WO2018099419A1 - 一种十六氢芘的制备方法 - Google Patents

一种十六氢芘的制备方法 Download PDF

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WO2018099419A1
WO2018099419A1 PCT/CN2017/113834 CN2017113834W WO2018099419A1 WO 2018099419 A1 WO2018099419 A1 WO 2018099419A1 CN 2017113834 W CN2017113834 W CN 2017113834W WO 2018099419 A1 WO2018099419 A1 WO 2018099419A1
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reaction
hydrogenation
catalyst
molecular sieve
hexadecahydroquinone
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PCT/CN2017/113834
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English (en)
French (fr)
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孙国权
方向晨
樊宏飞
姚春雷
全辉
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中国石油化工股份有限公司
中国石油化工股份有限公司抚顺石油化工研究院
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Priority to RU2019118573A priority Critical patent/RU2717334C1/ru
Priority to PL17877027.7T priority patent/PL3533856T3/pl
Priority to US16/464,405 priority patent/US11111191B2/en
Priority to EP17877027.7A priority patent/EP3533856B1/en
Priority to JP2019528074A priority patent/JP6772382B2/ja
Priority to KR1020197018875A priority patent/KR102294660B1/ko
Publication of WO2018099419A1 publication Critical patent/WO2018099419A1/zh
Priority to ZA2019/03463A priority patent/ZA201903463B/en

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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C5/00—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms
    • C07C5/02—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by hydrogenation
    • C07C5/10—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by hydrogenation of aromatic six-membered rings
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
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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
    • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
    • C10G45/44—Hydrogenation of the aromatic hydrocarbons
    • C10G45/46—Hydrogenation of the aromatic hydrocarbons characterised by the catalyst used
    • C10G45/52—Hydrogenation of the aromatic hydrocarbons characterised by the catalyst used containing platinum group metals or compounds thereof
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • B01J29/08—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the faujasite type, e.g. type X or Y
    • B01J29/10—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the faujasite type, e.g. type X or Y containing iron group metals, noble metals or copper
    • B01J29/12—Noble metals
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    • B01J29/00—Catalysts comprising molecular sieves
    • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
    • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • B01J29/08—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the faujasite type, e.g. type X or Y
    • B01J29/10—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the faujasite type, e.g. type X or Y containing iron group metals, noble metals or copper
    • B01J29/12—Noble metals
    • B01J29/126—Y-type faujasite
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
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    • B01J35/00—Catalysts, in general, characterised by their form or physical properties
    • B01J35/30—Catalysts, in general, characterised by their form or physical properties characterised by their physical properties
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J35/00—Catalysts, in general, characterised by their form or physical properties
    • B01J35/40—Catalysts, in general, characterised by their form or physical properties characterised by dimensions, e.g. grain size
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C13/00—Cyclic hydrocarbons containing rings other than, or in addition to, six-membered aromatic rings
    • C07C13/28—Polycyclic hydrocarbons or acyclic hydrocarbon derivatives thereof
    • C07C13/32—Polycyclic hydrocarbons or acyclic hydrocarbon derivatives thereof with condensed rings
    • C07C13/62—Polycyclic hydrocarbons or acyclic hydrocarbon derivatives thereof with condensed rings with more than three condensed rings
    • C07C13/66—Polycyclic hydrocarbons or acyclic hydrocarbon derivatives thereof with condensed rings with more than three condensed rings the condensed ring system contains only four rings
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    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
    • C10G45/44—Hydrogenation of the aromatic hydrocarbons
    • 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
    • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
    • C10G45/44—Hydrogenation of the aromatic hydrocarbons
    • C10G45/46—Hydrogenation of the aromatic hydrocarbons characterised by the catalyst used
    • C10G45/54—Hydrogenation of the aromatic hydrocarbons characterised by the catalyst used containing crystalline alumino-silicates, e.g. molecular sieves
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    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G65/00—Treatment of hydrocarbon oils by two or more hydrotreatment processes 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
    • 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
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2229/00—Aspects of molecular sieve catalysts not covered by B01J29/00
    • B01J2229/10—After treatment, characterised by the effect to be obtained
    • B01J2229/18—After treatment, characterised by the effect to be obtained to introduce other elements into or onto the molecular sieve itself
    • B01J2229/186—After treatment, characterised by the effect to be obtained to introduce other elements into or onto the molecular sieve itself not in framework positions
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2229/00—Aspects of molecular sieve catalysts not covered by B01J29/00
    • B01J2229/10—After treatment, characterised by the effect to be obtained
    • B01J2229/20—After treatment, characterised by the effect to be obtained to introduce other elements in the catalyst composition comprising the molecular sieve, but not specially in or on the molecular sieve itself
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2229/00—Aspects of molecular sieve catalysts not covered by B01J29/00
    • B01J2229/30—After treatment, characterised by the means used
    • B01J2229/32—Reaction with silicon compounds, e.g. TEOS, siliconfluoride
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2229/00—Aspects of molecular sieve catalysts not covered by B01J29/00
    • B01J2229/30—After treatment, characterised by the means used
    • B01J2229/36—Steaming
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2229/00—Aspects of molecular sieve catalysts not covered by B01J29/00
    • B01J2229/30—After treatment, characterised by the means used
    • B01J2229/42—Addition of matrix or binder particles
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2523/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00
    • C07C2523/38—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of noble metals
    • C07C2523/40—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of noble metals of the platinum group metals
    • C07C2523/44—Palladium
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2529/00—Catalysts comprising molecular sieves
    • C07C2529/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites, pillared clays
    • C07C2529/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • C07C2529/08—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the faujasite type, e.g. type X or Y
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2529/00—Catalysts comprising molecular sieves
    • C07C2529/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites, pillared clays
    • C07C2529/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • C07C2529/08—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the faujasite type, e.g. type X or Y
    • C07C2529/10—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the faujasite type, e.g. type X or Y containing iron group metals, noble metals or copper
    • C07C2529/12—Noble metals

Definitions

  • the invention relates to a method for preparing hexadecahydroquinone.
  • Niobium is an important component in coal tar. It is enriched in fractions of 300-360 ° C during the initial distillation of coal tar, and can be obtained by conventional methods such as rectification and crystallization.
  • Coal tar raw materials basically do not contain heptahydroquinone, and the extraction of hexadecahydroquinone directly from coal tar has not been reported. The operation cost is large, and the extracted hexadecahydroquinone has low purity and low yield, and is not economically feasible.
  • the preparation of 1,2,3,6,7,8-hexahydroindole is generally mainly carried out by selective catalytic hydrogenation of hydrazine, and it is difficult to select a single product due to the simultaneous reaction of the series reaction and the parallel reaction. It is more difficult to improve the method of obtaining high purity hexadecahydroquinone.
  • CN1351130A discloses a method for hydrogenating coal tar to produce diesel oil, which is mainly characterized in that the coal tar is subjected to fractional distillation, and the fractions below the diesel oil are hydrotreated, and the diesel fuel meeting the fuel index can be directly produced or the blending component can be produced as a diesel product.
  • it only hydrotreats the lighter fractions of coal tar, and cannot extract high value-added, high-purity hexadecahydroquinone products without fully utilizing coal tar.
  • CN1676583A discloses a medium-high temperature coal tar hydrocracking process.
