WO2017063309A1 - 一种劣质原料油的处理方法 - Google Patents

一种劣质原料油的处理方法 Download PDF

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WO2017063309A1
WO2017063309A1 PCT/CN2016/000577 CN2016000577W WO2017063309A1 WO 2017063309 A1 WO2017063309 A1 WO 2017063309A1 CN 2016000577 W CN2016000577 W CN 2016000577W WO 2017063309 A1 WO2017063309 A1 WO 2017063309A1
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Prior art keywords
oil
reaction
catalytic cracking
weight
hydrogenation
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PCT/CN2016/000577
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English (en)
French (fr)
Inventor
许友好
刘涛
王新
戴立顺
蓝天
聂红
李大东
Original Assignee
中国石油化工股份有限公司
中国石油化工股份有限公司石油化工科学研究院
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Priority claimed from CN201510671952.0A external-priority patent/CN106590742B/zh
Priority claimed from CN201510672058.5A external-priority patent/CN106590744B/zh
Application filed by 中国石油化工股份有限公司, 中国石油化工股份有限公司石油化工科学研究院 filed Critical 中国石油化工股份有限公司
Priority to GB1807649.7A priority Critical patent/GB2558157B/en
Priority to RU2018117582A priority patent/RU2720990C2/ru
Priority to KR1020187013817A priority patent/KR102648572B1/ko
Priority to SG11201803154QA priority patent/SG11201803154QA/en
Priority to US15/768,437 priority patent/US11365360B2/en
Publication of WO2017063309A1 publication Critical patent/WO2017063309A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/02—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils characterised by the catalyst used
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • 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
    • 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
    • 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
    • C10G65/10—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only including only cracking steps
    • 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
    • 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
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
    • C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
    • C10G69/04—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one step of catalytic cracking in the absence of hydrogen
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • 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
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1077—Vacuum residues
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/201—Impurities
    • C10G2300/205—Metal content
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/02—Gasoline
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/04—Diesel oil
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/26—Fuel gas

Definitions

  • the invention relates to a method for treating inferior raw material oil.
  • CN101210200B discloses a combined process of residue hydrotreating and catalytic cracking.
  • the residue, the catalytic cracking heavy cycle oil from which the solid impurities are removed, the distillate of the optional distillate oil and the optional catalytic cracking slurry are fed together into the residue hydrotreating unit, the resulting hydrocrack and optional
  • the vacuum gas oil enters the catalytic cracking unit together to obtain various products; the catalytic cracking heavy cycle oil for removing solid impurities is recycled to the residue hydrotreating device; the catalytic cracking oil slurry is subjected to distillation separation, and the catalytic cracking oil slurry is subjected to distillation separation
  • the distillate can be recycled to the residue hydrotreating unit.
  • CN102344829A discloses a combined process for residue hydrotreating, catalytic cracking heavy oil hydrogenation and catalytic cracking.
  • the residue hydrogenation tail oil obtained by fractional distillation of the liquid phase stream obtained from the residue hydrogenation reactor enters the catalytic cracking unit as a raw material for catalytic cracking, and the catalytic cracking heavy oil in the catalytic cracking product is mixed with the gas phase stream obtained from the residue hydrogenation reactor Entering the catalytic cracking heavy oil hydrogenation reactor, the hydrogenated catalytic cracking heavy oil is recycled to the catalytic cracking unit.
  • the invention provides a method for treating inferior raw material oil, the method comprising:
  • the inferior feedstock oil is subjected to a low severity hydrogenation reaction, and the obtained reaction product is separated to obtain a gas, a hydrogenated naphtha, a hydrogenated diesel oil and a hydrogenated residue; wherein, in a low severity hydrogenation reaction, Based on the inferior feedstock oil, the yield of the hydrocracked oil is from 85 to 95% by weight, and the properties of the hydrorebase residue are substantially constant;
  • the first hydrocracking reaction obtained in step a is subjected to a first catalytic cracking reaction, and the obtained reaction product is separated to obtain a first dry gas, a first liquefied gas, a first gasoline, a first diesel oil and a first wax oil;
  • the first wax oil obtained in the step b is subjected to a wax oil hydrogenation reaction, and the obtained reaction product is obtained.
  • the hydrogenated wax oil is isolated;
  • the hydrogenated wax oil obtained in the step c is subjected to the first catalytic cracking reaction or the second catalytic cracking reaction described in the step b.
  • the method further comprises the step of: e. the second wax oil obtained in the second catalytic cracking reaction described in step d is subjected to the wax oil hydrogenation reaction described in step c.
  • the yield of the hydrorebase is from 87 to 93% by weight based on the inferior feedstock oil, The properties of the hydrocure oil remain substantially constant.
  • the properties of the hydrocracking oil are undesirably changed (for example, the density is increased or the carbon residue value is increased), the severity of the hydrogenation reaction is increased, and the properties of the hydroredebric oil are added to the initial stage of operation (for example, 0-1000 h).
  • the properties of the hydrogen residue remain essentially constant.
  • the growth rate of the density of the hydroresin exceeds 0.005 g/cm 3 /(1000 hours), and/or when the growth rate of the residual carbon value of the hydro residue exceeds 0.5% by weight/(1000 hours)
  • Increase the severity of the hydrogenation reaction for example, increase the reaction temperature by 2-10 ° C / (1000 hours) or reduce the liquid hourly space velocity by 0.1-0.5 h -1 / (1000 hours).
  • the inferior feedstock oil has a desulfurization rate of 50 to 95% by weight, a denitrification rate of 10 to 70% by weight, and a decarburization ratio of 10 to 70% by weight.
  • the metal ratio is 50 to 95% by weight.
  • the reaction conditions for the low severity hydrogenation reaction include: a hydrogen partial pressure of 8-20 MPa, a reaction temperature of 330-420 ° C, a liquid hourly space velocity of 0.1-1.5 hr -1 , total hydrogen oil.
  • the volume ratio is 200-1500 standard cubic meters per cubic meter.
  • the low severity hydrogenation reaction has a reaction temperature of 350-370 ° C at an initial stage of operation (eg, 0-1000 h), such as 350-360 ° C, 350-355 ° C, or, for example, 350 ° C, 351°C, 352°C, 353°C, 354°C, 355°C, 356°C, 357°C, 358°C, 359°C, 360°C, 361°C, 362°C, 363°C, 364°C, 365°C, 366°C, 367°C , 368 ° C, 369 ° C or 370 ° C.
  • an initial stage of operation eg, 0-1000 h
  • the low severity hydrogenation reaction is carried out in a fixed bed reactor in the presence of a hydrogenation catalyst.
  • the hydrogenation catalyst used for the low-calorie hydrogenation reaction may sequentially include a hydrogenation protecting agent, a hydrodemetallization catalyst, a hydrodesulfurization catalyst, and addition according to the flow direction of the reactant. Hydrogen denitrification decarburization catalyst.
  • the hydrogenation protecting agent and the hydrodemetallization catalyst account for 20%-70%, for example 30%-50%, based on the total weight of the hydrogenation catalyst; the hydrodesulfurization catalyst accounts for 20%-70%, for example 40 %-60%; hydrodenitrogenation The catalyst accounts for 0% to 60%, for example, 10% to 40%, and the sum of the hydrogenation protecting agent, the hydrodemetallization catalyst, the hydrodesulfurization catalyst, and the hydrodenitrogenation-removing carbon catalyst is 100% by weight.
  • the hydrogenation catalysts are those conventionally used in the art.
  • the hydrodemetallization catalyst comprises 30% by weight or more based on the total weight of the hydrogenation catalyst.
  • the inferior feedstock oil is petroleum hydrocarbon and/or other mineral oil
  • the petroleum hydrocarbon is selected from the group consisting of atmospheric gas oil, vacuum gas oil, atmospheric residue, vacuum residue, hydrogenation Residual oil, coker gas oil, deasphalted oil, and any combination thereof
  • the other mineral oils are selected from the group consisting of coal and natural gas derived liquid oils, oil sand oils, tight oils, shale oils, and any combination thereof.
  • the inferior feedstock oil satisfies: (1) a density of 910-1000 kg/m 3 at 20 ° C; and/or (2) a carbon residue specific gravity of 4-15 wt%; / or (3) metal (Ni + V) content of 12-600ppm.
  • the inferior feedstock oil satisfies: (1) a density of 980-1000 kg/m 3 at 20 ° C; and/or (2) a residual carbon specific gravity of 10-15 wt% And/or (3) the metal (Ni + V) content is 60-600 ppm.
  • the first catalytic cracking reaction comprises the steps of: (1) first pre-heating the pre-heated hydrocrack with the first regenerated catalytic cracking catalyst in a lower portion of the first catalytic cracking reactor a reaction, the obtained reaction product is separated to obtain a first cracked product and a first semi-regenerated catalytic cracking catalyst; the first regenerated catalytic cracking catalyst has a micro-reverse evaluation activity of 35-60; (2) in step (1) The obtained first cracked product and the first semi-regenerated catalytic cracking catalyst are then subjected to a first recatalytic conversion reaction in an upper portion of the first catalytic cracking reactor, and the obtained reaction product is separated and fractionated to obtain a first dry gas.
  • a first liquefied gas a first gasoline, a first diesel oil, and a first wax oil.
  • the lower portion and the upper portion of the first catalytic cracking reactor are demarcated from a position between the first 1/3 portion of the reactor and the first 2/3 portion (in the direction in which the reactant flows); in a preferred implementation
  • the lower part refers to the first 1/2 part of the length of the reactor
  • the upper part refers to the last 1/2 part of the length of the reactor.
  • the first cracking reaction is carried out under the following conditions: a reaction temperature of 530-620 ° C, a weight hourly space velocity of 30-180 hr -1 , a ratio of agent to oil of 4-12, and a water-oil ratio of 0.03-0.3, the reaction pressure is 130 kPa-450 kPa; the first re-catalytic conversion reaction is carried out under the following conditions: a reaction temperature of 460 ° C - 520 ° C, a weight hourly space velocity of 20-100 hr -1 , a dose
  • the oil ratio is 3-15, the water-oil ratio is 0.03-0.3, and the reaction pressure is 130 kPa-450 kPa.
  • the first wax oil in the first catalytic cracking reaction, has a hydrogen content of 10.5-15% by weight; based on the hydrogenated residue, the first Production of a wax oil The rate is 15-50% by weight.
  • the secondary processing wax oil is subjected to the wax oil hydrogenation reaction together with the first wax oil; the secondary processing wax oil is selected from the group consisting of coking wax oil, deasphalted oil, Catalytic cracking wax oil produced by other devices, and any combination thereof.
  • the wax oil hydrogenation reaction is carried out in a fixed bed reactor in the presence of a hydrogenation catalyst.
  • the hydrogenation catalyst used for the hydrogenation reaction of the wax oil may sequentially include a hydrogenation protecting agent, a hydrodemetallization desulfurization catalyst, and a hydrotreating catalyst in accordance with the flow direction of the reactant.
  • the hydrogenation protecting agent comprises from 0 to 30% by weight, for example from 5 to 20% by weight, based on the total weight of the hydrogenation catalyst
  • the hydrodemetallization desulfurization catalyst comprises from 5 to 35% by weight, for example from 10 to 25% by weight.
  • the hydrotreating catalyst accounts for 35% to 95% by weight, for example, 55 to 85% by weight, and the sum of the hydrogenation protecting agent, the hydrodemetallization desulfurization catalyst, and the hydrotreating catalyst is 100% by weight.
