>350℃馏分转化率和第二加氢裂化单元>350℃馏分转化率相近条件下，进行反应。具体反应条件和产品性质如表4所示。
由表4可以看出，对比例1和对比例2产品轻馏分I链烷烃含量分别为54.9重量％和47.9重量％；产品轻馏分II环烷烃和芳烃含量分别为60.1重量％和58.6重量％，产品重馏分I>350℃馏分中环烷烃+芳烃含量分别为59.0重量％和72.4重量％，产品重馏分II>350℃馏分中环烷烃+芳烃含量分别为54.0重量％和68.2重量％，凝点分别为+28℃和+8℃。
以上结果表明，采用传统Y型或β型分子筛催化剂的加氢裂化实施方案，难以实现原料链烷烃和环烷烃的高效选择性转化，而采用本发明的方法可实现蜡油原料按照链结构和环结构的定向转化，从而实现生产优质化工原料和高附加值环烷烃特种油品。
实施例2
蜡油原料油在加氢处理单元依次与加氢保护剂(保护剂)、加氢脱金属催化剂(脱金属剂)、加氢精制催化剂(精制剂)接触进行反应，所得反应流出物进入第一加氢裂化单元，与含有ZSM-5分子筛的加氢裂化催化剂I(裂化剂1)接触进行反应，所得反应流出物经分离后，得到轻馏分I、中馏分I和重馏分I；所得中馏分I进入第二加氢裂化单元的分馏塔进行分馏；所得重馏分I进入第二加氢裂化单元，与加氢处理催化剂(处理剂)接触进行反应，所得反应流出物经分离后，得到轻馏分II和重馏分II。具体反应条件和产品性质如表5所示。
在本实施例的反应过程中，控制加氢处理反应单元的芳烃饱和率为38.6％，第一加氢裂化反应单元的>350℃馏分转化率为47.1％，第二加氢裂化反应单元的>350℃馏分转化率为5％。
由表5可以看出，所得轻馏分I的链烷烃含量为91.74重量％，可作为优质蒸汽裂解制乙烯装置原料；所得轻馏分II的环烷烃+芳烃含量为61.55重量％，可作为优质重整料；所得重馏分I>350℃馏分中的环烷烃+芳烃含量为84.7重量％；所得重馏分II(280-370℃)馏分的凝点为<-50℃，40℃运动粘度为7.790mm 2/s，稠环芳烃(PCA)含量小于3.0％，可作为变 压器油；所得重馏分II>350℃馏分中的环烷烃+芳烃含量为81.7重量％，凝点为-38℃，可作为优质环烷烃特种油品，如冷冻机油。
实施例3
采用与实施例2相同的工艺流程，具体反应条件和产品性质如表5所示。
在本实施例的反应过程中，控制加氢处理反应单元的芳烃饱和率为54.6％，第一加氢裂化反应单元的>350℃馏分转化率为44.4％，第二加氢裂化反应单元的>350℃馏分转化率为5％。
由表5可以看出，所得轻馏分I的链烷烃含量为90.18重量％，可作为优质蒸汽裂解制乙烯装置原料；所得轻馏分II的环烷烃+芳烃含量为60.82重量％，可作为优质重整料；所得重馏分I>350℃馏分中的环烷烃+芳烃含量为83.5重量％；所得重馏分II(280-370℃)馏分的凝点为<-50℃，40℃运动粘度为7.065mm 2/s，稠环芳烃(PCA)含量小于3.0％，可作为变压器油；所得重馏分II>350℃馏分中的环烷烃+芳烃含量为80.5重量％，凝点为-38℃，可作为优质环烷烃特种油品，如冷冻机油。
对比例3
采用与实施例2相同的工艺流程，与实施例2不同的是，控制加氢处理反应单元的芳烃饱和率为59.2％。具体反应条件和产品性质如表5所示。
由表5可以看出，所得轻馏分I的链烷烃含量为86.08重量％，所得轻馏分II的环烷烃+芳烃含量为56.42重量％；所得重馏分I>350℃馏分中的环烷烃+芳烃含量为81.3重量％，所得重馏分II>350℃馏分中的环烷烃+芳烃含量为79.8重量％，凝点为-38℃。以上结果表明，本对比例没有采用本发明优选的范围，将加氢处理反应单元的芳烃饱和率提高，会导致第一加氢裂化反应单元中环烷烃开环裂化反应增加，对蜡油原料油按照链结构和环结构定向转化反应效果产生不利影响。
实施例4