  • the process is as follows: medium and high temperature coal tar is heated to 250-300 ° C in a heating furnace, mixed with hydrogen into a hydrotreating reactor, refined to form oil through a distillation device, and fractionated gasoline, diesel, lubricating oil and hydrogenated tail oil, plus After the hydrogen tail oil is heated by the cracking furnace, it is mixed with hydrogen and then enters the cracking reactor to further produce the gasoline and diesel fraction.
  • the hexadecahydroquinone product cannot be extracted from the fraction, but the blended fuel oil is produced.
  • the direct entry of the process coal tar into the high temperature heating furnace will cause the furnace tube to coke, which affects the normal operation cycle of the device.
  • the object of the present invention is to provide a process for the preparation of hexadecahydroquinone, by which a high purity hexadecahydroquinone product can be prepared.
  • the present invention provides a process for preparing hexadecahydroquinone, which comprises: subjecting a hydrocarbon oil feedstock containing a lanthanoid compound to a hydrogenation reaction in the presence of a hydrogenation catalyst, wherein the lanthanoid compound is selected from the group consisting of ruthenium And at least one of an unsaturated hydrogenation product thereof, the hydrogenation catalyst comprising a carrier and an active metal component supported on a carrier, the active metal component being Pt and/or Pd, the carrier containing a small crystal Granular Y-type molecular sieve, alumina and amorphous silica-alumina, the small-grain Y-type molecular sieve has an average crystal grain diameter of 200-700 nm, a SiO 2 /Al 2 O 3 molar ratio of 40-120, and a relative crystallinity ⁇ 95 %, the specific surface area is 900-1200 m 2 / g, and the pore volume of the secondary pore of
  • a high purity hexadecahydroquinone product can be obtained.
  • FIG. 1 is a schematic flow chart showing an embodiment of a method for producing hexadecahydroquinone according to the present invention.
  • FIG. 2 is a schematic flow chart of another embodiment of a method for preparing hexadecahydroquinone according to the present invention.
  • the preparation method of hexadecahydroquinone according to the present invention comprises subjecting a hydrocarbon oil raw material containing a lanthanoid compound to a hydrogenation reaction in the presence of a hydrogenation catalyst.
  • the hydrogenation catalyst contains a carrier and an active metal component supported on a carrier.
  • the active metal component is Pt and/or Pd.
  • the support contains small grain Y-type molecular sieves, alumina, and amorphous silica-alumina.
  • the small-grain Y-type molecular sieve has an average crystal grain diameter of 200 to 700 nm, and specifically, for example, may be 200 nm, 250 nm, 300 nm, 350 nm, 400 nm, 450 nm, 500 nm, 550 nm, 600 nm, 650 nm, 700 nm, and among these values Any of the ranges of any two of them.
  • the small-grain Y-type molecular sieve has an average crystal grain diameter of 300 to 500 nm.
  • the average grain diameter of the small-grain Y-type molecular sieve is measured by SEM (Scanning Electron Microscopy).
  • the small-grain Y-type molecular sieve has a SiO 2 /Al 2 O 3 molar ratio of 40-120, specifically, for example, 40, 50, 60, 70, 80, 90, 100, 110, 120 and these point values Any of the ranges formed by any two of them.
  • the small crystal Y-type molecular sieve has a relative crystallinity of ⁇ 95%, preferably 95-120%, more preferably 98-115%.
  • the relative crystallinity of the small-grain Y-type molecular sieve is detected by an X-ray diffraction method.
  • the specific surface area is small crystal Y zeolite is 900-1200m 2 / g, in particular, for example, can be 900m 2 / g, 920m 2 / g, 950m 2 / g, 980m 2 / g, 1000m 2 / g, 2 / g 1050m 2 / g 1080m 2 / g 2 / g, 1120m 2 / g, 1150m 2 / g 1180m 2 / g / g and any two 1020m,, 1100m,, 1200m 2 points constituted values Any value in the range.
  • the specific surface area of the small-grain Y-type molecular sieve is detected according to a low-temperature liquid nitrogen physical adsorption method.
  • the small-grain Y-type molecular sieve has more secondary pores. Specifically, in the small-grain Y-type molecular sieve, the pore volume of the secondary pore of 1.7-10 nm accounts for more than 50% of the total pore volume, preferably It is 50-80%, further preferably 60-80%. In the present invention, the pore volume of the secondary pore of the small-grain Y-type molecular sieve is detected by a low-temperature liquid nitrogen physical adsorption method.
  • the unit cell constant of the small-grain Y-type molecular sieve may be 2.425-2.435 nm, for example, 2.425 nm, 2.426 nm, 2.427 nm, 2.428 nm, 2.429 nm, 2.43 nm, 2.431 nm, 2.432 nm, 2.433 nm, 2.434 nm, 2.435 nm and any value in the range formed by any two of these point values.
  • the unit cell constant of the small-grain Y-type molecular sieve is detected by an X-ray diffraction method.
  • the small-grain Y-type molecular sieve may have a pore volume of 0.5-0.8 mL/g, for example, 0.5 mL/g, 0.55 mL/g, 0.6 mL/g, 0.65 mL/g, 0.7 mL/g, and 0.75 mL/g. Any value in the range of 0.8 mL/g and any two of these point values.
  • the pore volume of the small-grain Y-type molecular sieve is detected by a low-temperature liquid nitrogen physical adsorption method.
  • the hydrogenation catalyst has properties as follows: the specific surface area may be from 350 to 550 m 2 /g, preferably from 380 to 500 m 2 /g; and the pore volume may be from 0.5 to 1 mL/g, preferably from 0.5 to 0.9 mL. /g.
  • the content of the active metal component may be 0.1 to 2% by weight, preferably 0.2 to 1.5% by weight based on the total weight of the hydrogenation catalyst; the content of the carrier may be It is 98-99.9 wt%, preferably 98.5-99.8 wt%.
  • the content of the small-grain Y-type molecular sieve may be 5-40% by weight, preferably 10-25% by weight, based on the total weight of the carrier; the content of the alumina may be 10-40% by weight, preferably 15-30% by weight; the amorphous silicon aluminum may be present in an amount of from 20 to 65% by weight, preferably from 30 to 60% by weight.
  • the hydrogenation catalyst may be selected from a suitable commercial catalyst, or may be according to the art.
  • the preparation is carried out by a conventional method, for example, according to the method reported in CN104588073A.
  • the preparation method of the hydrogenation catalyst may include: mechanically mixing, molding, and then drying and calcining a small-grain Y-type molecular sieve, an amorphous silicon-aluminum, and an adhesive made of alumina to prepare a catalyst carrier.
  • the Pt and/or Pd are supported on the support by impregnation, dried and calcined to obtain a hydrogenation catalyst.
  • the preparation method of the small-grain Y-type molecular sieve may include the following steps:
  • the molecular sieve obtained in the step (3) is treated with a mixed solution containing NH 4 + and H + , and then washed and dried to obtain a small-grain Y-type molecular sieve.
  • the properties of the small-grained NaY molecular sieve are as follows: the SiO 2 /Al 2 O 3 molar ratio is greater than 6 and not higher than 9, preferably 6.5-9, further preferably 7-8; the average grain diameter is 200-700 nm, It is preferably 300-500 nm; the specific surface area is 800-1000 m 2 /g, preferably 850-950 m 2 /g; the pore volume is 0.3-0.45 mL/g, the relative crystallinity is 90-130%, and the unit cell constant is 2.46-2.47.