  • the hydrogenation catalysts are those conventionally used in the art.
  • the wax oil hydrogenation reaction is carried out under the following conditions: a reaction pressure of 5.0 to 20.0 MPa, a reaction temperature of 300 to 430 ° C, and a liquid hourly space velocity of 0.2 to 5.0 hr -1 , hydrogen oil
  • the volume ratio is 200-1800 standard cubic meters per cubic meter.
  • the second catalytic cracking reaction is carried out under the following conditions: a reaction temperature of 450 ° C to 620 ° C, a weight hourly space velocity of 1-100 hr -1 , a ratio of 1 to 25 oil to oil, and water oil. The ratio is 0.03-0.3.
  • the second catalytic cracking reaction comprises the steps of: (1) performing a second cracking of the preheated hydrogenated wax oil and the second regenerated catalytic cracking catalyst in a lower portion of the second catalytic cracking reactor. a reaction, the obtained reaction product is separated to obtain a second cracked product and a second semi-regenerated catalytic cracking catalyst; (2) the second cracked product obtained in the step (1) and the second semi-regenerated catalytic cracking catalyst are subsequently Performing a second re-catalytic conversion reaction on the upper portion of the second catalytic cracking reactor, and the obtained reaction product is separated and fractionated to obtain a second dry gas, a second liquefied gas, a second gasoline, a second diesel oil, and a second wax oil. .
  • the lower portion and the upper portion of the second catalytic cracking reactor are demarcated by a position between the front 1/3 portion and the first 2/3 portion of the reactor (in the direction in which the reactant flows); in a preferred implementation
  • the lower part refers to the first 1/2 part of the length of the reactor
  • the upper part refers to the last 1/2 part of the length of the reactor.
  • the second cracking reaction is carried out under the following conditions: a reaction temperature of 530-620 ° C, a weight hourly space velocity of 30-180 hours -1 , a ratio of the agent to the oil of 4-12, and a water-oil ratio of 0.03-0.3, the reaction pressure is 130 kPa-450 kPa; the conditions of the second re-catalytic conversion reaction are: reaction temperature is 460 ° C - 520 ° C, weight hourly space velocity is 20-100 hr -1 , agent oil The ratio is 3-15, the water-oil ratio is 0.03-0.3, and the reaction pressure is 130 kPa-450 kPa.
  • a method for treating heavy feedstock oil comprising:
  • the heavy feedstock oil is subjected to a shallow hydrogenation reaction to obtain a hydrogenation gas, a hydrogenated naphtha, a hydrogenated diesel oil, and a hydrogenated residue; wherein the heavy feedstock oil is used as a reference to control the addition
  • the yield of the hydrogen residue oil is 85-95% by weight
  • step b contacting the hydrogen residue obtained in step a with the catalytic cracking catalyst and performing a first catalytic cracking reaction to obtain a first dry gas, a first liquefied gas, a first gasoline, a first light cycle oil, a first wax oil, and An external sputum slurry; wherein the catalytic cracking catalyst has a microreverse evaluation activity of 40-55;
  • the first wax oil obtained in step b is filtered and then subjected to a wax oil hydrogenation reaction to obtain a hydrogenated wax oil;
  • the outer eucalyptus oil slurry obtained in the step b is subjected to the first catalytic cracking reaction in the step b;
  • the hydrogenated wax oil obtained in the step c is subjected to a second catalytic cracking reaction or the first catalytic cracking reaction.
  • the method further comprises the step of: the second wax oil obtained in the second catalytic cracking reaction in the step d is subjected to the wax oil hydrogenation reaction in the step c.
  • the yield of the hydrorebase is controlled to be 87 to 93% by weight based on the heavy raw material oil.
  • the desulfurization rate of the heavy feedstock oil is controlled to be 50-95% by weight, the denitrification rate is 20-70% by weight, and the carbon removal rate is 20%. -70% by weight, the demetallization rate is 50-90% by weight.
  • the conditions of the shallow hydrogenation reaction are: a hydrogen partial pressure of 10-20 MPa, a reaction temperature of 320-420 ° C, and a liquid hour volume space velocity of 0.2- 1.0 hour -1, the total hydrogen oil volume ratio is 300-1500 standard cubic meters / cubic meter.
  • the heavy feedstock oil is petroleum hydrocarbon and/or other mineral oil
  • the petroleum hydrocarbon is selected from the group consisting of atmospheric gas oil, vacuum gas oil, atmospheric residue, and reduction.
  • reaction conditions of the first catalytic cracking in the step b are: a reaction temperature of 450-670 ° C, a weight hourly space velocity of 10-100 hours -1 , a regenerated catalyst and a raw material
  • the oil weight ratio is from 1 to 30, and the weight ratio of water vapor to raw material is from 0.03 to 1.0.
  • the yield is 15-50% by weight.
  • the outer mash slurry obtained in the step b has a solid content of less than 6 g/liter and a density at 20 ° C of 920 to 1150 kg/m 3 .
  • the secondary processing wax oil is subjected to a hydrogenation reaction of the wax oil in the step c together with the first wax oil;
  • the secondary processing wax oil is selected from the group consisting of coking wax oil At least one of catalytic cracking wax oil produced by deasphalted oil and other devices.
  • condition of the wax oil hydrogenation reaction in the step c is: a reaction pressure of 6.0-18.0 MPa, a reaction temperature of 270-420 ° C, and a volumetric space velocity of 0.2- 1.0 hour -1, the hydrogen oil volume ratio is 200-1800 standard cubic meters / cubic meter.
  • the second catalytic cracking reaction in the step d is: a reaction temperature of 450 ° C to 620 ° C, a weight hourly space velocity of 1-100 hours -1 , a ratio of the ratio of the agent to the oil It is 1-25 and the water to oil ratio is 0.03-0.3.
  • the inferior feedstock oil is subjected to a shallow hydrogenation reaction to obtain a gas, a hydrogenated naphtha, a hydrogenated diesel oil, and a hydrogenated residue; wherein the hydrous residue is controlled based on the inferior feedstock oil
  • the yield is 85-95% by weight
  • the first hydrocracking reaction obtained in step a is subjected to a first catalytic cracking reaction to obtain a first dry gas, a first liquefied gas, a first gasoline, a first diesel oil, and a first wax oil;
  • step b the first wax oil obtained in step b is subjected to a wax oil hydrogenation reaction to obtain a hydrogenated wax oil;
  • the second catalytic cracking reaction is carried out on the hydrogenated wax oil obtained in the step c to obtain a second dry gas, a second liquefied gas, a second gasoline, a second diesel oil and a second wax oil.
  • the yield of the hydrorelag is controlled to be 87 to 93% by weight based on the inferior feedstock oil.
  • step a the desulfurization rate of the inferior feedstock oil is controlled to be 50-95% by weight, the denitrification rate is 10-70% by weight, and the removal carbon ratio is 10- 70% by weight, the demetallization rate is 50-95% by weight.
  • condition of the shallow hydrogenation reaction is: a hydrogen partial pressure of 8-20 MPa, a reaction temperature of 330-420 ° C, and a liquid hour volume space velocity of 0.1- 1.5 hours -1, the total hydrogen oil volume ratio is 200-1500 standard cubic meters / cubic meter.
  • the inferior feedstock oil is petroleum hydrocarbon and/or other mineral oil
  • the petroleum hydrocarbon is selected from the group consisting of atmospheric gas oil, vacuum gas oil, atmospheric residue, and decompression. At least one of residual oil, hydrocracked oil, coker gas oil and deasphalted oil, the other mineral oil is selected from at least one of coal and natural gas derived liquid oil, oil sand oil, tight oil and shale oil .
  • the inferior feedstock oil has a density of from 920 to 1100 kg/m 3 at 20 ° C and a residual carbon specific gravity of from 8 to 20% by weight.
  • (1) the pre-heated hydrocrack and the first regenerated catalytic cracking catalyst are subjected to a first cracking reaction together with a lower portion of the first catalytic cracking reactor to obtain a first cracked product and a first semi-regenerated catalytic cracking catalyst;
  • the first cracked product obtained in the step (1) and the first semi-regenerated catalytic cracking catalyst are subsequently subjected to a first recatalytic conversion reaction in the upper portion of the first catalytic cracking reactor, and subjected to separation and fractionation to obtain the The first kilo gas, the first liquefied gas, the first gasoline, the first diesel oil, and the first wax oil.
  • condition of the first cracking reaction in the step (1) is: a reaction temperature of 530-620 ° C, a weight hourly space velocity of 30-180 hours -1, a ratio of the ratio of the agent to the oil 4-12, the water-oil ratio is 0.03-0.3, the reaction pressure is 130 kPa-450 kPa;
  • the first re-catalyzed conversion reaction in the step (2) is: the reaction temperature is 460 ° C -520 ° C,
  • the weight hourly space velocity is 20-100 hours-1, the ratio of agent to oil is 3-15, the ratio of water to oil is 0.03-0.3, and the reaction pressure is 130 kPa-450 kPa.
  • the yield is 15-50% by weight.
  • the secondary processing wax oil is subjected to the wax oil hydrogenation reaction in the step c together with the first wax oil; the secondary processing wax oil is selected from the group consisting of coking wax oil At least one of catalytic cracking wax oil produced by deasphalted oil and other devices.
  • condition of the wax oil hydrogenation reaction in the step c is: a reaction pressure of 5.0-20.0 MPa, a reaction temperature of 300-430 ° C, and a volumetric space velocity of 0.2- 5.0 hours-1, the hydrogen oil volume ratio is 200-1800 standard cubic meters / cubic meter.
  • condition of the second catalytic cracking reaction in the step d is: a reaction temperature of 450 ° C to 620 ° C, a weight hourly space velocity of 1-100 hours -1 , a ratio of the ratio of the agent to the oil It is 1-25 and the water to oil ratio is 0.03-0.3.
  • the preheated hydrogenated wax oil and the second regenerated catalytic cracking catalyst are subjected to a second cracking reaction together with the second catalytic cracking reactor to obtain a second cracking product and a second semi-regenerated catalytic cracking catalyst;
  • the second cracked product obtained in the step ( ⁇ ) and the second semi-regenerated catalytic cracking catalyst are subsequently subjected to a second recatalytic conversion reaction in the upper portion of the second catalytic cracking reactor, and subjected to separation and fractionation to obtain the a second dry gas, a second liquefied gas, a second gasoline, a second diesel oil, and a second wax oil.
  • the second cracking reaction in the step ( ⁇ ) is carried out under the conditions of a reaction temperature of 530-620 ° C and a weight hourly space velocity of 30-180 hr-1, a ratio of the ratio of the agent to the oil. 4-12, the water-oil ratio is 0.03-0.3, the reaction pressure is 130 kPa-450 kPa; the second re-catalyzed conversion reaction in the step ( ⁇ ) is: the reaction temperature is 460 ° C -520 ° C, The weight hourly space velocity is 20-100 hours-1, the ratio of agent to oil is 3-15, the ratio of water to oil is 0.03-0.3, and the reaction pressure is 130 kPa-450 kPa.
  • the invention also includes any possible combination of the above-described embodiments and/or technical solutions.
  • the life of the hydrogenation catalyst can be increased, the operating cycle of the hydrogenation unit can be significantly prolonged, and the chemical hydrogen consumption can be reduced.