蜡油原料油在加氢处理单元依次与加氢保护剂(保护剂)、加氢脱金属催化剂(脱金属剂)、加氢精制催化剂(精制剂)接触进行反应，所得反应流出物进入第一加氢裂化单元，与含有ZSM-5分子筛的加氢裂化催化剂I(裂化剂1)接触进行反应，所得反应流出物经分离后，得到轻馏分I和重馏分I；所得重馏分I进入第二加氢裂化单元，与加氢裂化催化剂II(裂化剂4) 接触进行反应，所得反应流出物经分离后，得到轻馏分II和重馏分II。具体反应条件和产品性质如表6所示。
在本实施例的反应过程中，控制加氢处理反应单元的芳烃饱和率为38.6％，第一加氢裂化反应单元的>350℃馏分转化率为47.1％，第二加氢裂化反应单元的>350℃馏分转化率为56.25％。
由表6可以看出，所得轻馏分I的链烷烃含量为91.74重量％，可作为优质蒸汽裂解制乙烯装置原料；所得轻馏分II的环烷烃+芳烃含量为64.37重量％，可作为优质重整料；所得重馏分I>350℃馏分中的环烷烃+芳烃含量为84.7重量％；所得重馏分II(280-370℃)馏分的凝点为<-50℃，40℃运动粘度为7.801mm 2/s，稠环芳烃(PCA)含量小于3.0％，可作为变压器油；所得重馏分II>350℃馏分中的环烷烃+芳烃含量为65.1重量％，凝点为-38℃，可作为优质环烷烃特种油品，如冷冻机油。
实施例5
采用与实施例4相同的工艺流程，具体反应条件和产品性质如表6所示。
在本实施例的反应过程中，控制加氢处理反应单元的芳烃饱和率为38.6％，第一加氢裂化反应单元的>350℃馏分转化率为47.1％，第二加氢裂化反应单元的>350℃馏分转化率为72.4％。
由表6可以看出，所得轻馏分I的链烷烃含量为91.74重量％，可作为优质蒸汽裂解制乙烯装置原料；所得轻馏分II的环烷烃+芳烃含量为59.64重量％，可作为优质重整料；所得重馏分I>350℃馏分中的环烷烃+芳烃含量为84.7重量％；所得重馏分II(280-370℃)馏分的凝点为<-50℃，40℃运动粘度为6.725mm 2/s，稠环芳烃(PCA)含量小于3.0％，可作为变压器油；所得重馏分II>350℃馏分中的环烷烃+芳烃含量为63.0重量％，凝点为-35℃，可作为优质环烷烃特种油品，如冷冻机油。
对比例4
采用与实施例4相同的工艺流程，与实施例4不同的是，第二加氢裂化反应单元采用较高的>350℃馏分转化率，所述>350℃馏分转化率为88.5％。具体反应条件和产品性质如表6所示。
由表6可知，采用对比例的方法，产品轻馏分I链烷烃含量也为91.74重量％；但产品轻馏分II环烷烃+芳烃含量仅为55.9重量％，且产品重馏分II>350℃馏分中环烷烃+芳烃含量为43.9重量％，凝点为-30.0℃，产品 性质无法满足优质环烷基特种油品质量要求。
以上结果表明，第二加氢裂化反应单元控制过高的>350℃转化率，不仅会降低轻馏分II环烷烃和芳烃含量，而且会导致产品重馏分II质量指标无法满足环烷基特种油品质量要求。
对比例5
采用与实施例4相同的工艺流程，与实施例4不同的是，第一加氢裂化反应单元采用较高的>350℃馏分转化率，所述>350℃馏分转化率为65％。具体反应条件和产品性质如表6所示。
由表6可知，采用对比例5的方法，产品轻馏分I链烷烃含量为88.25重量％；产品轻馏分II环烷烃+芳烃含量为61.36重量％，所得重馏分I>350℃馏分中的环烷烃+芳烃含量为80.4重量％，所得重馏分II>350℃馏分中环烷烃+芳烃含量为59.9重量％，凝点为-45℃，需要说明的是，尽管控制较高的第一加氢裂化转化率也能得到性质合格的产品，但第一加氢裂化单元转化率过高下，产品低碳轻烃(C3+C4)质量分数高达18.5重量％，反应化学氢耗过高，且目标产品分布不合理。
以上结果表明，第一加氢裂化反应单元控制过高的>350℃转化率存在反应过程不经济的问题。
实施例6、对比例6和7
蜡油原料油在加氢处理单元依次与加氢保护剂(保护剂)、加氢脱金属催化剂(脱金属剂)、加氢精制催化剂(精制剂)接触进行反应，所得反应流出物进入第一加氢裂化单元，与含有ZSM-5分子筛的加氢裂化催化剂I(裂化剂5-7)接触进行反应，所得反应流出物经分离后，得到轻馏分I和重馏分I；所得重馏分I进入第二加氢裂化单元，与加氢处理催化剂(处理剂)接触进行反应，所得反应流出物经分离后，得到轻馏分II和重馏分II。具体反应条件和产品性质如表7所示。