  • the relative crystallinity after calcination in air at 650 ° C for 3 hours is 90% or more, preferably 90-110%, more preferably 90-105%.
  • the hydrogenation reaction process comprises two reaction stages which are carried out in sequence, correspondingly, the catalyst used in the first reaction stage is hydrogenation catalyst A, and the second reaction stage is used.
  • the catalyst is hydrogenation catalyst B.
  • the content percentage x 1 of the active metal component in the hydrogenation catalyst A is lower than the content percentage x 2 of the active metal component in the hydrogenation catalyst B, preferably x 1 is 0.1 to 1.5 percentage points lower than x 2 More preferably, x 1 is 0.3-1.5 percentage points lower than x 2 .
  • the content percentage y 1 of the small-grain Y-type molecular sieve in the hydrogenation catalyst A is higher than the content percentage y 2 of the small-grain Y-type molecular sieve in the hydrogenation catalyst B, preferably y 1 is higher than y 2 -35 percentage points, more preferably y 1 is 10-35 percentage points higher than y 2 .
  • y 1 is higher than y 2 -35 percentage points, more preferably y 1 is 10-35 percentage points higher than y 2 .
  • Higher purity hexadecahydroquinone can be obtained in accordance with the preferred embodiment described above, and the yield is higher.
  • the process conditions of the hydrogenation reaction can be a conventional choice in the art.
  • the hydrogenation reaction process conditions include: hydrogen partial pressure of 4-20 MPa, liquid hour volumetric space velocity of 0.05-6 h -1 , hydrogen oil volume ratio of 50-3000, and average reaction temperature of 150- 380 ° C.
  • the average reaction temperature of the second reaction stage is lower than the average reaction temperature of the first reaction stage by 10 - 150 ° C, preferably 30-120 ° C. More preferably, the average reaction temperature of the first reaction stage is 180-380 ° C, further preferably 220-350 ° C; the average reaction temperature of the second reaction stage is 150-350 ° C, preferably 180-330 ° C.
  • the first reaction stage and the second reaction stage may be carried out in the same reactor or in two or more reactors connected in series.
  • the lanthanoid compound is at least one selected from the group consisting of ruthenium and its unsaturated hydrogenation product.
  • the unsaturated hydrogenation product of hydrazine may be, for example, indoline, tetrahydroanthracene, hexahydroanthracene, octahydroquinone or the like.
  • the content of the lanthanoid compound may be 0.5% by weight or more, and specifically, for example, may be 0.5 to 10% by weight, such as 0.5% by weight, 0.8% by weight, or 1.0% by weight. %, 1.2% by weight, 1.5% by weight, 2% by weight, 3% by weight, 4% by weight, 5% by weight, 6% by weight, 7% by weight, 8% by weight, 9% by weight, and 10% by weight.
  • the hydrocarbon oil raw material containing a lanthanoid compound may be a hydrocarbon oil raw material conventional in the art as long as it contains a predetermined amount of a lanthanoid compound.
  • the hydrocarbon oil feedstock containing a lanthanide compound is a heavy distillate having an initial boiling point of from 130 to 220 ° C (preferably from 160 to 200 ° C).
  • the hydrocarbon oil feedstock containing a lanthanide compound is a heavy distillate having an initial boiling point of from 130 to 220 ° C and a final boiling point of from 300 to 400 ° C.
  • the hydrocarbon oil raw material containing a lanthanoid compound is a diesel fraction, and has an initial boiling point of 160 to 200 ° C and a final boiling point of 300 to 350 ° C.
  • the hydrocarbon oil raw material containing a lanthanoid compound is prepared according to the method comprising the following steps:
  • reaction effluent obtained by the hydrocracking reaction is subjected to gas-liquid separation, and then the separated liquid phase is fractionated, and the fractionated heavy fraction is used as the hydrocarbon oil raw material containing the lanthanoid compound.
  • the coal tar may be at least one of low temperature coal tar, medium temperature coal tar or high temperature coal tar, or may be after the coal tar extracts at least one of naphthalene, anthracene, phenanthrene, carbazole and fluoranthene. Distillate.
  • the coal tar generally has an aromatic content of 20 to 100% by weight, and a density of 20 ° C is generally 1.023 to 1.235 g/cm 3 .
  • the distillation range of the coal tar is in any range of 200 to 700 ° C, and the temperature difference between the initial boiling point and the final boiling point is generally between 100 and 400 ° C.
  • the coal tar raw material is a residual fraction of at least one of cerium, phenanthrene, oxazole and fluoranthene extracted from high temperature coal tar or high temperature coal tar.
  • the pretreatment described in the step (1) generally includes mechanical de-ingufacturing, dehydration, electric desalting, and the like, and may optionally include extraction and removal of ruthenium, phenanthrene, and the like.
  • the catalyst used in the hydrofinishing reaction described in the step (2) may be a hydrofinishing catalyst conventional in the art, such as a diesel hydrotreating catalyst or a hydrocracking pretreatment catalyst.
  • the hydrotreating catalyst generally has a VIB group and/or a Group VIII metal as an active component, an alumina or a silica-containing alumina as a carrier, a Group VIB metal is generally Mo and/or W, and a Group VIII metal is generally Co. And / or Ni.
  • the Group VIB metal content is 10 to 50% by weight based on the oxide, and the Group VIII metal content is 3 to 15% by weight based on the oxide, based on the weight of the catalyst; the properties are as follows: a specific surface area of 100 to 350 m 2 / g, pore volume is 0.15 ⁇ 0.6mL / g.
  • the alternative commercial catalysts are 3936, 3996, FF-16, FF-26, FF-36, FF-46, FF-56, FF-66 and other hydrogenated products developed by China Petroleum and Chemical Corporation Fushun Petrochemical Research Institute.
  • the purified catalyst may also be HC-K, HC-P catalyst of UOP, TK-555, TK-565 catalyst of Topsoe, and KF-847, KF-848 of AKZO.
  • the process conditions of the hydrotreating reaction described in the step (2) are generally: hydrogen partial pressure of 3 to 19 MPa, average reaction temperature of 260 to 440 ° C, liquid hour volumetric space velocity of 0.1 to 4 h -1 , hydrogen oil volume ratio It is 300:1 to 3000:1.
  • the hydrotreating described in the step (2) may be selected from a conventional reactor form such as a fixed bed or a bubbling bed.
  • the fixed bed reactor can be in the form of an upflow (parallel) reactor, a downflow (parallel) reactor or a gas liquid countercurrent reactor.
  • the catalyst used in the hydrocracking reaction process described in the step (3) may be a conventional hydrocracking catalyst in the art, such as a light oil type hydrocracking catalyst, a flexible hydrocracking catalyst, and a (high) medium oil.
  • Type hydrocracking catalyst generally has a Group VIB and/or Group VIII metal as the active component, Group VIB metals are generally Mo and/or W, and Group VIII metals are typically Co and/or Ni.
  • the carrier of the catalyst is one or more of alumina, silica-containing alumina, and molecular sieves.