  • FIG. 1 is a schematic flow chart of a method for treating a poor quality feedstock oil of the present invention, wherein:
  • the invention provides a method for treating inferior feedstock oil, the method comprising: a, performing low-calorie hydrogenation reaction on the inferior feedstock oil, and separating the obtained reaction product to obtain gas, hydrogenated naphtha, hydrogenated diesel oil and a hydrocracked oil; wherein, in a low-calorie hydrogenation reaction, the yield of the hydrorebranched oil is from 85 to 95% by weight, preferably from 87 to 93% by weight, based on the inferior feedstock oil.
  • the properties of the hydrocracking oil are kept substantially constant; b, the hydrocracking oil obtained in the step a is subjected to a first catalytic cracking reaction, and the obtained reaction product is separated to obtain a first dry gas, a first liquefied gas, and a first Gasoline, first diesel oil and first wax oil; c, the first wax oil obtained in step b is subjected to wax oil hydrogenation reaction, and the obtained reaction product is separated to obtain hydrogenated wax oil; d, obtained in step c The hydrogenated wax oil is subjected to the first catalytic cracking reaction or the second catalytic cracking reaction described in the step b.
  • the method of the present invention may further comprise the step of: reacting the second catalytic cracking described in step d
  • the second wax oil obtained in the above is subjected to the wax oil hydrogenation reaction described in the step c.
  • the yield of the hydrogenated residue is from 85 to 95% by weight, preferably from 87 to 93, based on the inferior feedstock oil. %, the properties of the hydroresin remain substantially constant.
  • the nature of the hydrorebase remains substantially constant means that at least one of the following conditions is satisfied:
  • the percentage change of the desulfurization rate of the inferior feedstock oil ( ⁇ desulfurization rate) is less than 20%;
  • the percentage change of the denitrification rate of the inferior feedstock oil (the delta denitrification rate) is less than 40%
  • ⁇ desulfurization rate [(maximum desulfurization rate - minimum desulfurization rate) / minimum desulfurization rate] * 100%;
  • ⁇ denitrification rate [(maximum denitrification rate - minimum denitrification rate) / minimum denitrification rate] * 100%;
  • ⁇ removal carbon rate [(maximum decarburization rate - minimum decarburization rate) / minimum decarburization rate] * 100%;
  • ⁇ demetallization rate [(maximum demetallization rate - minimum demetallization rate) / minimum demetallization rate] * 100%;
  • the maximum value and the minimum value refer to the maximum value and the minimum value in each batch, respectively.
  • “the nature of the hydrorebase remains substantially constant” means that the delta desulfurization rate is less than 20%; the delta denitrification rate is less than 40%; the delta decarburization rate is less than 40%; and the delta demetallization The rate is less than 20%.
  • “the nature of the hydrorebase remains substantially constant” means that the delta desulfurization rate is less than 10%; the delta denitrification rate is less than 20%; the delta decarburization rate is less than 20%; The metal rate is less than 10%.
  • the severity of the hydrogenation reaction is increased, so that the properties of the hydrorebricated oil and the properties of the hydrorebricated oil at the initial stage of operation are basically keep constant.
  • the density of the hydrorebase is increased by more than 0.001 to 0.005 g/cm 3 or the residual carbon is increased by more than 0.1% to 0.5%, the reaction severity of the hydrotreating unit is increased.
  • the growth rate of the density of the hydroresin exceeds 0.005 g/cm 3 /(1000 hours), and/or when the growth rate of the carbon residue of the hydro residue exceeds 0.5% by weight/(1000 hours)
  • increase the severity of the hydrogenation reaction for example, increase the reaction temperature by 2-10 ° C / (1000 hours) or reduce the liquid hourly space velocity by 0.1-0.5 h -1 / (1000 hours).
  • the low severity hydrogenation reaction can control the reaction temperature over time during the entire reaction period, such as increasing the reaction temperature at a constant rate (temperature increase rate is 10-50 ° C / (8000 hours)), or can divide the entire operation cycle into n Stages (n is an integer greater than 1), each phase maintains its own response Temperature, and the temperature difference of any two consecutive stages (reaction temperature at the end of the latter stage minus the reaction temperature at the end of the previous stage) is 10-50 ° C / (n-1); wherein the low severity hydrogenation The reaction temperature of the reaction is 350-370 ° C in the 0-1000 h operating period.
  • the reaction temperature means the volume average temperature of the reactor for the hydrogenation reaction and the outlet temperature of the reactor for the catalytic cracking reaction unless otherwise stated.
  • the inventors of the present invention have unexpectedly found that when the inferior feedstock oil is subjected to a hydrogenation reaction, when the yield of the hydrorebase residue is between 85 and 95% by weight, the metal on the catalyst increases as the operation time of the apparatus increases.
  • the coke deposition amount is increasing more and more slowly, and the operation cycle of the residue hydrogenation reaction device can be remarkably improved.
  • the present invention refers to such a hydrogenation reaction as a low-calorie hydrogenation reaction.
  • the invention combines low-quality feedstock oil in a low-calorie hydrotreating unit for low-status hydrogenation reaction, and dynamically adjusts the reaction conditions to obtain a hydrogen residue yield and impurities obtained after separation and fractionation of the product.
  • the rate of removal is relatively stable. Specifically, as the operating time of the apparatus increases, as the yield of the hydrocracking oil increases and/or the rate of removal of impurities decreases, the severity of the hydrogenation reaction is increased (e.g., the reaction temperature is increased).
  • the reaction conditions of the low severity hydrogenation reaction may be: hydrogen partial pressure of 8-20 MPa, preferably 9-16 MPa, reaction temperature of 330-420 ° C, preferably 350 ° C-400. °C, the liquid hourly space velocity is 0.1-1.5 hours -1 , preferably 0.2-1.0 hours -1 , and the total hydrogen oil volume ratio is 200-1500 standard cubic meters/millimeter, preferably 500-1000 standard cubic meters/m3.
  • the low severity hydrogenation reaction has a reaction temperature of 350-370 ° C at the initial stage of operation (for example, 0-1000 h).
  • the main purpose of using a low severity hydrogenation reaction is to control the desulfurization rate, denitrification rate, carbon removal rate and demetallization rate of inferior feedstock oil to a low level.
  • the desulfurization rate of the inferior feedstock oil can be controlled to be 50 to 95% by weight, preferably 65 to 85% by weight
  • the denitrification rate is 10 to 70% by weight, preferably 25 to 45% by weight
  • the carbon is removed.
  • the ratio is 10 to 70% by weight, preferably 25 to 45% by weight
  • the demetallization ratio is 50 to 95% by weight, preferably 65 to 80% by weight.
  • the low severity hydrogenation reaction is carried out in a fixed bed reactor.
  • the low severity hydrogenation reaction is carried out in the presence of a hydrogenation catalyst.
  • the hydrogenation catalyst used for the low-calorie hydrogenation reaction may sequentially include a hydrogenation protecting agent, a hydrodemetallization catalyst, a hydrodesulfurization catalyst, and addition according to the flow direction of the reactant. Hydrogen denitrification decarburization catalyst.
  • the hydrogenation protecting agent and the hydrodemetallization catalyst account for 20%-70%, for example 30%-50%, based on the total weight of the hydrogenation catalyst; the hydrodesulfurization catalyst accounts for 20%-70%, for example 40 %-60%; hydrodenitrogenation decarburization catalyst accounts for 0%-60%, for example 10%-40%, and hydrogenation protecting agent, hydrodemetallization catalyst, hydrodesulfurization catalyst, and hydrogenation
  • the sum of the denitrification and decarburization catalysts was 100% by weight.
  • the hydrogenation catalysts are those conventionally used in the art.
  • the hydrodemetallization catalyst comprises 30% by weight or more based on the total weight of the hydrogenation catalyst.
  • the inferior feedstock oil is conventionally used in the art.
  • the inferior feedstock oil may be petroleum hydrocarbons and/or other mineral oils, wherein the petroleum hydrocarbons may be selected from the group consisting of atmospheric gas oil, vacuum gas oil, atmospheric residue, vacuum residue, hydrocrack, Coking gas oil, deasphalted oil, and any combination thereof, other mineral oils may be selected from the group consisting of coal and natural gas derived liquid oils, oil sands oils, tight oils, shale oils, and any combination thereof.
  • the inferior feedstock oil may satisfy: (1) a density of 910-1000 kg/m 3 at 20 ° C; and/or (2) a carbon residue specific gravity of 4-15 wt%; and / Or (3) the metal (Ni + V) content is 12-600 ppm.
  • the inferior feedstock oil satisfies: (1) a density of 980-1000 kg/m 3 at 20 ° C; and/or (2) a carbon residue specific gravity of 10-15 wt%; and/or (3)
  • the metal (Ni + V) content is 60-600 ppm.
  • the first catalytic cracking reaction is a high-selective catalytic cracking process, which does not pursue the highest single-pass conversion rate of the feedstock oil, and controls the conversion rate to an appropriate level, thereby effectively improving dry gas and coke.
  • Selectivity while producing a relatively large amount of catalytic cracking wax oil for further hydrotreating.
  • the process can effectively compensate for the problem of insufficient processing depth of low-grade residue hydrotreating for inferior raw materials and optimize product distribution.
  • the first catalytic cracking reaction may include the following steps: (1) performing a first cracking reaction between the preheated hydrocrack and the first regenerated catalytic cracking catalyst in a lower portion of the first catalytic cracking reactor, and the obtained reaction The product is separated to obtain a first cracked product and a first semi-regenerated catalytic cracking catalyst; the first regenerated catalytic cracking catalyst has a micro-reverse evaluation activity of 35-60; (2) the first cracked product obtained in the step (1) And the first semi-regenerated catalytic cracking catalyst is subsequently subjected to a first re-catalytic conversion reaction in an upper portion of the first catalytic cracking reactor, and the obtained reaction product is separated and fractionated to obtain a first dry gas, a first liquefied gas, First gasoline, first diesel oil and first wax oil.
  • the lower portion and the upper portion of the first catalytic cracking reactor are demarcated from a position between the first 1/3 portion of the reactor and the first 2/3 portion (in the direction in which the reactant flows); in a preferred implementation
  • the lower part refers to the first 1/2 part of the length of the reactor
  • the upper part refers to the last 1/2 part of the length of the reactor.
  • the first cracking reaction is mainly a macromolecular cracking reaction
  • the first recatalytic conversion reaction is mainly a reaction such as selective cracking, selective hydrogen transfer and isomerization.
  • the first cracking reaction can be carried out under the following conditions: a reaction temperature of 530-620 ° C, a weight hourly space velocity of 30-180 hr -1 , a ratio of the ratio of the agent to the oil (weight ratio of the catalyst to the feedstock oil) of 4-12, water
  • the oil ratio (weight ratio of steam to feedstock oil) is 0.03-0.3, and the reaction pressure is 130 kPa-450 kPa.
  • the first recatalytic conversion reaction is carried out under the following conditions: a reaction temperature of 460 ° C - 520 ° C, a weight hourly space velocity of 20-100 hr -1 , a ratio of the agent to the oil of 3 - 15 , a ratio of water to oil (steam and raw materials)
  • the oil weight ratio is 0.03-0.3, and the reaction pressure is 130 kPa-450 kPa.