分子筛含量过低，会带来链烷烃转化率不足，但过低和过高情况下都存在环烷烃和芳烃开环率高的问题，对于分子筛含量过高情况下，还存在副产物轻烃含量高的问题。
实施例7
采用与实施例6相同的工艺流程，与实施例6不同的是，使用其它的分子筛如IM-5和ZSM-48代替ZSM-5，得到性质合格的产品。
表8中列出了变压器油和冷冻机油的产品质量指标。
表1
项目 中间基VGO原料
密度(20℃)/(g/cm 3) 0.9091
硫/重量％ 2.19
氮/μg·g -1 703
馏程/℃  
IBP/50％/95％ 305/415/482
>350℃馏分质量分数，％ 90.0
烃类组成/质量％  
链烷烃 20.4
一环烷烃 7.0
二环烷烃 10.3
三环烷烃 7
四环烷烃 3.9
五环烷烃 1.3
六环烷烃 0.1
总环烷烃 29.6
单环芳烃 21.9
双环芳烃 11.5
三环芳烃 3.9
四环芳烃 1.8
五环芳烃 0.5
总噻吩+未鉴定芳烃 9.4
总芳烃 50.0
总重量 100.0
A(蜡油原料油中链烷烃质量分数) 20.4
B(蜡油原料油中链烷烃、一环烷烃和单环芳烃质量分数之和) 49.3
表2
项目 保护剂 脱金属剂 精制剂 处理剂
商品牌号 RG-30A/B RAM-100 RJW-3 RN-32V
金属 Ni/Mo Ni/Mo Ni/Mo/W Ni/Mo/W
NiO，重量％ 0.5-1.5 ≮1 ≮3 ≮2.4
MoO3，重量％ 2-6 ≮6 ≮1 ≮2.3
WO3，重量％ / / ≮26 ≮23.0
载体 氧化铝 氧化铝 氧化铝 氧化铝和氧化硅-氧化铝
表3
项目 裂化剂1 裂化剂2 裂化剂3 裂化剂4 裂化剂5 裂化剂6 裂化剂7
金属 NiW NiW NiW NiMo NiW NiW NiW
NiO，重量％ ≮4 ≮3 ≮2.5 ≮4.5 ≮4 ≮4 ≮4
MoO3，重量％ / / / ≮15.5 / / /
WO3，重量％ ≮18 ≮23 ≮25 / ≮18 ≮18 ≮18
分子筛类型 ZSM-5 Y β Y ZSM-5 ZSM-5 ZSM-5
分子筛含量，重量％ 35 15 15 30 45 5 80
孔径/nm 0.5 0.7 0.8 0.7 0.5 0.5 0.5
表4
Figure PCTCN2022118720-appb-000001
Figure PCTCN2022118720-appb-000002
表5
Figure PCTCN2022118720-appb-000003
Figure PCTCN2022118720-appb-000004
表6
Figure PCTCN2022118720-appb-000005
Figure PCTCN2022118720-appb-000006
Figure PCTCN2022118720-appb-000007
表7
Figure PCTCN2022118720-appb-000008
Figure PCTCN2022118720-appb-000009
表8
Figure PCTCN2022118720-appb-000010

Claims (30)

  1. 一种加氢裂化的方法，包括：
    (1)加氢处理单元，蜡油原料油与氢气的混合物料在加氢处理单元依次与加氢保护剂、任选的加氢脱金属催化剂、加氢精制催化剂接触进行反应，得到反应流出物；
    (2)第一加氢裂化单元，步骤(1)所得反应流出物进入第一加氢裂化单元，在氢气的存在下与加氢裂化催化剂I接触进行反应，所得反应流出物经分离后，至少得到轻馏分I和重馏分I；所述轻馏分I中富含链烷烃，在轻馏分I中链烷烃的质量分数至少为82％，所述重馏分I中富含环烷烃和芳烃，在重馏分I中>350℃馏分的烃类组成中，环烷烃和芳烃的质量分数之和大于82％；