  • the Group VIB metal content is 10 to 35 wt% based on the oxide, the Group VIII metal content is 3 to 15 wt% based on the oxide, and the molecular sieve content is 5 to 40 wt%, and the alumina content is based on the weight of the catalyst. It is 10 to 80% by weight; its specific surface area is 100 to 650 m 2 /g, and the pore volume is 0.15 to 0.50 mL/g.
  • the catalysts for the selection of products are FC-26, FC-28, FC-14, ZHC-01, ZHC-02, ZHC-04 and other single-stage hydrocracking developed by China Petroleum and Chemical Corporation Fushun Petrochemical Research Institute.
  • hydrocracking catalysts such as UHC's DHC39, DHC-8, and CHERON's ICR126.
  • ZHC-02 and ICR126 are hydrocracking catalysts with amorphous silicon aluminum and Y type molecular sieves as cracking components, which are more suitable for the hydrocracking reaction process of the present invention.
  • a (high) medium oil type hydrocracking catalyst is preferably used.
  • a medium oil type hydrocracking catalyst such as FC-26 catalyst
  • FC-26 catalyst FC-26 catalyst
  • the catalyst has good chain-breaking function for alkane and side chain aromatics under hydrogenation conditions, and can be used in the cyclic hydrocarbons (including cycloalkanes, side chain cycloalkanes, aromatic hydrocarbons, aromatic hydrocarbons with side chains) in the raw materials.
  • the side chain alkane is broken.
  • the catalyst has suitable fused aromatic hydrocarbons (without side chain) saturation function, and almost no open loop.
  • the resulting oil obtained by hydrocracking is subjected to fractional distillation, and the component containing the precursor of the intended product can be as concentrated as possible in a suitable narrow fraction. Therefore, the use of a medium oil type hydrocracking catalyst can maintain the maximum amount of cyclic hydrocarbons in the product, and contributes to an increase in the yield of the final product of interest.
  • the reactor used in the hydrocracking system is a conventional fixed bed hydrogenation reactor, more preferably a downflow fixed bed reactor.
  • the process conditions of hydrocracking in step (3) are generally: hydrogen partial pressure of 3 to 19 MPa, average reaction temperature of 260 to 440 ° C, volumetric space velocity of liquid is 0.3 to 4 h -1 , and hydrogen oil volume ratio of 300:1 to 5000:1.
  • step (3) “optional” means that the separation (e.g., gas-liquid separation) process may or may not be performed (e.g., gas-liquid separation).
  • the fractionation operation described in the step (4) can select conventional techniques in the art.
  • the initial distillation point of the fractionated heavy fraction may be 130-220 °C. It is preferably 160-200 °C.
  • the heavy fraction obtained by fractional distillation in step (4) is divided into a diesel fraction, and more preferably, the initial fraction of the diesel fraction is 130 to 220 ° C, more preferably 160 to 200 ° C; and the final boiling point is 280 to 400 ° C. More preferably, it is 300-350 degreeC.
  • the method may further include removing the naphtha fraction from the heavy fraction obtained in the step (4), and using the obtained remaining liquid fraction as the hydrocarbon oil raw material containing the lanthanoid compound.
  • the method may further comprise separating and fractionating the reaction effluent obtained by the hydrogenation reaction to obtain a hexadecahydroquinone-rich component and a heavy component, and recycling at least a portion of the heavy component back to the above step (
  • the hydrocracking reaction is carried out in 3).
  • the fractionation process herein can employ conventional fractionation techniques in the art.
  • the product obtained by fractional distillation, in addition to the hexadecahydroquinone component (intermediate component) and the heavy component also includes a liquid light component.
  • the liquid light component and the intermediate component (rich in the hexadecahydroquinone component) have a cutting temperature of 130 to 280 ° C, preferably 200 to 260 ° C.
  • the intermediate component and the heavy component have a cutting temperature of 300 to 360 ° C, preferably 320 to 340 ° C.
  • the product is a high-purity hexadecahydroquinone product, and the purity thereof can be more than 95 wt%.
  • the liquid heavy component obtained above the cutting temperature contains a heptacyclic or higher hydrocarbon such as dibenzopyrene or indenofluorene, and can be converted into hydrazine by cyclic hydrocracking, thereby improving the yield of the target product.
  • the method further comprises separating and fractionating the reaction effluent obtained by the hydrogenation reaction to obtain a hexadecahydroquinone-rich component and a heavy component,
  • the hexadecahydroquinone-rich component is cooled and cooled, and then filtered and extracted to obtain a solid hexadecahydroquinone.
  • the fractionation process herein can employ conventional fractionation techniques in the art.
  • the initial boiling point of the liquid fraction rich in hexadecahydroquinone obtained by fractional distillation is generally 220-300 ° C, preferably 260-280 ° C; the final boiling point is generally >300-360 ° C (greater than 300 ° C and less than or equal to 360 ° C), It is preferably 320 to 340 °C.
  • the liquid fraction is cooled and cooled, and the resulting hexadecahydroquinone is crystallized and precipitated in the liquid, and then filtered and optionally subjected to centrifugation to obtain a high-purity hexadecahydroquinone product.
  • the preparation method of the hexadecahydroquinone comprises:
  • the hydrotreating reaction effluent obtained in the step (2) optionally after separation, enters the hydrocracking reaction zone together with the hydrogen, and is contacted with the hydrocracking catalyst for reaction;
  • hydrocracking effluent is subjected to gas-liquid separation, and the liquid is subjected to fractional distillation to obtain a heavy fraction, and the initial boiling point of the heavy fraction is 130 to 220 ° C;
  • Step (5) After the reaction effluent is separated and fractionated, a heptahydroquinone-rich component and a heavy component are obtained, and the hexadecahydroquinone-rich component is cooled and cooled, and then filtered and vacuum-extracted. The resulting solid is the hexadecahydroquinone product.
  • the hydrogenation reaction process of the above step (5) comprises two reaction stages which are sequentially carried out, correspondingly, the catalyst used in the first reaction stage is hydrogenation catalyst A, and the catalyst used in the second reaction stage is Hydrogenation catalyst B.
  • the content percentage x 1 of the active metal component in the hydrogenation catalyst A is lower than the content percentage x 2 of the active metal component in the hydrogenation catalyst B, preferably x 1 is 0.1 to 1.5 percentage points lower than x 2 More preferably, x 1 is 0.3-1.5 percentage points lower than x 2 .
  • the content percentage y 1 of the small-grain Y-type molecular sieve in the hydrogenation catalyst A is higher than the content percentage y 2 of the small-grain Y-type molecular sieve in the hydrogenation catalyst B, preferably y 1 is higher than y 2 -35 percentage points, more preferably y 1 is 10-35 percentage points higher than y 2 .
  • the hydrogenation catalyst used has different properties depending on the content of the active metal component and the content of the small-grain Y-type molecular sieve, respectively.
  • the hydrogenation catalyst A has a relatively low active metal component content and a higher Y-type molecular sieve content, and thus the cracking performance of the catalyst is high.
  • the paraffin and the side-chain polycyclic aromatic hydrocarbon still contained in the diesel fraction obtained by hydrocracking coal tar are further contacted with the hydrogenation catalyst A, and the side chain on the polycyclic aromatic hydrocarbon is almost completely stripped from the aromatic ring after the reaction.