  • the first wax oil has a hydrogen content of 10.5-15% by weight; based on the hydrogenated residue, the yield of the first wax oil is 15-50% by weight, preferably 30-45% by weight.
  • the secondary processing wax oil may be subjected to the wax oil hydrogenation reaction described in the step c together with the first wax oil to increase the source of the second catalytic cracking raw material.
  • the secondary processing wax oil may be selected from the group consisting of coking wax oil, deasphalted oil, catalytic cracking wax oil produced by other devices, and any combination thereof.
  • the catalytic cracking wax oil is not limited to the first wax oil and the second wax oil of the present invention, and may be derived from other catalytic cracking devices.
  • the wax oil hydrogenation reaction can be carried out under the following conditions: a reaction pressure of 5.0 to 20.0 MPa, preferably 6.0 to 15.0 MPa, and a reaction temperature of 300 to 430 ° C, preferably 320 to 390 °C, the liquid hourly space velocity may be 0.2-5.0 hours -1 , preferably 0.3-2.5 hours -1 , and the hydrogen oil volume ratio may be 200-1800 standard cubic meters/m3, preferably 400-1100 standard cubic meters/m3. .
  • the wax oil hydrogenation reaction is carried out in a fixed bed reactor in the presence of a hydrogenation catalyst.
  • the hydrogenation catalyst used for the hydrogenation reaction of the wax oil may sequentially include a hydrogenation protecting agent, a hydrodemetallization desulfurization catalyst, and a hydrotreating catalyst in accordance with the flow direction of the reactant.
  • the hydrogenation protecting agent comprises from 0 to 30% by weight, for example from 5 to 20% by weight, based on the total weight of the hydrogenation catalyst
  • the hydrodemetallization desulfurization catalyst comprises from 5 to 35% by weight, for example from 10 to 25% by weight.
  • the hydrotreating catalyst accounts for 35% to 95% by weight, for example, 55 to 85% by weight, and the sum of the hydrogenation protecting agent, the hydrodemetallization desulfurization catalyst, and the hydrotreating catalyst is 100% by weight.
  • the hydrogenation catalysts are those conventionally used in the art.
  • the second catalytic cracking reaction can be carried out under conventional conditions in the art, for example, the reaction temperature is from 450 ° C to 620 ° C, the weight hourly space velocity is from 1 to 100 hours -1 , and the ratio of the agent to the oil is from 1 to 25 The water to oil ratio is 0.03-0.3.
  • the second catalytic cracking reaction may also adopt a high selective catalytic cracking process.
  • the second catalytic cracking reaction may include the following steps: (1) preheating the hydrogenated wax oil with the second regenerated catalytic cracking
  • the catalyst is subjected to a second cracking reaction in a lower portion of the second catalytic cracking reactor, and the obtained reaction product is separated to obtain a second cracking product and a second semi-regenerated catalytic cracking catalyst; (2) the second obtained in the step (1)
  • the cracked product and the second semi-regenerated catalytic cracking catalyst are then subjected to a second recatalytic conversion reaction in the upper portion of the second catalytic cracking reactor, and the obtained reaction product is separated, fractionated to obtain a second dry gas, and the second liquefaction Gas, second gasoline, second diesel and second wax oil.
  • the lower portion and the upper portion of the second catalytic cracking reactor are demarcated by a position between the front 1/3 portion and the first 2/3 portion of the reactor (in the direction in which the reactant flows); in a preferred implementation
  • the lower part refers to the first 1/2 part of the length of the reactor
  • the upper part refers to the last 1/2 part of the length of the reactor.
  • the hydrogenation catalyst may comprise at least one metal component selected from Group VIII and/or at least one selected from Group VIB (as an active ingredient), and alumina and/or silica (as a carrier).
  • the catalytic cracking catalyst may comprise a zeolite (as an active component), preferably a medium pore zeolite and/or optionally a large pore zeolite; wherein the medium pore zeolite may be selected from the ZSM series and/or the ZRP series.
  • the catalytic cracking reactor can be selected from the group consisting of a riser, a fluidized bed, and combinations thereof.
  • the hydrogenation reactor can be selected from the group consisting of a fixed bed, a suspended bed, a bubbling bed, a moving bed, and combinations thereof (preferably a fixed bed).
  • the number of the catalytic cracking reactor and the hydrogenation reactor may be 1, 2, 3 or more. When the number of reactors is 2, the reactors may be connected in series or in parallel; When the number of the devices is three or more, the reactors may be connected in series, in parallel, or in a mixture.
  • the inferior feedstock oil from line 9 is mixed with the new hydrogen and recycle hydrogen mixed gas from line 11, and then enters the low-chassis hydrogenation reactor 1, and is subjected to impurity removal and hydrodemetallization under low severity hydrogenation reaction conditions. Hydrodesulfurization, hydrodenitrogenation and hydrocracking reaction.
  • the resulting product is passed through line 13 to a separation unit 2 of a low severity hydrogenation reaction product which is passed via line 14 to a recycle gas treatment system 3, via line 15 to a recycle hydrogen compressor 4, and then via line 16
  • the new hydrogen from line 10 is mixed.
  • the liquid phase stream from separation unit 2 enters hydrogenation fractionation unit 5 via line 17 to provide hydrogenation gas (line 18), hydrogenated naphtha (line 19), hydrogenated diesel (line 20), and hydrocracking, respectively.
  • the first wax oil is mixed with the mixed hydrogen from line 12 via line 29 and sent to the wax oil hydrogenation reaction. 7.
  • the stream leaving the wax oil hydrogenation reactor 7 is separated in a separation unit 8 of the wax oil hydrogenation product, and the resulting hydrogen-rich phase stream is mixed with the hydrogen-rich phase stream from line 14 via line 23 and sent to a recycle gas treatment system. 3.
  • the resulting liquid phase stream (hydrogenated wax oil) is mixed via line 24 with the hydrocedure from line 21 and sent to the first catalytic cracking reactor 6.
  • the inferior feedstock oil used in the examples and comparative examples was a mixed residue of a vacuum residue and an atmospheric residue, and the properties thereof are shown in Table 1.
  • the catalysts used in the examples and comparative examples were produced by Sinopec Catalyst Branch.
  • Example 1 provides a regulatable low severity hydrogenation reaction of the present invention in which the reaction temperature and liquid hourly space velocity are adjusted stepwise with reaction time, while the hydrogen oil volume ratio and hydrogen partial pressure are maintained at 800 standard cubic meters/ Cubic meters and 15 MPa.
  • the cut point of the hydrocluster was 350 °C.
  • the hydrogenation test was carried out on a continuous high-pressure fixed-bed pilot plant consisting of three reactors connected in series, each containing a 5:45:50 by volume of hydrogenation protecting agent (RG-10A), hydrodemetallization. Catalyst (RDM-2B), hydrodesulfurization catalyst (RMS-1B). At the time of the test, the pilot plant was in the initial stage of operation and the operation time was less than 50 hours.
  • RPM-2B hydrodesulfurization catalyst
  • the catalytic cracking test was carried out on a medium-sized catalytic cracking unit using a riser reactor using an MLC-500 catalyst.
  • the wax oil hydrotreating test was carried out on a fixed bed hydrogenation reactor packed with a hydrogenation protection catalyst A (RG-30A) in a volume ratio of 4:4:15:77, and hydrogenation protection.
  • Catalyst B (RG-30B), hydrodemetallization desulfurization catalyst (RMS-30) and hydrotreating catalyst (RDA-1).
  • Comparative Example 1 is a conventional residue hydrogenation test, and the test apparatus and test materials were the same as in Example 1. The difference is that the temperature of the hydrogenation reaction of the inferior feedstock oil and the liquid hourly space velocity are constant at 390 ° C and 0.25 h -1 , respectively .
  • Example 1 The reaction conditions and reaction results of Example 1 and Comparative Example 1 are shown in Table 2.
  • Example 3 The reaction product obtained in 5000 to 500 hours of Example 1 (see Table 3) was the subject of subsequent studies. Hydrogen residue is used as a feedstock oil for the first catalytic cracking reaction. Hydrogen residue undergoes first catalytic cracking After the reaction and separation, the first dry gas, the first liquefied gas, the first gasoline, the first diesel oil and the first wax oil are obtained. The first wax oil cut point is 330 ° C, accounting for 33.23% of the feed amount. The first wax oil is sent to the wax oil hydrotreating unit, and the obtained product is separated by gas and liquid, and the liquid hydrogenated wax oil passes through the second catalytic cracking reaction to obtain the second dry gas, the second liquefied gas, and the second Gasoline, second diesel and second wax. The second wax oil is sent to the wax oil hydrotreating unit.
  • Comparative Example 2 is a combination of existing residue hydrogenation-heavy oil catalytic cracking.
  • the reaction product in 5000-5500 hours of Comparative Example 1 (see Table 3) was used as a subject for subsequent studies.
  • the hydrogen residue is subjected to reaction, separation and fractionation to obtain dry gas, liquefied gas, gasoline, diesel oil, oil slurry and coke.
  • the operating conditions are shown in Table 4, and the product distribution is shown in Table 5.
  • Comparative Example 3 The process flow and reaction conditions of Comparative Example 3 were substantially the same as in Example 2, except that in Comparative Example 3, the reaction product of 5000-5500 hours of Comparative Example 1 (see Table 3) was used as the object of the subsequent study. .
  • the operating conditions are shown in Table 4, and the product distribution is shown in Table 5.