    (3)第二加氢裂化单元，步骤(2)所得重馏分I进入第二加氢裂化单元，在氢气的存在下与加氢裂化催化剂II和/或加氢处理催化剂接触进行反应，所得反应流出物经分离后，至少得到轻馏分II和重馏分II。
  2. 根据前述权利要求中任一项所述的方法，其特征在于，蜡油原料油的初馏点为300-350℃，选自常压蜡油、减压蜡油、加氢蜡油、焦化蜡油、催化裂化重循环油、脱沥青油中的一种或几种。
  3. 根据前述权利要求中任一项所述的方法，其特征在于，加氢处理单元中，以加氢处理单元整体催化剂为基准，加氢保护剂、任选的加氢脱金属催化剂、加氢精制催化剂的装填体积分数分别为：3％-10％；0％-20％；70％-90％。
  4. 根据前述权利要求中任一项所述的方法，其特征在于，加氢处理单元的反应条件为：氢分压为3.0MPa-20.0MPa，例如8.0MPa-17.0MPa，反应温度为280℃-400℃，例如340-430℃，液时体积空速(以加氢精制催化剂计)为0.5h -1-6h -1，例如0.5h -1-2.0h -1，氢油体积比为300-2000，例如600-1000。
  5. 根据前述权利要求中任一项所述的方法，其特征在于，所述的加氢保护剂含有载体和负载在载体上的活性金属组分，载体选自氧化铝、氧化硅和氧化钛中的一种或几种，活性金属组分选自第VIB族金属、第VIII族非贵金属中的一种或几种，以加氢保护剂的重量为基准，以氧化物计，活性金属组分为0.1-15重量％，加氢保护剂的粒径为0.5-50.0mm， 堆密度为0.3-1.2g/cm 3，比表面积为50-300m 2/g。
  6. 根据前述权利要求中任一项所述的方法，其特征在于，所述的加氢脱金属催化剂含有载体和负载在载体上的活性金属组分，载体选自氧化铝、氧化硅和氧化钛中的一种或几种，活性金属组分选自第VIB族金属、第VIII族非贵金属中的一种或几种，以加氢脱金属催化剂的重量为基准，以氧化物计，活性金属组分为3-30重量％，加氢脱金属催化剂的粒径为0.2-2.0mm，堆密度为0.3-0.8g/cm 3，比表面积为100-250m 2/g。
  7. 根据前述权利要求中任一项所述的方法，其特征在于，所述加氢精制催化剂是负载型催化剂，载体为氧化铝和/或氧化硅-氧化铝，活性金属组分为至少一种选自第VIB族金属和/或至少一种选自第VIII族金属；所述第VIII族金属选自镍和/或钴，所述第VIB族金属选自钼/或钨，以加氢精制催化剂的总重量为基准，以氧化物计，第VIII族金属的含量为1-15重量％，第VIB族金属的含量为5-40重量％。
  8. 根据权利要求7所述的方法，其特征在于，加氢精制催化剂的活性金属组分选自镍、钼和钨金属中的两种或三种。
  9. 根据前述权利要求中任一项所述的方法，其特征在于，在加氢处理单元，控制原料油的芳烃饱和率小于等于58％；任选地，所述原料油的芳烃饱和率＝100％*(原料油中芳烃含量-加氢处理单元反应流出物芳烃含量)/原料油芳烃含量。
  10. 根据前述权利要求中任一项所述的方法，其特征在于，第一加氢裂化单元的反应条件为：氢分压为3.0MPa-20.0MPa，例如8.0MPa-17.0MPa，反应温度为280℃-430℃，例如280℃-400℃，或340-430℃，液时体积空速为0.5h -1-6h -1，例如0.7h -1-3.0h -1，氢油体积比为300-2000，例如800-1500。
  11. 根据前述权利要求中任一项所述的方法，其特征在于，第一加氢裂化反应单元的>350℃馏分转化率控制范围为：
    由100*(A重量％/蜡油原料油中>350℃馏分质量分数)至100*(B重量％/蜡油原料油中>350℃馏分质量分数)，
    其中，A为蜡油原料油中链烷烃质量分数，B为蜡油原料油中链烷烃、一环烷烃、单环芳烃的质量分数总和，