  • the hydrogenation catalyst B has a higher hydrogenation performance and a weaker cleavage activity because of its relatively high active metal component content and relatively low content of small-grain Y-type molecular sieve.
  • the hydrogenation product of the first reaction stage is further contacted with the hydrogenation catalyst B During the reaction, the non-perhydrohydroquinone (such as hexahydroanthracene) formed by partial hydrogenation is restricted at the lower reaction temperature due to the cracking activity of the catalyst, and the saturation function is strong, further hydrogenation, and all carbon and carbon Both of the double bonds are saturated to give a hexadecahydroquinone (perhydrohydroquinone) product, thereby increasing the yield of the production of hexadecahydroquinone by the process of the present invention.
  • the non-perhydrohydroquinone such as hexahydroanthracene
  • the hydrogenation catalyst is graded and combined in the above preferred manner in the hydrogenation reaction process, and the hydrogenation saturation of the condensed aromatic hydrocarbons, especially the crude ruthenium in the diesel fraction, is well realized, so that the hydrogenation method can be directly used. Production of high purity hexadecahydroquinone products.
  • a coal tar raw material is used as a starting material, and a suitable process flow is selected, and a high-purity hexadecahydroquinone can be prepared by a hydrogenation process, and a solvent oil excellent in performance can also be obtained. product.
  • the method of the present invention greatly broadens the potential of coal tar to produce high value-added products. It not only provides a low value-added coal tar with a processing method to improve its economy, but also develops a new raw material and a new process route for the hexadecahydroquinone product.
  • the invention firstly passes the hydrorefining, hydrocracking and fractionation processes, and has high aromatics content in the coal tar hydrocracking diesel oil, and generates more light and heavy components which are easily soluble in the hydrocracking process.
  • the hydrocracking oil is distilled to fractionate the narrow fraction rich in hexadecahydroquinone, thereby enriching the tricyclic or higher fused aromatic hydrocarbon component into the diesel fraction, and realizing the separation of the components of the compatible product.
  • the hydrogenation catalyst used in the present invention uses a small-grain Y-type molecular sieve as an acidic component, and the Y-type molecular sieve has the characteristics of high silicon-to-aluminum ratio, high crystallinity, multiple secondary pores, and large specific surface area.
  • Cooperating with amorphous silicon aluminum and hydrogenation active metal components Pt and Pd not only promotes the hydrogenation saturation activity of aromatic hydrocarbons, but also facilitates selective ring opening and chain scission of aromatic hydrocarbons, and is beneficial to the diffusion of reaction products.
  • the capacity of the carbon is greatly enhanced, thereby improving the activity, selectivity and stability of the catalyst.
  • the catalyst is particularly suitable as a naphthenic starting material, especially in a hydrodearomatization reaction of a cycloalkyl starting material having a high viscosity and a high content of fused aromatic hydrocarbons.
  • the hydrogenation catalyst A used in the first reaction stage has a relatively high Y-type molecular sieve content and a relatively low metal content, and thus exhibits a partial cleavage activity.
  • the partial cracking performance of hydrogenation catalyst A can effectively cleave the fused aromatic hydrocarbons with side paraffins, further The chain is stripped from the aromatic ring.
  • the hydrogenation catalyst B used in the second reaction stage has a relatively high metal content and a low Y-type molecular sieve content, and the hydrogenation performance is strong, and the suitable cracking activity is also for the condensed aromatic hydrocarbon. Hydrogenation has an important catalytic effect. Therefore, the non-perhydrohydroquinone (such as hexahydroindole) which has been partially hydrogenated in the first reaction stage can complete the hydrogenation saturation of all the aromatic rings in the second reaction stage at a lower reaction temperature, thereby A hexadecahydroquinone (perhydrohydrazine) product is obtained.
  • the invention provides a process for the stepwise saturation of various heterocyclic aromatic hydrocarbons in the diesel fraction obtained by hydrocracking for the production of hexadecahydroquinone product and low aromatic solvent oil, which can avoid the high temperature in the single-stage process to the greatest extent.
  • the aromatic hydrocarbon condensation deposits and cracking reactions in the coal tar hydrocracking diesel oil fraction seriously affect the life of the catalyst.
  • a process flow of the present invention is: after the pretreatment (the pretreatment unit is omitted in the figure), the coal tar passes through the pipeline 1 and is mixed with the hydrogen passing through the pipeline 2 to enter the hydrotreating reactor 3 .
  • the gas-liquid separator 7 generally comprises a high pressure separator and a low pressure separator), and the resulting hydrogen rich gas is mixed with fresh hydrogen entering the line 9 after passing through the line 10 and optionally subjected to dehydrogenation treatment to obtain recycled hydrogen.
  • the liquid obtained by the gas-liquid separator passes through the line 8 and enters the fractionation column 11 for separation, and the gas product, the light distillate oil and the heavy distillate are respectively discharged through the pipeline 12, the pipeline 13 and the pipeline 15, and the obtained diesel fraction is passed through the pipeline 14 and the pipeline.
  • the hydrogen of 17 After the hydrogen of 17 is mixed, it enters the first supplementary hydrotreating reactor 16 and is contacted with the low-activity hydrogenation catalyst A to carry out a hydrogenation reaction; the resulting reaction effluent is passed through line 18 to the second supplementary hydrotreating reactor 19, In the presence of hydrogen, in contact with the highly active hydrogenation catalyst B, the tetracyclic and small amounts of tricyclic aromatic hydrocarbons are saturated while maintaining the ring-shaped integrity of the cycloalkane after saturation of the polycyclic aromatic hydrocarbons, becoming a tricyclic or tetracyclic cycloalkane.
  • the effluent obtained after the supplementary hydrotreating is passed through the line 20 to the gas-liquid separator 21 (the gas-liquid separator 21 usually includes a high-pressure separator and a low-pressure separator), and the hydrogen-rich gas obtained after the separation passes through the line 22 and the line 23
  • the introduced fresh hydrogen is mixed to obtain recycled hydrogen; after the separation, the obtained liquid is subjected to an optional stripping treatment (omitted in the drawing), and is subjected to fractionation through the line 24 to the fractionation column 25, and a small amount of gas is discharged through the line 26, and the obtained is rich in ten.
  • the hexahydroquinone liquid enters the cooling and cooling, filtration and vacuum extraction unit 29 via line 28, and the obtained solid product, heptahydroquinone, is discharged through line 30; the fractionated low-boiling solvent oil is passed through line 27, high-boiling solvent oil through line 32, and After the extraction, the liquid obtained through the line 31 is mixed and used as a low aromatic solvent oil product.
  • another process flow of the present invention is: after the pretreatment (the pretreatment unit is omitted in the figure), the coal tar passes through the pipeline 1, and is mixed with the hydrogen passing through the pipeline 2 to enter the hydrotreating reactor. 3, to remove Hydrogenation reaction of sulfur, nitrogen, oxygen, metal, etc., the purification reaction effluent enters the hydrocracking reactor 5 through the pipeline 4 for cracking reaction, and the hydrocracking reaction effluent enters the gas-liquid separator 7 through the pipeline 6 (gas-liquid separator) 7 typically comprises a high pressure separator and a low pressure separator), and the resulting hydrogen rich gas is passed through line 10 and optionally after dehydrogenation treatment, and mixed with fresh hydrogen entering line 9 to provide recycled hydrogen.