  • Table 3 Properties and product distribution of the partially hydrocedure obtained in Example 1 and Comparative Example 1 (the hydrorebase obtained in Example 1 was used in Example 2, and the hydrorebase obtained in Comparative Example 1 was used for Proportion 2 and Comparative 3)

Abstract

一种劣质原料油的处理方法，包括：a、将劣质原料油进行低苛刻度加氢反应，所得到的反应产物经分离得到气体、加氢石脑油、加氢柴油和加氢渣油；其中，在低苛刻度加氢反应中，以所述的劣质原料油为基准，所述加氢渣油的产率为85-95重量％，所述加氢渣油的性质基本保持恒定；b、将步骤a中所得的加氢渣油进行第一催化裂化反应，所得的反应产物经分离得到第一干气、第一液化气、第一汽油、第一柴油和第一蜡油；c、将步骤b中所得的第一蜡油进行加氢反应，所得到的反应产物经分离得到加氢蜡油；d、将步骤c中所得的加氢蜡油进行步骤b中所述的第一催化裂化反应或第二催化裂化反应。该方法可延长劣质原料油的加氢处理装置的运行周期并降低化学氢耗。

Description

一种劣质原料油的处理方法 技术领域
本发明涉及一种劣质原料油的处理方法。
背景技术
CN101210200B公开了一种渣油加氢处理与催化裂化组合工艺方法。渣油、脱除固体杂质的催化裂化重循环油、任选的馏分油和任选的催化裂化油浆的蒸出物一起进入渣油加氢处理装置，所得的加氢渣油与任选的减压瓦斯油一起进入催化裂化装置，得到各种产品；将脱除固体杂质的催化裂化重循环油循环至渣油加氢处理装置；将催化裂化油浆进行蒸馏分离，而催化裂化油浆的蒸出物可循环至渣油加氢处理装置。
CN102344829A公开了一种渣油加氢处理、催化裂化重油加氢和催化裂化的组合方法。渣油加氢反应器所得的液相物流分馏得到的渣油加氢尾油作为催化裂化的原料进入催化裂化装置，催化裂化产物中的催化裂化重油与渣油加氢反应器所得的气相物流混合进入催化裂化重油加氢反应器，加氢后的催化裂化重油循环回催化裂化装置。
仍然需要新的劣质原料油的处理方法，其具有延长的加氢装置的运行周期、降低的化学氢耗、和提高的液体产品收率。
发明内容
本发明的目的是提供一种新的劣质原料油的处理方法，该方法可以延长加氢装置的运行周期，并且具有低化学氢耗和高液体产品收率。
本发明提供了一种劣质原料油的处理方法，该方法包括：
a、将劣质原料油进行低苛刻度加氢反应，所得到的反应产物经分离得到气体、加氢石脑油、加氢柴油和加氢渣油；其中，在低苛刻度加氢反应中，以所述的劣质原料油为基准，所述加氢渣油的产率为85-95重量％，所述加氢渣油的性质基本保持恒定；
b、将步骤a中所得的加氢渣油进行第一催化裂化反应，所得到的反应产物经分离得到第一干气、第一液化气、第一汽油、第一柴油和第一蜡油；
c、将步骤b中所得的第一蜡油进行蜡油加氢反应，所得到的反应产物 经分离得到加氢蜡油；
d、将步骤c中所得的加氢蜡油进行步骤b中所述的第一催化裂化反应或第二催化裂化反应。
在一种实施方案中，该方法还包括步骤e：将在步骤d中所述的第二催化裂化反应中所得的第二蜡油进行步骤c中所述的蜡油加氢反应。
在一种实施方案中，在步骤a中所述的低苛刻度加氢反应中，以所述的劣质原料油为基准，所述加氢渣油的产率为87-93重量％，所述加氢渣油的性质基本保持恒定。
当加氢渣油的性质不期望地变化(例如，密度增加或残炭值增加)时，提高加氢反应的苛刻度，使加氢渣油的性质与运转初期(例如0-1000h)的加氢渣油的性质基本保持恒定。例如，当加氢渣油的密度的增速超过0.005g/cm3/(1000小时)时，和/或当加氢渣油的残炭值的增速超过0.5重量％/(1000小时)时，提高加氢反应的苛刻度(例如，以2-10℃/(1000小时)来提高反应温度或以0.1-0.5h-1/(1000小时)来降低液时空速)。
在一种实施方案中，在步骤a中，所述的劣质原料油的脱硫率为50-95重量％，脱氮率为10-70重量％，脱残炭率为10-70重量％，脱金属率为50-95重量％。
在一种实施方案中，低苛刻度加氢反应的反应条件包括：氢分压为8-20兆帕，反应温度为330-420℃，液时空速为0.1-1.5小时-1，总氢油体积比为200-1500标准立方米/立方米。
在一种实施方案中，所述的低苛刻度加氢反应在运转初期(例如0-1000h)的反应温度为350-370℃，例如350-360℃、350-355℃，或者例如350℃、351℃、352℃、353℃、354℃、355℃、356℃、357℃、358℃、359℃、360℃、361℃、362℃、363℃、364℃、365℃、366℃、367℃、368℃、369℃或者370℃。
在一种实施方案中，所述的低苛刻度加氢反应在固定床反应器中在加氢催化剂的存在下进行。按照加氢催化剂的功能，按照反应物的流动方向，用于所述的低苛刻度加氢反应的加氢催化剂可以依次包括加氢保护剂、加氢脱金属催化剂、加氢脱硫催化剂、和加氢脱氮脱残炭催化剂。优选地，以加氢催化剂的总重量为基准，加氢保护剂和加氢脱金属催化剂占20％-70％，例如30％-50％；加氢脱硫催化剂占20％-70％，例如40％-60％；加氢脱氮脱残炭 催化剂占0％-60％，例如10％-40％，并且加氢保护剂、加氢脱金属催化剂、加氢脱硫催化剂、和加氢脱氮脱残炭催化剂的总和是100重量％。所述加氢催化剂为本领域中常规使用的那些。在一种优选的实施方案中，以加氢催化剂的总重量为基准，加氢脱金属催化剂占30重量％或更高。
在一种实施方案中，所述的劣质原料油为石油烃和/或其他矿物油，其中石油烃选自常压瓦斯油、减压瓦斯油、常压渣油、减压渣油、加氢渣油、焦化瓦斯油、脱沥青油、和其任意的组合，其他矿物质油选自煤与天然气衍生的液体油、油砂油、致密油、页岩油、和其任意的组合。
在一种实施方案中，所述的劣质原料油满足：(1)在20℃时的密度为910-1000千克/立方米；和/或(2)残炭比重为4-15重量％；和/或(3)金属(Ni+V)含量为12-600ppm。在一种优选的实施方案中，所述的劣质原料油满足：(1)在20℃时的密度为980-1000千克/立方米；和/或(2)残炭比重为10-15重量％；和/或(3)金属(Ni+V)含量为60-600ppm。
在一种实施方案中，所述的第一催化裂化反应包括如下步骤：(1)将预热的加氢渣油与第一再生催化裂化催化剂在第一催化裂化反应器的下部进行第一裂化反应，所得到的反应产物经分离得到第一裂化产物和第一半再生催化裂化催化剂；所述第一再生催化裂化催化剂的微反评价活性为35-60；(2)将步骤(1)中所得的第一裂化产物与所述的第一半再生催化裂化催化剂随后在第一催化裂化反应器的上部进行第一再催化转化反应，并且所得到的反应产物经过分离、分馏得到第一干气、第一液化气、第一汽油、第一柴油和第一蜡油。所述的第一催化裂化反应器的下部与上部由反应器的前1/3部分与前2/3部分(以反应物流动的方向)之间的某一位置分界；在一种优选的实施方案中，下部是指反应器长度的前1/2部分，上部是指反应器长度的后1/2部分。
在一种实施方案中，所述的第一裂化反应在下述条件进行：反应温度为530-620℃，重时空速为30-180小时-1，剂油比为4-12，水油比为0.03-0.3，反应压力为130千帕-450千帕；所述的第一再催化转化反应在下述条件进行：反应温度为460℃-520℃，重时空速为20-100小时-1，剂油比为3-15，水油比为0.03-0.3，反应压力为130千帕-450千帕。
在一种实施方案中，在所述的第一催化裂化反应中，所述的第一蜡油的氢含量为10.5-15重量％；以所述的加氢渣油为基准，所述的第一蜡油的产 率为15-50重量％。
在一种实施方案中，将二次加工蜡油与所述的第一蜡油一起进行所述的蜡油加氢反应；所述的二次加工蜡油选自焦化蜡油、脱沥青油、其它装置所产的催化裂化蜡油、和其任意的组合。
在一种实施方案中，所述的蜡油加氢反应在固定床反应器中在加氢催化剂的存在下进行。按照加氢催化剂的功能，按照反应物的流动方向，用于所述的蜡油加氢反应的加氢催化剂可以依次包括加氢保护剂、加氢脱金属脱硫催化剂和加氢处理催化剂。优选地，以加氢催化剂的总重量为基准，加氢保护剂占0-30重量％，例如5-20重量％、加氢脱金属脱硫催化剂占5-35重量％，例如10-25重量％；和加氢处理催化剂占35％-95重量％，例如55-85重量％，并且加氢保护剂、加氢脱金属脱硫催化剂和加氢处理催化剂的总和是100重量％。所述加氢催化剂为本领域中常规使用的那些。
在一种实施方案中，所述的蜡油加氢反应在下述条件进行：反应压力为5.0-20.0兆帕，反应温度为300-430℃，液时空速为0.2-5.0小时-1，氢油体积比为200-1800标准立方米/立方米。
在一种实施方案中，所述的第二催化裂化反应在下述条件进行：反应温度为450℃-620℃，重时空速为1-100小时-1，剂油比为1-25，水油比为0.03-0.3。
在一种实施方案中，所述的第二催化裂化反应包括如下步骤：(1)将预热的加氢蜡油与第二再生催化裂化催化剂在第二催化裂化反应器的下部进行第二裂化反应，所得到的反应产物经分离得到第二裂化产物和第二半再生催化裂化催化剂；(2)将步骤(1)中所得的第二裂化产物与所述的第二半再生催化裂化催化剂随后在第二催化裂化反应器的上部进行第二再催化转化反应，并且所得到的反应产物经过分离、分馏得到第二干气、第二液化气、第二汽油、第二柴油和第二蜡油。所述的第二催化裂化反应器的下部与上部由反应器的前1/3部分与前2/3部分(以反应物流动的方向)之间的某一位置分界；在一种优选的实施方案中，下部是指反应器长度的前1/2部分，上部是指反应器长度的后1/2部分。
在一种实施方案中，所述的第二裂化反应在下述条件进行：反应温度为530-620℃，重时空速为30-180小时-1，剂油比为4-12，水油比为0.03-0.3，反应压力为130千帕-450千帕；所述的第二再催化转化反应的条件为：反应 温度为460℃-520℃，重时空速为20-100小时-1，剂油比为3-15，水油比为0.03-0.3，反应压力为130千帕-450千帕。
本发明还提供了下述技术方案：
技术方案1、一种重质原料油的处理方法，该方法包括：
a、将重质原料油进行浅度加氢反应，得到加氢气体、加氢石脑油、加氢柴油和加氢渣油；其中，以所述重质原料油为基准，控制所述加氢渣油的产率为85-95重％；
b、将步骤a中所得加氢渣油与催化裂化催化剂接触并进行第一催化裂化反应，得到第一干气、第一液化气、第一汽油、第一轻循环油、第一蜡油和外甩油浆；其中，所述催化裂化催化剂的微反评价活性为40-55；
c、将步骤b中所得第一蜡油过滤后进行蜡油加氢反应，得到加氢蜡油；将步骤b中所得外甩油浆进行步骤b中所述第一催化裂化反应；