    其中，第一加氢裂化反应单元的>350℃馏分转化率＝100％*(蜡油原料油中>350℃馏分质量分数—第一加氢裂化反应单元反应产物中>350℃馏 分质量分数)/蜡油原料油中>350℃馏分质量分数。
  12. 根据前述权利要求中任一项所述的方法，其特征在于，所述的加氢裂化催化剂I包括载体和活性金属组分，所述载体包括耐热无机氧化物和分子筛，所述耐热无机氧化物选自氧化硅或氧化铝的一种或几种，所述活性金属组分选自第VIB族金属和第VIII族金属中至少两种金属组分；以加氢裂化催化剂I整体为基准，以氧化物计，第VIB族金属为10重量％-35重量％，第VIII族金属为2重量％-8重量％；
    以载体为基准，分子筛为10重量％-75重量％，优选地，20重量％-60重量％，例如35重量％-45重量％，余量为耐热无机氧化物；
    所述分子筛的硅铝摩尔比为20-50，孔径为0.4nm-0.58nm，优选地，比表面积为200m 2/g-400m 2/g。
  13. 根据权利要求12所述的方法，其特征在于，所述分子筛选自ZSM-5、ZSM-11、ZSM-12、ZSM-22、ZSM-23、ZSM-48、ZSM-50、IM-5、MCM-22、EU-1分子筛的一种或几种，优选为ZSM-5分子筛。
  14. 根据前述权利要求中任一项所述的方法，其特征在于，第二加氢裂化单元的反应条件为：氢分压为3.0MPa-20.0MPa，例如8.0MPa-17.0MPa，反应温度为280℃-430℃，例如，280-400℃，液时体积空速为0.5h -1-6h -1，例如0.7h -1-3.0h -1，氢油体积比为300-2000，例如800-1800。
  15. 根据前述权利要求中任一项所述的方法，其特征在于，第二加氢裂化反应单元的>350℃馏分转化率控制范围为5％-80％，
    其中，第二加氢裂化反应单元的>350℃馏分转化率＝100％*(重馏分I中>350℃馏分的质量分数—重馏分II中>350℃馏分的质量分数)/重馏分I中>350℃馏分的质量分数。
  16. 根据前述权利要求中任一项所述的方法，其特征在于，所述的加氢裂化催化剂II包括载体和活性金属组分，所述载体包括耐热无机氧化物和Y型分子筛，所述耐热无机氧化物选自氧化硅、氧化铝、氧化钛的一种或几种，所述活性金属组分选自第VIB族金属和第VIII族金属中至少两种金属组分；以加氢裂化催化剂II整体为基准，以氧化物计，第VIB族金属为10重量％-35重量％，第VIII族金属为2重量％-8重量％；
    以载体为基准，Y型分子筛为5重量％-55重量％，余量为耐热无机氧化物；
    任选地，所述的第二加氢裂化反应单元装填加氢裂化催化剂时，第二加氢裂化反应单元反应温度较第一加氢裂化反应单元温度高0-30℃。
  17. 根据前述权利要求中任一项所述的方法，其特征在于，所述加氢处理催化剂是负载型催化剂，载体为氧化铝和氧化硅-氧化铝，活性金属组分为至少一种选自第VIB族金属和/或至少一种选自第VIII族金属，所述第VIII族金属选自镍和/或钴，所述第VIB族金属选自钼/或钨，以加氢处理催化剂的总重量为基准，以氧化物计，第VIII族金属的含量为1-15重量％，第VIB族金属的含量为5-40重量％；
    任选地，所述的第二加氢裂化反应单元装填加氢处理催化剂时，第二加氢裂化反应单元反应温度较第一加氢裂化反应单元温度低0-35℃。
  18. 根据前述权利要求中任一项所述的方法，其特征在于，第一加氢裂化单元所得反应流出物经分离后得到轻馏分I和重馏分I，轻馏分I的初馏点为20℃-30℃，轻馏分I和重馏分I的切割点为65℃-120℃，优选为65-105℃；在轻馏分I中链烷烃的质量分数至少为85％。
  19. 根据前述权利要求中任一项所述的方法，其特征在于，第一加氢裂化单元所得反应流出物经分离后得到轻馏分I和重馏分I，轻馏分I的初馏点为20℃-30℃，轻馏分I和中馏分I的切割点为65℃-120℃，优选为65-105℃，中馏分I和重馏分I的切割点为160-180℃。所述轻馏分I中富含链烷烃，优选在轻馏分I中链烷烃的质量分数至少为85％。