  • the liquid obtained by the gas-liquid separator passes through the line 8 and enters the fractionation column 11 for separation.
  • the obtained gas product and the light distillate are discharged through the lines 12 and 13, respectively, and the resulting heavy distillate is mixed with the hydrogen from the line 16 through the line 14, and then enters.
  • the first supplementary hydrofinishing reactor 15 is contacted with the low activity hydrogenation catalyst A to carry out a hydrogenation reaction; the resulting reaction effluent is passed through line 17 to the second supplementary hydrotreating reactor 18, in the presence of hydrogen, with high activity.
  • Hydrogenation catalyst B is contacted to saturate the tetracyclic and minor tricyclic aromatic hydrocarbons while maintaining the integrity of the cycloalkane ring after saturation of the polycyclic aromatic hydrocarbons, becoming a tricyclic or tetracyclic cycloalkane.
  • the effluent obtained after the supplementary hydrotreating is passed through a line 19 to the gas-liquid separator 20 (the gas-liquid separator 20 usually includes a high-pressure separator and a low-pressure separator), and the hydrogen-rich gas obtained after the separation passes through the line 21 and the line 22
  • the introduced fresh hydrogen is mixed to obtain recycled hydrogen; after the separation, the liquid obtained is subjected to an optional stripping treatment (omitted in the drawing), enters the fractionation column 24 through the line 23 for fractionation, and a small amount of gas is discharged through the line 25, and the obtained is rich in ten.
  • the hexahydroquinone liquid enters the cooling and cooling, filtration and vacuum extraction unit 28 via line 27, and the resulting solid product, heptahydroquinone, is discharged via line 29; the fractionated low-boiling solvent oil is fractionated through line 26 and the liquid extracted by line 30. After mixing, as a low aromatic solvent oil product, the high boiling point solvent oil is returned to the hydrocracking reactor 5 via line 31 for cracking reaction to obtain more hexadecahydroquinone component.
  • the purity of hexadecahydroquinone is qualitatively analyzed by GC-MS gas chromatography-mass spectrometry, and the Saybolt color of the solvent oil is detected by the method of GB/T3555-1992, and the aromatic hydrocarbon content of the solvent oil is GB/T 17474.
  • the method detects that the content of lanthanide compounds in the product obtained by hydrocracking of coal tar is determined by the method of ISO13877-1998.
  • the high temperature coal tar raw materials used in the following examples and comparative examples are shown in Table 1 below.
  • the high temperature coal tar is obtained by dry distillation of coal produced in Anyang, Henan province, at 1000 ° C, and removal of naphthalene fraction obtained from naphthalene.
  • the supplementary hydrofinishing catalysts used in the following examples were all prepared according to the method disclosed in CN104588073A. Specifically, the properties of the supplementary hydrotreating catalyst used are shown in Table 2 below.
  • Catalyst A Catalyst B Active metal wt% Pt/Pd 0.12/0.28 0.25/0.55 Carrier Amorphous silica-alumina+alumina+Y molecular sieve Amorphous silica-alumina+alumina+Y molecular sieve Y molecular sieve, wt% 30 16 Y molecular sieve properties Average grain diameter, nm 370 370 Relative crystallinity, % 110 110 Secondary hole (1.7-10nm), % 62 62 SiO 2 /Al 2 O 3 molar ratio 85 85 Unit cell constant, nm 2.432 2.432 Specific surface area, m 2 /g 990 990 Porosity, mL/g 0.59 0.59
  • the supplementary hydrotreating catalyst used in the following comparative examples was prepared according to the method of CN104588073A, wherein the preparation of the small-grain Y-type molecular sieve was referred to Comparative Examples 1 and 2 in CN104588073A, specifically, the nature of the supplementary hydrofinishing catalyst used. As shown in Table 3 below.
  • the distillation range of the diesel fraction separated from the hydrocracking reaction effluent was 160 to 340 ° C, and the product distribution of the product obtained by the hydrocracking reaction included: ⁇ 160 °C fraction 8.3 wt%, 160-340 ° C fraction 55.5 wt%, > 340 ° C fraction 36.2 wt%, and 160-340 ° C fraction ruthenium content 1.5 wt%; from the supplemental hydrotreating reaction effluent
  • the distillation range of the liquid fraction containing hexadecahydroquinone is 280-320 °C.
  • the catalyst in the supplementary hydrotreating reaction zone is not segmented, but only one catalyst A is used, wherein the operating conditions of the hydrotreating reaction zone and the yield and purity of the obtained hexadecahydroquinone are as follows. Table 5 shows.
  • the catalyst used in the supplementary hydrotreating reaction zone is Catalyst C, wherein the operating conditions of the supplementary hydrotreating reaction zone and the yield and purity of the obtained hexadecahydroquinone are as shown in Table 5. Show.
  • the catalyst in the supplementary hydrotreating reaction zone is not segmented, but only one catalyst B is used, wherein the operating conditions of the hydrotreating reaction zone and the yield and purity of the obtained hexadecahydroquinone are as follows. Table 5 shows.
  • the catalyst used in the supplementary hydrotreating reaction zone is Catalyst D, wherein the operating conditions of the supplementary hydrotreating reaction zone and the yield and purity of the obtained hexadecahydroquinone are as shown in Table 5. Show.
  • the supplementary hydrotreating reaction is divided into two supplementary hydrotreating reaction stages, and the grading scheme of catalyst A and catalyst B is adopted, and the first supplementary hydrotreating reaction stage uses catalyst A, the second supplement Catalyst B was used in the hydrotreating reaction stage, wherein the operating conditions of the supplementary hydrotreating reaction zone and the yield and purity of the obtained hexadecahydroquinone are shown in Table 5.
  • the catalyst grade used in the supplementary hydrotreating reaction zone is Catalyst C and Catalyst D
  • the first supplemental hydrofining reaction stage uses Catalyst C
  • the second supplementary hydrofining reaction stage is used.
  • Catalyst D wherein the operating conditions of the supplementary hydrotreating reaction zone and the yield and purity of the resulting hexadecahydroquinone are shown in Table 5.
  • Example 3 the difference is that the gradation sequence of the supplementary hydrotreating catalyst is changed, the first supplementary hydrotreating reaction stage uses the catalyst B, and the second supplementary hydrofining reaction stage uses the catalyst A, wherein the supplementary hydrogenation
  • Table 5 The operating conditions of the refining reaction zone and the yield and purity of the resulting hexadecahydroquinone are shown in Table 5.
  • the coal tar raw material in the present invention can obtain a higher purity hexadecahydroquinone product through the processes of pretreatment, hydrotreating, hydrocracking, and supplemental hydrofining.
  • the supplementary refining reaction zone adopts a catalyst grading scheme, and the obtained hexadecahydroquinone has higher yield and purity, and has a more ideal hydrogenation effect.
  • the initial fraction of the heavy fraction separated from the hydrocracking reaction effluent was 160 ° C, and the product distribution of the product obtained by the hydrocracking reaction included: ⁇ 160 ° C
  • the fraction is 8.4 wt%, the fraction above 160 ° C is 91.6 wt%, and the niobium content in the fraction above 160 ° C is 1.2 wt%;
  • the distillation range of the liquid fraction rich in hexadecahydroquinone separated from the supplement hydrotreating reaction effluent is 250 to 340 ° C.