d、将步骤c中所得加氢蜡油进行第二催化裂化反应或所述第一催化裂化反应。
技术方案2、根据技术方案1的方法，该方法还包括步骤e：将步骤d中所述第二催化裂化反应所得第二蜡油进行步骤c中所述蜡油加氢反应。
技术方案3、根据技术方案1的方法，其中，步骤a中，以所述重质原料油为基准，控制所述加氢渣油的产率为87-93重％。
技术方案4、根据技术方案1的方法，其中，步骤a中，控制所述重质原料油的脱硫率为50-95重％，脱氮率为20-70重％，脱残炭率为20-70重％，脱金属率为50-90重％。
技术方案5、根据技术方案1的方法，其中，所述浅度加氢反应的条件为：氢分压为10-20兆帕，反应温度为320-420℃，液时体积空速为0.2-1.0小时-1，总氢油体积比为300-1500标准立方米/立方米。
技术方案6、根据技术方案1的方法，其中，所述重质原料油为石油烃和/或其他矿物油，其中石油烃选自常压瓦斯油、减压瓦斯油、常压渣油、减压渣油、加氢渣油、焦化瓦斯油和脱沥青油中的至少一种，其他矿物质油选自煤与天然气衍生的液体油、油砂油、致密油和页岩油中的至少一种。
技术方案7、根据技术方案1的方法，其中，所述重质原料油满足：在20℃时的密度为910-1000千克/立方米和/或残炭比重为4-15重％和/或金属含量为12-600ppm。
技术方案8、根据技术方案1的方法，其中，步骤b中所述第一催化裂化的反应条件为：反应温度为450-670℃，重时空速为10-100小时-1，再生催化剂与原料油重量比为1-30，水蒸气与原料的重量比为0.03-1.0。
技术方案9、根据技术方案1的方法，其中，控制所述第一蜡油的氢含量为9.0-13.0重％；以步骤b中所述加氢渣油为基准，控制所述第一蜡油的产率为15-50重％。
技术方案10、根据技术方案1的方法，其中，步骤b中所得外甩油浆的固含量小于6克/升，在20℃时的密度为920-1150千克/立方米。
技术方案11、根据技术方案1的方法，其中，步骤c中过滤后的所述第一蜡油的固体含量小于10ppm。
技术方案12、根据技术方案1的方法，其中，将二次加工蜡油与第一蜡油一起进行步骤c中所述蜡油加氢反应；所述二次加工蜡油为选自焦化蜡油、脱沥青油和其它装置所产催化裂化蜡油中的至少一种。
技术方案13、根据技术方案1的方法，其中，步骤c中所述蜡油加氢反应在固定床反应器中进行；按反应物流向，在所述固定床反应器内依次填装加氢保护剂、加氢脱金属脱硫剂和加氢处理催化剂。
技术方案14、根据技术方案1的方法，其中，步骤c中所述蜡油加氢反应的条件为：反应压力为6.0-18.0兆帕，反应温度为270-420℃，体积空速为0.2-1.0小时-1，氢油体积比为200-1800标准立方米/立方米。
技术方案15、根据技术方案1的方法，其中，步骤d中所述第二催化裂化反应的条件为：反应温度为450℃-620℃，重时空速为1-100小时-1，剂油比为1-25，水油比为0.03-0.3。
技术方案16、一种劣质原料油的处理方法，该方法包括：
a、将劣质原料油进行浅度加氢反应，得到气体、加氢石脑油、加氢柴油和加氢渣油；其中，以所述劣质原料油为基准，控制所述加氢渣油的产率为85-95重％；
b、将步骤a中所得加氢渣油进行第一催化裂化反应，得到第一干气、第一液化气、第一汽油、第一柴油和第一蜡油；
c、将步骤b中所得第一蜡油进行蜡油加氢反应，得到加氢蜡油；
d、将步骤c中所得加氢蜡油进行第二催化裂化反应，得到第二干气、第二液化气、第二汽油、第二柴油和第二蜡油。
技术方案17、根据技术方案16的方法，该方法还包括步骤e：将步骤d中所得第二蜡油进行步骤c中所述蜡油加氢反应。
技术方案18、根据技术方案16的方法，其中，步骤a中，以所述劣质原料油为基准，控制所述加氢渣油的产率为87-93重％。
技术方案19、根据技术方案16的方法，其中，步骤a中，控制所述劣质原料油的脱硫率为50-95重％，脱氮率为10-70重％，脱残炭率为10-70重％，脱金属率为50-95重％。
技术方案20、根据技术方案16的方法，其中，所述浅度加氢反应的条件为：氢分压为8-20兆帕，反应温度为330-420℃，液时体积空速为0.1-1.5小时-1，总氢油体积比为200-1500标准立方米/立方米。
技术方案21、根据技术方案16的方法，其中，所述劣质原料油为石油烃和/或其他矿物油，其中石油烃选自常压瓦斯油、减压瓦斯油、常压渣油、减压渣油、加氢渣油、焦化瓦斯油和脱沥青油中的至少一种，其他矿物质油选自煤与天然气衍生的液体油、油砂油、致密油和页岩油中的至少一种。
技术方案22、根据技术方案16的方法，其中，所述劣质原料油在20℃时的密度为920-1100千克/立方米，残炭比重为8-20重％。
技术方案23、根据技术方案16的方法，其中，所述将步骤a中所得加氢渣油进行第一催化裂化反应包括如下步骤：
(1)、将预热的所述加氢渣油与第一再生催化裂化催化剂在第一催化裂化反应器下部一起进行第一裂化反应，得到第一裂化产物和第一半再生催化裂化催化剂；
(2)、将步骤(1)中所得第一裂化产物与所述第一半再生催化裂化催化剂随后在第一催化裂化反应器上部进行第一再催化转化反应，并经过分离、分馏得到所述第一千气、第一液化气、第一汽油、第一柴油和第一蜡油。
技术方案24、根据技术方案23的方法，其中，步骤(1)中所述第一裂化反应的条件为：反应温度为530-620℃，重时空速为30-180小时-1，剂油比为4-12，水油比为0.03-0.3，反应压力为130千帕-450千帕；步骤(2)中所述第一再催化转化反应的条件为：反应温度为460℃-520℃，重时空速为20-100小时-1，剂油比为3-15，水油比为0.03-0.3，反应压力为130千帕-450千帕。
技术方案25、根据技术方案16的方法，其中，控制所述第一蜡油的氢含量为10.5-15重％；以步骤b中所述加氢渣油为基准，控制所述第一蜡油 的产率为15-50重％。
技术方案26、根据技术方案16的方法，其中，将二次加工蜡油与第一蜡油一起进行步骤c中所述蜡油加氢反应；所述二次加工蜡油为选自焦化蜡油、脱沥青油和其它装置所产催化裂化蜡油中的至少一种。
技术方案27、根据技术方案16的方法，其中，步骤c中所述蜡油加氢反应在固定床反应器中进行；按反应物流向，在所述固定床反应器内依次填装加氢保护剂、加氢脱金属脱硫剂和加氢处理催化剂。
技术方案28、根据技术方案16的方法，其中，步骤c中所述蜡油加氢反应的条件为：反应压力为5.0-20.0兆帕，反应温度为300-430℃，体积空速为0.2-5.0小时-1，氢油体积比为200-1800标准立方米/立方米。
技术方案29、根据技术方案16的方法，其中，步骤d中所述第二催化裂化反应的条件为：反应温度为450℃-620℃，重时空速为1-100小时-1，剂油比为1-25，水油比为0.03-0.3。
技术方案30、根据技术方案16的方法，其中，所述将步骤c中所得加氢蜡油进行第二催化裂化反应包括如下步骤：
(α)、将预热的所述加氢蜡油与第二再生催化裂化催化剂在第二催化裂化反应器下部一起进行第二裂化反应，得到第二裂化产物和第二半再生催化裂化催化剂；
(β)、将步骤(α)中所得第二裂化产物与所述第二半再生催化裂化催化剂随后在第二催化裂化反应器上部进行第二再催化转化反应，并经过分离、分馏得到所述第二干气、第二液化气、第二汽油、第二柴油和第二蜡油。
技术方案31、根据技术方案30的方法，其中，步骤(α)中所述第二裂化反应的条件为：反应温度为530-620℃，重时空速为30-180小时-1，剂油比为4-12，水油比为0.03-0.3，反应压力为130千帕-450千帕；步骤(β)中所述第二再催化转化反应的条件为：反应温度为460℃-520℃，重时空速为20-100小时-1，剂油比为3-15，水油比为0.03-0.3，反应压力为130千帕-450千帕。
本发明还包括上述实施方案和/或技术方案的任意可能的组合。
通过降低劣质原料油的加氢反应的苛刻度并且控制催化裂化装置对加氢渣油的转化深度，能够提高加氢催化剂的寿命，显著延长加氢装置的运行周期，并且能够降低化学氢耗。本发明的其他特征和优点将在随后的具体实施方式部分予以详细说明。
附图说明
附图是用来提供对本发明的进一步理解，并且构成说明书的一部分，与下面的具体实施方式一起用于解释本发明，但并不构成对本发明的限制。在附图中：
图1是本发明劣质原料油的处理方法的流程示意图，其中：
1           低苛刻度加氢反应器
2           低苛刻度加氢反应产物的分离单元
3           循环气处理系统
4           循环氢压缩机
5           加氢分馏单元
6           第一催化裂化反应器
7           蜡油加氢反应器
8           蜡油加氢产物的分离单元
9-30        管线
具体实施方式
以下结合附图对本发明的具体实施方式进行详细说明。应当理解的是，此处所描述的具体实施方式仅用于说明和解释本发明，并不用于限制本发明。
本发明提供一种劣质原料油的处理方法，该方法包括：a、将劣质原料油进行低苛刻度加氢反应，所得到的反应产物经分离得到气体、加氢石脑油、加氢柴油和加氢渣油；其中，在低苛刻度加氢反应中，以所述的劣质原料油为基准，所述加氢渣油的产率为85-95重量％，优选87-93重量％，所述加氢渣油的性质基本保持恒定；b、将步骤a中所得的加氢渣油进行第一催化裂化反应，所得到的反应产物经分离得到第一干气、第一液化气、第一汽油、第一柴油和第一蜡油；c、将步骤b中所得的第一蜡油进行蜡油加氢反应，所得到的反应产物经分离得到加氢蜡油；d、将步骤c中所得的加氢蜡油进行步骤b中所述的第一催化裂化反应或第二催化裂化反应。
本发明的方法还可以包括步骤e：将在步骤d中所述的第二催化裂化反 应中所得的第二蜡油进行步骤c中所述的蜡油加氢反应。
根据本发明，在步骤a中所述的低苛刻度加氢反应中，以所述的劣质原料油为基准，所述加氢渣油的产率为85-95重量％，优选87-93重量％，所述加氢渣油的性质基本保持恒定。在低苛刻度加氢反应中，“加氢渣油的性质基本保持恒定”是指满足下述条件中的至少一种：
(1)劣质原料油的脱硫率的变化百分比(Δ脱硫率)小于20％；
(2)劣质原料油的脱氮率的变化百分比(Δ脱氮率)小于40％；
(3)劣质原料油的脱残炭率的变化百分比(Δ脱残炭率)小于40％；
(4)劣质原料油的脱金属率的变化百分比(Δ脱金属率)小于20％；
其中
Δ脱硫率＝【(最大脱硫率-最小脱硫率)/最小脱硫率]*100％；
Δ脱氮率＝[(最大脱氮率-最小脱氮率)/最小脱氮率]*100％；
Δ脱残炭率＝[(最大脱残炭率-最小脱残炭率)/最小脱残炭率]*100％；
Δ脱金属率＝[(最大脱金属率-最小脱金属率)/最小脱金属率]*100％；