  20. 根据前述权利要求中任一项所述的方法，其特征在于，轻馏分II的初馏点为65℃-100℃，轻馏分II和重馏分II的切割点为155-180℃；
    在轻馏分II中环烷烃和芳烃的质量分数之和至少为58％，重馏分II>350℃馏分中环烷烃质量分数至少为50％。
  21. 根据前述权利要求中任一项所述的方法，其特征在于，蜡油原料油的烃类组成中芳烃+环烷烃质量含量为大于70％，例如70％-90％，75％-90％，80％-90％，85-90％，75％-85％，80％-85％。
  22. 根据前述权利要求中任一项所述的方法，其特征在于，
    通过调整第一加氢裂化反应单元的反应温度、体积空速、氢油比和反应压力工艺条件参数控制使得原料油中链烷烃的转化率为56％-95％，环烷烃和芳烃之和的转化率为10％-65％。
  23. 根据前述权利要求中任一项所述的方法，其特征在于，被输入第一加氢裂化单元进行处理的物流，其芳烃质量含量为10wt％-40wt％， 并且以芳烃含量为100wt％计，单环芳烃含量为60wt％-85wt％。
  24. 根据前述权利要求中任一项所述的方法，其特征在于，被输入第二加氢裂化单元进行处理的物流，其环烷烃和芳烃的质量含量之和为75wt％-90wt％。
  25. 根据前述权利要求中任一项所述的方法，其特征在于，在第一加氢裂化单元中，所述的加氢裂化催化剂I包括载体和活性金属组分，所述载体包括耐热无机氧化物和分子筛，以载体为基准，分子筛为10重量％-75重量％，优选地，20重量％-60重量％，例如35重量％-45重量％，余量为耐热无机氧化物；所述分子筛的硅铝摩尔比为20-50，孔径为0.4nm-0.58nm，优选地，比表面积为200m 2/g-400m 2/g。
  26. 根据前述权利要求中任一项所述的方法，其特征在于，在第一加氢裂化单元中，在65℃-120℃，优选为65-105℃进行馏分切割，并且任选地在160-180℃进行馏分切割。
  27. 根据前述权利要求中任一项所述的方法，其特征在于，蜡油原料油的初馏点为300-350℃，终馏点为520-650℃，20℃密度为0.890g/cm 3-0.940g/cm 3，蜡油原料油的烃类组成中芳烃+环烷烃质量含量为大于70％，例如70％-90％，75％-90％，80％-90％，85-90％，75％-85％，80％-85％，选自常压蜡油、减压蜡油、加氢蜡油、焦化蜡油、催化裂化重循环油、脱沥青油中的一种或几种。
  28. 根据前述权利要求中任一项所述的方法，其特征在于，在第一加氢裂化反应单元中，通过调整第一加氢裂化反应单元的反应温度、体积空速、氢油比和反应压力中的一个或多个工艺条件参数使得原料油的链烷烃转化率为56％-95％，环烷烃和芳烃之和的转化率为10％-65％，优选通过反应温度和体积空速调控链烷烃转化率，以及环烷烃和芳烃的转化率，
    其中
    所述的链烷烃转化率为＝(原料中链烷烃含量-第一加氢裂化反应单元产物中>350℃馏分链烷烃含量*第一加氢裂化反应单元产物中>350℃馏分质量分数)/原料中链烷烃含量；
    所述的环烷烃和芳烃之和的转化率为＝(原料中环烷烃和芳烃的含量之和-第一加氢裂化反应单元产物中>350℃馏分环烷烃和芳烃的含量之和*第一加氢裂化反应单元产物中>350℃馏分质量分数)/原料中环烷烃和芳 烃的含量之和。
  29. 用于前述权利要求中任一方法的系统，包括加氢处理单元、第一加氢裂化单元、第二加氢裂化单元；
    所述加氢处理单元设置蜡油原料油入口、氢气入口、反应流出物出口，在加氢处理单元内依次装填加氢保护剂、任选的加氢脱金属催化剂、加氢精制催化剂；
    所述第一加氢裂化单元设置第一加氢裂化反应系统和第一分离系统，在第一加氢裂化反应系统内装填加氢裂化催化剂I，第一加氢裂化反应系统设置加氢处理单元的反应流出物入口并与加氢处理单元的反应流出物出口连通，第一加氢裂化反应系统的反应流出物出口与第一分离系统的入口连通，第一分离系统至少设置第一富氢气体出口、轻馏分I出口和重馏分I出口；