  • the catalyst in the supplementary hydrotreating reaction zone is not segmented, but only one catalyst A is used, wherein the operating conditions of the hydrotreating reaction zone and the yield and purity of the obtained hexadecahydroquinone are as follows. Table 7 shows.
  • the catalyst used in the supplementary hydrotreating reaction zone is Catalyst C, wherein the operating conditions of the supplementary hydrotreating reaction zone and the yield and purity of the obtained hexadecahydroquinone are as shown in Table 7. Show.
  • the catalyst in the supplementary hydrotreating reaction zone is not segmented, but only one catalyst B is used, wherein the operating conditions of the hydrotreating reaction zone and the yield and purity of the obtained hexadecahydroquinone are as follows. Table 7 shows.
  • the catalyst used in the supplementary hydrotreating reaction zone is Catalyst D, wherein the operating conditions of the supplementary hydrotreating reaction zone and the yield and purity of the obtained hexadecahydroquinone are shown in Table 7. Show.
  • the supplementary hydrotreating reaction is divided into two supplementary hydrofining reaction stages, and the catalysis scheme of catalyst A and catalyst B is adopted, and the first supplementary hydrotreating reaction stage uses catalyst A, the second supplement Catalyst B was used in the hydrotreating reaction stage, wherein the operating conditions of the supplementary hydrotreating reaction zone and the yield and purity of the obtained hexadecahydroquinone are shown in Table 7.
  • Example 7 except that the catalyst grade used in the supplementary hydrotreating reaction zone is Catalyst C and Catalyst D, the first supplementary hydrofining reaction stage uses Catalyst C, and the second supplementary hydrofining reaction stage is used.
  • Catalyst D wherein the operating conditions of the supplementary hydrotreating reaction zone and the yield and purity of the resulting hexadecahydroquinone are shown in Table 7.
  • Example 7 the difference is that the gradation sequence of the supplementary hydrotreating catalyst is changed, the first supplementary hydrofining reaction stage uses the catalyst B, and the second supplementary hydrofining reaction stage uses the catalyst A, wherein the supplementary hydrogenation
  • Table 7 The operating conditions of the refining reaction zone and the yield and purity of the resulting hexadecahydroquinone are shown in Table 7.
  • the process of pretreatment, hydrorefining, hydrocracking, and supplemental hydrofining of the coal tar raw material in the present invention can obtain the heptahydroquinone product with higher purity.
  • the supplementary refining reaction zone adopts a catalyst grading scheme, and the obtained hexadecahydroquinone has higher yield and purity, and has a more ideal hydrogenation effect.

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Abstract

一种十六氢芘的制备方法，包括：在加氢催化剂的存在下，将含有芘系化合物的烃油原料进行加氢反应，其中，所述芘系化合物选自芘及其不饱和加氢产物中的至少一种，所述加氢催化剂含有载体和负载于载体上的活性金属组分，所述活性金属组分为Pt和/或Pd，所述载体含有小晶粒Y型分子筛、氧化铝和无定形硅铝，所述小晶粒Y型分子筛的晶粒平均直径为200-700nm，SiO 2/Al 2O 3摩尔比为40-120，相对结晶度≥95％，比表面积为900-1200m 2/g，1.7-10nm的二次孔的孔容占总孔容的50％以上，可以制得高纯度的十六氢芘产品，并且收率较高。

Description

一种十六氢芘的制备方法 技术领域
本发明涉及一种十六氢芘的制备方法。
背景技术
高纯度十六氢芘的生产工艺苛刻，价格昂贵，因此，满足市场要求的高纯度十六氢芘产品成为稀有资源。
目前，合成的方法是生产高纯度十六氢芘的重要途径之一。芘是煤焦油中的重要组分，在煤焦油初馏时富集于300～360℃的馏分中，通过传统的精馏、结晶等方法可获取。煤焦油原料中基本不含有十六氢芘，直接从煤焦油中提取十六氢芘还未见报道，操作费用大，提取出的十六氢芘纯度低、收率低，经济上不可行。1,2,3,6,7,8-六氢芘的制备一般主要是通过芘的选择性催化加氢制得，由于连串反应和平行反应同时进行，从而使单一产物的选择性很难提高，获取高纯度十六氢芘的方法更难。
CN1351130A公开了一种煤焦油加氢生产柴油的方法，主要是煤焦油经过分馏，得到的柴油以下的馏分进行加氢精制，可直接生产符合燃油指标的柴油或者生产作为柴油产品的调和组分，但是它只是对煤焦油中比较轻质的馏分进行加氢处理，并不能提取出高附加值、高纯度的十六氢芘产品，没有完全利用煤焦油。