在以上各式中，最大值和最小值分别是指各个批次中的最大值和最小值。在一种优选的实施方案中，“加氢渣油的性质基本保持恒定”是指Δ脱硫率小于20％；Δ脱氮率小于40％；Δ脱残炭率小于40％；并且Δ脱金属率小于20％。在一种最优选的实施方案中，“加氢渣油的性质基本保持恒定”是指Δ脱硫率小于10％；Δ脱氮率小于20％；Δ脱残炭率小于20％；并且Δ脱金属率小于10％。
当加氢渣油的性质不期望地变化(例如，密度增加或残炭值增加)时，提高加氢反应的苛刻度，使加氢渣油的性质与运转初期的加氢渣油的性质基本保持恒定。例如，以加氢渣油密度增加超过0.001-0.005g/cm3，或残炭增加超过0.1％-0.5％时，就增加加氢处理装置的反应苛刻度。又例如，当加氢渣油的密度的增速超过0.005g/cm3/(1000小时)时，和/或当加氢渣油的残炭值的增速超过0.5重量％/(1000小时)时，提高加氢反应的苛刻度(例如，以2-10℃/(1000小时)来提高反应温度或以0.1-0.5h-1/(1000小时)来降低液时空速)。
所述低苛刻度加氢反应可以在整个反应期间随时间控制反应温度，例如匀速增加反应温度(温度增加速率为10-50℃/(8000小时))，或者可以将整个操作周期平均分为n个阶段(n为大于1的整数)，每个阶段保持各自的反应 温度，并且任何两个连续阶段的温度差(后一阶段末端的反应温度减去前一阶段末端的反应温度)为10-50℃/(n-1)；其中所述的低苛刻度加氢反应的反应温度在0-1000h操作时期内为350-370℃。
在本发明中，除非另有说明，对于加氢反应来说，反应温度是指反应器的体积平均温度，对于催化裂化反应来说，反应温度是指反应器的出口温度。
本发明的发明人意外地发现，将劣质原料油进行加氢反应时，当加氢渣油的产率在85-95重量％之间时，随着装置运转时间的增加，催化剂上的金属和焦炭沉积量越来越慢地增加，渣油加氢反应装置的运转周期能够显著提高，本发明将此种加氢反应称为低苛刻度加氢反应。本发明将劣质原料油在低苛刻度加氢处理单元进行可调控性低苛刻度加氢反应，通过动态调变反应条件，使产物经分离、分馏后得到的加氢渣油产率和杂质脱除率相对稳定。具体来说，随装置运转时间增加，当加氢渣油产率增加和/或杂质脱除率降低时，增加加氢反应的苛刻度(例如提高反应温度)。
整体来说，所述低苛刻度加氢反应的反应条件可以为：氢分压为8-20兆帕，优选为9-16兆帕，反应温度为330-420℃，优选为350℃-400℃，液时空速为0.1-1.5小时-1，优选为0.2-1.0小时-1，总氢油体积比为200-1500标准立方米/立方米，优选为500-1000标准立方米/立方米。其中，所述的低苛刻度加氢反应在运转初期(例如0-1000h)的反应温度为350-370℃。
采用低苛刻度的加氢反应的主要目的在于将劣质原料油的脱硫率、脱氮率、脱残炭率和脱金属率控制在较低水平。具体地，可以控制所述的劣质原料油的脱硫率为50-95重量％，优选为65-85重量％，脱氮率为10-70重量％，优选为25-45重量％，脱残炭率为10-70重量％，优选为25-45重量％，脱金属率为50-95重量％，优选为65-80重量％。
根据本发明，所述的低苛刻度加氢反应在固定床反应器中进行。根据本发明，所述的低苛刻度加氢反应在加氢催化剂的存在下进行。按照加氢催化剂的功能，按照反应物的流动方向，用于所述的低苛刻度加氢反应的加氢催化剂可以依次包括加氢保护剂、加氢脱金属催化剂、加氢脱硫催化剂、和加氢脱氮脱残炭催化剂。优选地，以加氢催化剂的总重量为基准，加氢保护剂和加氢脱金属催化剂占20％-70％，例如30％-50％；加氢脱硫催化剂占20％-70％，例如40％-60％；加氢脱氮脱残炭催化剂占0％-60％，例如10％-40％，并且加氢保护剂、加氢脱金属催化剂、加氢脱硫催化剂、和加氢 脱氮脱残炭催化剂的总和是100重量％。所述加氢催化剂为本领域中常规使用的那些。在一种优选的实施方案中，以加氢催化剂的总重量为基准，加氢脱金属催化剂占30重量％或更高。
根据本发明，所述的劣质原料油是本领域所常规使用的。例如，所述的劣质原料油可以为石油烃和/或其他矿物油，其中石油烃可以选自常压瓦斯油、减压瓦斯油、常压渣油、减压渣油、加氢渣油、焦化瓦斯油、脱沥青油、和其任意的组合，其他矿物质油可以选自煤与天然气衍生的液体油、油砂油、致密油、页岩油、和其任意的组合。
从性质方面看，所述的劣质原料油可以满足：(1)在20℃时的密度为910-1000千克/立方米；和/或(2)残炭比重为4-15重量％；和/或(3)金属(Ni+V)含量为12-600ppm。优选地，所述的劣质原料油满足：(1)在20℃时的密度为980-1000千克/立方米；和/或(2)残炭比重为10-15重量％；和/或(3)金属(Ni+V)含量为60-600ppm。
根据本发明，所述的第一催化裂化反应是一种高选择性催化裂化工艺，该工艺不追求原料油单程转化率最高，而将转化率控制在适当水平，从而能够有效提高干气、焦炭选择性，同时产生较大量的催化裂化蜡油，用于进一步加氢处理。该工艺能够有效弥补低苛刻度渣油加氢对劣质原料加工深度不足问题，并能够优化产物分布。
所述的第一催化裂化反应可以包括如下步骤：(1)将预热的加氢渣油与第一再生催化裂化催化剂在第一催化裂化反应器的下部进行第一裂化反应，所得到的反应产物经分离得到第一裂化产物和第一半再生催化裂化催化剂；所述第一再生催化裂化催化剂的微反评价活性为35-60；(2)将步骤(1)中所得的第一裂化产物与所述的第一半再生催化裂化催化剂随后在第一催化裂化反应器的上部进行第一再催化转化反应，并且所得到的反应产物经过分离、分馏得到第一干气、第一液化气、第一汽油、第一柴油和第一蜡油。所述的第一催化裂化反应器的下部与上部由反应器的前1/3部分与前2/3部分(以反应物流动的方向)之间的某一位置分界；在一种优选的实施方案中，下部是指反应器长度的前1/2部分，上部是指反应器长度的后1/2部分。所述的第一裂化反应主要为大分子裂化反应，所述的第一再催化转化反应主要为选择性裂化、选择性氢转移和异构化等反应。所述的第一裂化反应可以在下述条件进行：反应温度为530-620℃，重时空速为30-180小时-1，剂油比(催化剂与 原料油的重量比)为4-12，水油比(水蒸气与原料油重量比)为0.03-0.3，反应压力为130千帕-450千帕。所述的第一再催化转化反应在下述条件进行：反应温度为460℃-520℃，重时空速为20-100小时-1，剂油比为3-15，水油比(水蒸气与原料油重量比)为0.03-0.3，反应压力为130千帕-450千帕。在所述的第一催化裂化反应中，所述的第一蜡油的氢含量为10.5-15重量％；以所述的加氢渣油为基准，所述的第一蜡油的产率为15-50重量％，优选为30-45重量％。
根据本发明，可以将二次加工蜡油与所述的第一蜡油一起进行步骤c中所述的蜡油加氢反应，以增加第二催化裂化的原料来源。所述的二次加工蜡油可以选自焦化蜡油、脱沥青油、其它装置所产的催化裂化蜡油、和其任意的组合。催化裂化蜡油不限于本发明的第一蜡油和第二蜡油，可以来自其它催化裂化装置。
根据本发明，所述的蜡油加氢反应可以在下述条件进行：反应压力可以为5.0-20.0兆帕，优选为6.0-15.0兆帕，反应温度可以为300-430℃，优选为320-390℃，液时空速可以为0.2-5.0小时-1，优选为0.3-2.5小时-1，氢油体积比可以为200-1800标准立方米/立方米，优选为400-1100标准立方米/立方米。
所述的蜡油加氢反应在固定床反应器中在加氢催化剂的存在下进行。按照加氢催化剂的功能，按照反应物的流动方向，用于所述的蜡油加氢反应的加氢催化剂可以依次包括加氢保护剂、加氢脱金属脱硫催化剂和加氢处理催化剂。优选地，以加氢催化剂的总重量为基准，加氢保护剂占0-30重量％，例如5-20重量％、加氢脱金属脱硫催化剂占5-35重量％，例如10-25重量％；和加氢处理催化剂占35％-95重量％，例如55-85重量％，并且加氢保护剂、加氢脱金属脱硫催化剂和加氢处理催化剂的总和是100重量％。所述加氢催化剂为本领域中常规使用的那些。
根据本发明，所述的第二催化裂化反应可以在本领域常规条件下进行，例如，反应温度为450℃-620℃，重时空速为1-100小时-1，剂油比为1-25，水油比为0.03-0.3。所述的第二催化裂化反应也可以采用高选择性催化裂化工艺，例如，所述的第二催化裂化反应可以包括如下步骤：(1)将预热的加氢蜡油与第二再生催化裂化催化剂在第二催化裂化反应器的下部进行第二裂化反应，所得到的反应产物经分离得到第二裂化产物和第二半再生催化裂化催化剂；(2)将步骤(1)中所得的第二裂化产物与所述的第二半再生催化裂化 催化剂随后在第二催化裂化反应器的上部进行第二再催化转化反应，并且所得到的反应产物经过分离、分馏得到第二干气、第二液化气、第二汽油、第二柴油和第二蜡油。所述的第二催化裂化反应器的下部与上部由反应器的前1/3部分与前2/3部分(以反应物流动的方向)之间的某一位置分界；在一种优选的实施方案中，下部是指反应器长度的前1/2部分，上部是指反应器长度的后1/2部分。
需要说明的是，本发明的方法所采用的加氢催化剂、催化裂化催化剂、加氢反应器和催化裂化反应器可以采用本领域常规使用的那些。加氢催化剂可以包含至少一种选自VIII族和/或至少一种选自VIB族的金属组分(作为活性成分)，以及氧化铝和/或二氧化硅(作为载体)。催化裂化催化剂可以包含沸石(作为活性组分)，优选中孔沸石和/或任选的大孔沸石；其中，中孔沸石可以选自ZSM系列和/或ZRP系列。所述的催化裂化反应器可以选自提升管、流化床、及其组合。所述的加氢反应器可以选自固定床、悬浮床、沸腾床、移动床、及其组合(优选固定床)。所述的催化裂化反应器和所述的加氢反应器的数量可以是1个、2个、3个或更多个，当反应器数量为2个时，反应器可以串联或并联；当反应器是3个或更多个时，反应器可以串联、并联或混联。
下面将结合附图提供本发明的一种具体实施方式。
来自管线9的劣质原料油与来自管线11的新氢和循环氢混合气体混合后，进入低苛刻度加氢反应器1，在低苛刻度加氢反应条件下，进行脱杂质、加氢脱金属、加氢脱硫、加氢脱氮与加氢脱残炭反应。得到的产物经管线13，进入低苛刻度加氢反应产物的分离单元2，富氢气相物流经管线14进入循环气处理系统3，经管线15送入循环氢压缩机4，然后经管线16与来自管线10的新氢混合。来自分离单元2的液相物流经管线17进入加氢分馏单元5，分别得到加氢气体(管线18)、加氢石脑油(管线19)、加氢柴油(管线20)和加氢渣油(管线21)。加氢渣油经管线21进入第一催化裂化反应器6，在高选择性催化裂化反应条件下进行反应，并经过分离、分馏后依次得到第一干气(管线25)、第一液化气(管线26)、第一汽油(管线27)、第一轻循环油(管线28)、第一蜡油(管线29)和油浆(管线30)，油浆经管线30由油浆泵送至第一催化裂化反应器6进一步反应。
第一蜡油经管线29与来自管线12的混合氢混合，并送入蜡油加氢反应 器7。离开蜡油加氢反应器7的物流在蜡油加氢产物的分离单元8中分离，得到的富氢气相物流经管线23与来自管线14的富氢气相物流混合并被送入循环气处理系统3，得到的液相物流(加氢蜡油)经管线24与来自管线21的加氢渣油混合并且被送入第一催化裂化反应器6。