    所述第二加氢裂化单元设置第二加氢裂化反应系统和第二分离系统，在第二加氢裂化反应系统内装填加氢裂化催化剂II和/或加氢处理催化剂，第二加氢裂化反应系统设置重馏分I入口并与第一分离系统的重馏分I出口连通，第二加氢裂化反应系统的反应流出物出口与第二分离系统的入口连通，第二分离系统至少设置第二富氢气体出口、轻馏分II出口和重馏分II出口。
  30. 前述权利要求中任一项的系统，其中
    在第一加氢裂化单元中，所述的加氢裂化催化剂I包括载体和活性金属组分，所述载体包括耐热无机氧化物和分子筛，以载体为基准，分子筛为10重量％-75重量％，优选地，20重量％-60重量％，例如35重量％-45重量％，余量为耐热无机氧化物；所述分子筛的硅铝摩尔比为20-50，孔径为0.4nm-0.58nm；
    在第一加氢裂化单元中，设置有用于控制在65℃-120℃，优选为65-105℃进行馏分切割的控制装置，并且任选地设置有用于控制在160-180℃进行馏分切割的控制装置。
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Cited By (1)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN117089368A (zh) * 2023-08-01 2023-11-21 山东公泉化工股份有限公司 一种生产变压器油的加氢裂化方法

Families Citing this family (3)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN119529892B (zh) * 2023-08-31 2025-08-12 中国石油化工股份有限公司 一种多产重整料的加氢裂化方法
CN120192792B (zh) * 2025-05-23 2025-08-12 中石油克拉玛依石化有限责任公司 Sebs类热塑性弹性体专用白油及其制备方法和应用
CN120383514B (zh) * 2025-06-26 2025-10-28 北京石油化工工程有限公司 一种高碳醇的生产方法和系统

Citations (9)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN87105808A (zh) 1986-08-27 1988-07-06 切夫尔昂研究公司 加氢裂化润滑油基础油料的加氢脱蜡改进方法
AU6563901A (en) * 2001-09-04 2003-03-06 Uop Llc Hydrocracking process product recovery method
CN102311785A (zh) 2010-07-07 2012-01-11 中国石油化工股份有限公司 一种环烷基馏分油加氢生产润滑油基础油的方法
CN102959054A (zh) 2010-06-29 2013-03-06 埃克森美孚研究工程公司 烃的联合加氢裂化和脱蜡
CN102971401A (zh) 2010-06-29 2013-03-13 埃克森美孚研究工程公司 烃类的联合加氢裂化和脱蜡
CN103666555A (zh) * 2012-08-31 2014-03-26 中国石油化工股份有限公司 一种增产中间馏分油的加氢裂化方法
CN105295998A (zh) * 2014-07-25 2016-02-03 中国石油化工股份有限公司 一种由柴油原料生产小分子芳烃的方法
CN106669803A (zh) 2015-11-09 2017-05-17 中国石油化工股份有限公司 生产高粘度指数加氢裂化尾油的催化剂及其制备方法