下面将通过实施例来进一步说明本发明，但是本发明并不因此而受到任何限制。
本发明实施例所采用的仪器、装置和试剂，如无特别说明，均为本领域常规使用的那些。
实施例中所使用的分析方法如下所述：
金属(Ni+V)含量 石油化工分析方法RIPP 124-90
硫含量 石油化工分析方法RIPP 62-90
氮含量 石油化工分析方法RIPP63-90
残炭含量 石油化工分析方法RIPP148-90
上述方法记载在《石油化工分析方法(RIPP试验方法)》(杨翠定等编，科学出版社，1990)中。
按照下列公式分别计算硫、残炭、氮和金属的脱除率：
Figure PCTCN2016000577-appb-000001
实施例与对比例中所用的劣质原料油为减压渣油与常压渣油的混合渣油，其性质见表1。
表1：劣质原料油的性质
原料  
密度(20℃)，g/cm3 0.996
运动粘度(100℃)，mm2/s 162.5
碳，重量％ 82.68
氢，重量％ 10.45
硫，重量％ 4.36
氮，重量％ 0.20
残炭值，重量％ 13.5
金属(Ni+V)，ppm 94.9
饱和烃，重量％ 23.6
芳烃，重量％ 47.2
胶质，重量％ 23.6
沥青质(C7不溶物)，重量％ 5.6
实施例与对比例中所用的催化剂由中国石化催化剂分公司生产。
实施例1
实施例1提供了本发明的可调控的低苛刻度加氢反应，其中反应温度和液时空速随反应时间进行阶段性调节，而氢油体积比与氢分压分别维持在800标准立方米/立方米和15兆帕。劣质原料油的加氢产物中，加氢渣油的切割点为350℃。
加氢试验在连续高压固定床中试装置上进行，该装置包括3个串联的反应器，分别装有体积比为5∶45∶50的加氢保护剂(RG-10A)、加氢脱金属催化剂(RDM-2B)、加氢脱硫催化剂(RMS-1B)。试验进行时，该中试装置处于运转初期，运转时间不足50小时。
催化裂化试验在中型催化裂化装置上进行，采用提升管反应器，使用MLC-500催化剂。
蜡油加氢处理试验在固定床加氢反应器上进行，该固定床加氢反应器装填有体积比为4∶4∶15∶77的加氢保护催化剂A(RG-30A)、加氢保护催化剂B(RG-30B)、加氢脱金属脱硫催化剂(RMS-30)和加氢处理催化剂(RDA-1)。
对比例1
对比例1为常规渣油加氢试验，试验装置、试验原料与实施例1相同。不同之处在于劣质原料油的加氢反应的温度与液时空速分别恒定在390℃和0.25h-1。
实施例1和对比例1的反应条件和反应结果列于表2中。
实施例2
实施例1的5000-5500小时内获得的反应产物(见表3)作为后续研究的对象。加氢渣油作为第一催化裂化反应的原料油。加氢渣油经过第一催化裂化 反应、分离分馏后得到第一干气、第一液化气、第一汽油、第一柴油和第一蜡油。第一蜡油切割点为330℃，占进料量的33.23％。第一蜡油被送入蜡油加氢处理单元，所得的产物经过气、液分离，液相的加氢蜡油经过第二催化裂化反应后得到第二干气、第二液化气、第二汽油、第二柴油和第二蜡油。第二蜡油被送入蜡油加氢处理单元。
操作条件见表4，产品分布见表5。
对比例2
对比例2为现有的渣油加氢-重油催化裂化的组合。对比例1的5000-5500小时内的反应产物(见表3)作为后续研究的对象。加氢渣油经过反应、分离分馏后得到干气、液化气、汽油、柴油、油浆和焦炭。操作条件见表4，产品分布见表5。
对比例3
对比例3的工艺流程以及反应条件与实施例2基本相同，不同之处在于，在对比例3中，使用对比例1的5000-5500小时内的反应产物(见表3)作为后续研究的对象。操作条件见表4，产品分布见表5。
表2
Figure PCTCN2016000577-appb-000002
Figure PCTCN2016000577-appb-000003
表3：实施例1和对比例1所得的部分加氢渣油的性质和产物分布(实施例1所得的加氢渣油用于实施例2，对比例1所得的加氢渣油用于对比例2和对比例3)
项目 实施例1 对比例1
反应条件    
反应温度，℃ 380 390
液时空速，h-1 0.25 0.25
氢油体积比，v/v 800 800
氢分压，兆帕 15.0 15.0
氢耗 1.75 2.03
加氢渣油(切割点＞350℃)性质    
密度(20℃)，g/cm3 0.944 0.936
硫，重量％ 0.62 0.43
氮，重量％ 0.13 0.12
残炭，重量％ 8.1 6.2
金属(Ni+V)，ppm 20.0 17.8
产物分布，重量％    
气体 3.75 5.92
加氢石脑油 1.98 2.87
加氢柴油 7.46 11.95
加氢渣油 88.56 81.29
合计 101.75 102.03
操作周期/月 16 12
表4：实施例2、对比例2和对比例3的反应条件
Figure PCTCN2016000577-appb-000004
表5：实施例2、对比例2和对比例3的反应结果。
Figure PCTCN2016000577-appb-000005
*以加氢渣油原料为100％计算

Claims (18)

  1. 一种劣质原料油的处理方法，该方法包括：
    a、将劣质原料油进行低苛刻度加氢反应，所得到的反应产物经分离得到气体、加氢石脑油、加氢柴油和加氢渣油；其中，在低苛刻度加氢反应中，以所述的劣质原料油为基准，所述加氢渣油的产率为85-95重量％，优选87-93重量％，所述加氢渣油的性质基本保持恒定，其中所述的低苛刻度加氢反应在运转初期的反应温度为350-370℃，例如350-360℃；
    b、将步骤a中所得的加氢渣油进行第一催化裂化反应，所得到的反应产物经分离得到第一干气、第一液化气、第一汽油、第一柴油和第一蜡油；
    c、将步骤b中所得的第一蜡油进行蜡油加氢反应，所得到的反应产物经分离得到加氢蜡油；
    d、将步骤c中所得的加氢蜡油进行步骤b中所述的第一催化裂化反应或第二催化裂化反应。
  2. 根据前述权利要求中任一项的方法，该方法还包括步骤e：将在步骤d中所述的第二催化裂化反应中所得的第二蜡油进行步骤c中所述的蜡油加氢反应。
  3. 根据前述权利要求中任一项的方法，其中当加氢渣油的性质不期望地变化时，提高加氢反应的苛刻度，使加氢渣油的性质与运转初期的加氢渣油的性质基本保持恒定。
  4. 根据前述权利要求中任一项的方法，其中当加氢渣油密度增加超过0.001-0.005g/cm3，或残炭值增加超过0.1％～0.5％时，就增加加氢处理装置的反应苛刻度。
  5. 根据前述权利要求中任一项的方法，其中在步骤a中，所述的劣质原料油的脱硫率为50-95重量％，脱氮率为10-70重量％，脱残炭率为10-70重量％，脱金属率为50-95重量％。
  6. 根据前述权利要求中任一项的方法，其中低苛刻度加氢反应的反应条件包括：氢分压为8-20兆帕，反应温度为330-420℃，液时空速为0.1-1.5小时-1，总氢油体积比为200-1500标准立方米/立方米。
  7. 根据前述权利要求中任一项的方法，当加氢渣油的密度的增速超过0.005g/cm3/(1000小时)时，和/或当加氢渣油的残炭值的增速超过0.5重 量％/(1000小时)时，提高加氢反应的苛刻度，例如以2-10℃/(1000小时)来提高反应温度或以0.1-0.5h-1/(1000小时)来降低液时空速。
  8. 根据前述权利要求中任一项的方法，其中所述的劣质原料油为石油烃和/或其他矿物油，其中石油烃选自常压瓦斯油、减压瓦斯油、常压渣油、减压渣油、加氢渣油、焦化瓦斯油、脱沥青油、和其任意的组合，其他矿物质油选自煤与天然气衍生的液体油、油砂油、致密油、页岩油、和其任意的组合。
  9. 根据前述权利要求中任一项的方法，其中所述的劣质原料油满足：(1)在20℃时的密度为980-1000千克/立方米；和/或(2)残炭比重为10-15重量％；和/或(3)金属(Ni+V)含量为60-600ppm。
  10. 根据前述权利要求中任一项的方法，其中所述的第一催化裂化反应包括如下步骤：(1)将预热的加氢渣油与第一再生催化裂化催化剂在第一催化裂化反应器的下部进行第一裂化反应，所得到的反应产物经分离得到第一裂化产物和第一半再生催化裂化催化剂；所述第一再生催化裂化催化剂的微反评价活性为35-60；(2)将步骤(1)中所得的第一裂化产物与所述的第一半再生催化裂化催化剂随后在第一催化裂化反应器的上部进行第一再催化转化反应，并且所得到的反应产物经过分离、分馏得到第一干气、第一液化气、第一汽油、第一柴油和第一蜡油。
  11. 根据权利要求10的方法，其中所述的第一裂化反应在下述条件进行：反应温度为530-620℃，重时空速为30-180小时-1，剂油比为4-12，水油比为0.03-0.3，反应压力为130千帕-450千帕；所述的第一再催化转化反应的条件为：反应温度为460℃-520℃，重时空速为20-100小时-1，剂油比为3-15，水油比为0.03-0.3，反应压力为130千帕-450千帕。
  12. 根据前述权利要求中任一项的方法，其中所述的第一蜡油的氢含量为10.5-15重量％；以所述的加氢渣油为基准，所述的第一蜡油的产率为15-50重量％。
  13. 根据前述权利要求中任一项的方法，其中，将二次加工蜡油与所述的第一蜡油一起进行所述的蜡油加氢反应；所述的二次加工蜡油选自焦化蜡油、脱沥青油、其它装置所产的催化裂化蜡油、和其任意的组合。
  14. 根据前述权利要求中任一项的方法，其中所述的蜡油加氢反应在固定床反应器中在加氢催化剂的存在下进行。
  15. 根据前述权利要求中任一项的方法，其中所述的蜡油加氢反应在下述条件进行：反应压力为5.0-20.0兆帕，反应温度为300-430℃，液时空速为0.2-5.0小时-1，氢油体积比为200-1800标准立方米/立方米。
  16. 根据前述权利要求中任一项的方法，其中，所述的第二催化裂化反应在下述条件进行：反应温度为450℃-620℃，重时空速为1-100小时-1，剂油比为1-25，水油比为0.03-0.3。
  17. 根据前述权利要求中任一项的方法，其中，所述的第二催化裂化反应包括如下步骤：
    (1)将预热的加氢蜡油与第二再生催化裂化催化剂在第二催化裂化反应器的下部进行第二裂化反应，所得到的反应产物经分离得到第二裂化产物和第二半再生催化裂化催化剂；
    (2)将步骤(1)中所得的第二裂化产物与所述的第二半再生催化裂化催化剂随后在第二催化裂化反应器的上部进行第二再催化转化反应，并且所得到的反应产物经过分离、分馏得到第二干气、第二液化气、第二汽油、第二柴油和第二蜡油。
  18. 根据前述权利要求中任一项的方法，其中所述的低苛刻度加氢反应的运转初期是指0-1000小时的阶段。
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