CN112745920A (zh) * 2019-10-30 2021-05-04 中国石油化工股份有限公司 一种生产高辛烷值汽油的加氢裂化方法

Family Cites Families (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US5611912A (en) * 1993-08-26 1997-03-18 Mobil Oil Corporation Production of high cetane diesel fuel by employing hydrocracking and catalytic dewaxing techniques
CN104560156B (zh) * 2013-10-29 2016-10-05 中国石油化工股份有限公司 一种加氢裂化方法
CN109563418B (zh) * 2016-08-03 2022-03-18 埃克森美孚研究工程公司 用于生产高性能基础油料的萃余油加氢转化
CN112538385B (zh) * 2019-09-20 2022-04-12 中国石油化工股份有限公司 一种加氢与催化裂化的组合方法
CN112725027B (zh) * 2019-10-28 2025-10-03 中国石油化工股份有限公司 一种生产重石脑油的加氢裂化方法
CN113122324B (zh) * 2019-12-31 2022-11-11 中国石油化工股份有限公司 一种催化油浆加氢生产特种油品的方法

Patent Citations (9)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN87105808A (zh) 1986-08-27 1988-07-06 切夫尔昂研究公司 加氢裂化润滑油基础油料的加氢脱蜡改进方法
AU6563901A (en) * 2001-09-04 2003-03-06 Uop Llc Hydrocracking process product recovery method
CN102959054A (zh) 2010-06-29 2013-03-06 埃克森美孚研究工程公司 烃的联合加氢裂化和脱蜡
CN102971401A (zh) 2010-06-29 2013-03-13 埃克森美孚研究工程公司 烃类的联合加氢裂化和脱蜡
CN102311785A (zh) 2010-07-07 2012-01-11 中国石油化工股份有限公司 一种环烷基馏分油加氢生产润滑油基础油的方法
CN103666555A (zh) * 2012-08-31 2014-03-26 中国石油化工股份有限公司 一种增产中间馏分油的加氢裂化方法
CN105295998A (zh) * 2014-07-25 2016-02-03 中国石油化工股份有限公司 一种由柴油原料生产小分子芳烃的方法
CN106669803A (zh) 2015-11-09 2017-05-17 中国石油化工股份有限公司 生产高粘度指数加氢裂化尾油的催化剂及其制备方法
CN112745920A (zh) * 2019-10-30 2021-05-04 中国石油化工股份有限公司 一种生产高辛烷值汽油的加氢裂化方法

Non-Patent Citations (1)

* Cited by examiner, † Cited by third party
Title
See also references of EP4403614A4

Cited By (1)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN117089368A (zh) * 2023-08-01 2023-11-21 山东公泉化工股份有限公司 一种生产变压器油的加氢裂化方法

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