CN1676583A公开了一种中高温煤焦油加氢裂化工艺。工艺过程为：中高温煤焦油经加热炉加热到250～300℃，与氢气混合进入加氢精制反应器，精制生成油经蒸馏装置，分馏出汽油、柴油、润滑油及加氢尾油，加氢尾油经裂化加热炉加热后，与氢气混合后进入裂化反应器，进一步生产汽柴油馏分。馏分中不能提取出十六氢芘产品，只是生产调和燃料油，该工艺煤焦油直接进入高温加热炉会导致炉管结焦，影响装置的正常运转周期。
发明内容
本发明的目的是为了提供一种十六氢芘的制备方法，采用该方法可以制备高纯度的十六氢芘产品。
在现有技术中，虽然有研究对粗芘加氢进行了大量的探索，但其通常仅能够得到六氢芘，而未有得到十六氢芘的记载。本发明的发明人通过研究意外发现，在加氢催化 剂中，采用具有高硅铝比、高结晶度、二次孔多、大比表面积的小晶粒Y型分子筛与无定形硅铝以及活性金属组分Pt和Pd相互配合，可以对粗芘原料(含有芘系化合物的烃油原料)进行全加氢，并且具有较高的催化活性、选择性和稳定性，从而完成了本发明。
本发明提供了一种十六氢芘的制备方法，该方法包括：在加氢催化剂的存在下，将含有芘系化合物的烃油原料进行加氢反应，其中，所述芘系化合物选自芘及其不饱和加氢产物中的至少一种，所述加氢催化剂含有载体和负载于载体上的活性金属组分，所述活性金属组分为Pt和/或Pd，所述载体含有小晶粒Y型分子筛、氧化铝和无定形硅铝，所述小晶粒Y型分子筛的晶粒平均直径为200-700nm，SiO2/Al2O3摩尔比为40-120，相对结晶度≥95％，比表面积为900-1200m2/g，1.7-10nm的二次孔的孔容占总孔容的50％以上。
按照本发明所述的制备十六氢芘的方法，可以制得高纯度的十六氢芘产品。
附图说明
图1是本发明所述的十六氢芘的制备方法的一种实施方式的流程示意图。
图2是本发明所述的十六氢芘的制备方法的另一种实施方式的流程示意图。
具体实施方式
在本文中所披露的范围的端点和任何值都不限于该精确的范围或值，这些范围或值应当理解为包含接近这些范围或值的值。对于数值范围来说，各个范围的端点值之间、各个范围的端点值和单独的点值之间，以及单独的点值之间可以彼此组合而得到一个或多个新的数值范围，这些数值范围应被视为在本文中具体公开。
本发明所述的十六氢芘的制备方法包括：在加氢催化剂的存在下，将含有芘系化合物的烃油原料进行加氢反应。
在本发明中，所述加氢催化剂含有载体和负载于载体上的活性金属组分。所述活性金属组分为Pt和/或Pd。所述载体含有小晶粒Y型分子筛、氧化铝和无定形硅铝。
所述小晶粒Y型分子筛的晶粒平均直径为200-700nm，具体地，例如可以为200nm、250nm、300nm、350nm、400nm、450nm、500nm、550nm、600nm、650nm、700nm以及这些点值中的任意两个所构成的范围中的任意值。在优选情况下，所述小晶粒Y型分子筛的晶粒平均直径为300-500nm。在本发明中，所述小晶粒Y型分子筛的晶粒平均直径根据SEM(扫描电子显微镜)法检测。
所述小晶粒Y型分子筛的SiO2/Al2O3摩尔比为40-120，具体地，例如可以为40、50、60、70、80、90、100、110、120以及这些点值中的任意两个所构成的范围中的任意值。
所述小晶粒Y型分子筛的相对结晶度≥95％，优选为95-120％，更优选为98-115％。在本发明中，所述小晶粒Y型分子筛的相对结晶度根据X光衍射法检测。
所述小晶粒Y型分子筛的比表面积为900-1200m2/g，具体地，例如可以为900m2/g、920m2/g、950m2/g、980m2/g、1000m2/g、1020m2/g、1050m2/g、1080m2/g、1100m2/g、1120m2/g、1150m2/g、1180m2/g、1200m2/g以及这些点值中的任意两个所构成的范围中的任意值。在本发明中，所述小晶粒Y型分子筛的比表面积根据低温液氮物理吸附法检测。
所述小晶粒Y型分子筛具有较多的二次孔，具体地，在所述小晶粒Y型分子筛中，1.7-10nm的二次孔的孔容占总孔容的50％以上，优选为50-80％，进一步优选为60-80％。在本发明中，所述小晶粒Y型分子筛的二次孔的孔容根据低温液氮物理吸附法检测。
所述小晶粒Y型分子筛的晶胞常数可以为2.425-2.435nm，例如2.425nm、2.426nm、2.427nm、2.428nm、2.429nm、2.43nm、2.431nm、2.432nm、2.433nm、2.434nm、2.435nm以及这些点值中的任意两个所构成的范围中的任意值。在本发明中，所述小晶粒Y型分子筛的晶胞常数根据X光衍射法检测。
所述小晶粒Y型分子筛的孔容可以为0.5-0.8mL/g，例如0.5mL/g、0.55mL/g、0.6mL/g、0.65mL/g、0.7mL/g、0.75mL/g、0.8mL/g以及这些点值中的任意两个所构成的范围中的任意值。在本发明中，所述小晶粒Y型分子筛的孔容根据低温液氮物理吸附法检测。
在本发明中，所述加氢催化剂的性质如下：比表面积可以为350-550m2/g，优选为380-500m2/g；孔容可以为0.5-1mL/g，优选为0.5-0.9mL/g。
在所述加氢催化剂中，以所述加氢催化剂的总重量为基准，所述活性金属组分的含量可以为0.1-2重量％，优选为0.2-1.5重量％；所述载体的含量可以为98-99.9重量％，优选为98.5-99.8重量％。
在所述载体中，以所述载体的总重量为基准，所述小晶粒Y型分子筛的含量可以为5-40重量％，优选为10-25重量％；所述氧化铝的含量可以为10-40重量％，优选为15-30重量％；所述无定形硅铝的含量可以为20-65重量％，优选为30-60重量％。
在本发明中，所述加氢催化剂可以选择适宜的商品催化剂，也可以根据本领域的 常规方法进行制备，如按照CN104588073A中报道的方法制得。具体地，所述加氢催化剂的制备方法可以包括：将小晶粒Y型分子筛、无定型硅铝和用氧化铝制成的粘合剂机械混合、成型，然后干燥和焙烧，制成催化剂载体；采用浸渍法在载体上负载Pt和/或Pd，经干燥和焙烧，得到加氢催化剂。
所述小晶粒Y型分子筛的制备方法可以包括以下步骤：
(1)将小晶粒NaY分子筛制成Na2O含量≤2.5重量％的小晶粒NH4NaY分子筛；
(2)将小晶粒NH4NaY分子筛进行水热处理，然后用六氟硅酸铵水溶液进行脱铝补硅；
(3)将步骤(3)得到的分子筛用含NH4 +和H+和的混合溶液处理，然后洗涤和干燥，得到小晶粒Y型分子筛。
所述小晶粒NaY分子筛的性质如下：SiO2/Al2O3摩尔比大于6且不高于9，优选为6.5-9，进一步优选为7-8；晶粒平均直径为200-700nm，优选为300-500nm；比表面积为800-1000m2/g，优选为850-950m2/g；孔容0.3-0.45mL/g，相对结晶度为90-130％，晶胞常数为2.46-2.47，经650℃空气中焙烧3小时后相对结晶度为90％以上，优选为90-110％，更优选为90-105％。
根据本发明的一种优选实施方式，所述加氢反应的过程包括依次进行的两个反应阶段，与之相对应，第一反应阶段使用的催化剂为加氢催化剂A，第二反应应阶段使用的催化剂为加氢催化剂B。所述加氢催化剂A中的活性金属组分的含量百分比x1低于所述加氢催化剂B中的活性金属组分的含量百分比x2，优选地x1比x2低0.1-1.5个百分点，更优选地x1比x2低0.3-1.5个百分点。所述加氢催化剂A中的小晶粒Y型分子筛的含量百分比y1高于所述加氢催化剂B中的小晶粒Y型分子筛的含量百分比y2，优选地y1比y2高5-35个百分点，更优选地y1比y2高10-35个百分点。按照上述优选实施方式可以获得较高纯度的十六氢芘，并且收率更高。
在本发明所述的方法中，所述加氢反应的工艺条件可以为本领域的常规选择。在优选情况下，所述加氢反应的工艺条件包括：氢分压为4-20MPa，液时体积空速为0.05-6h-1，氢油体积比为50-3000，平均反应温度为150-380℃。
在进一步优选的实施方式中，当所述加氢反应的过程包括依次进行的两个反应阶段时，所述第二反应阶段的平均反应温度比所述第一反应阶段的平均反应温度低10-150℃，优选30-120℃。更优选地，所述第一反应阶段的平均反应温度为180-380℃，进一步优选为220-350℃；所述第二反应阶段的平均反应温度为150-350℃，优选为 180-330℃。
在上述优选实施方式中，所述第一反应阶段和所述第二反应阶段可以在同一个反应器内进行，也可以在串联的两个以上反应器内进行。
在本发明中，所述芘系化合物选自芘及其不饱和加氢产物中的至少一种。芘的不饱和加氢产物例如可以为二氢芘、四氢芘、六氢芘、八氢芘等。
在所述含有芘系化合物的烃油原料中，所述芘系化合物的含量可以为0.5重量％以上，具体地，例如可以为0.5-10重量％，如0.5重量％、0.8重量％、1.0重量％、1.2重量％、1.5重量％、2重量％、3重量％、4重量％、5重量％、6重量％、7重量％、8重量％、9重量％、10重量％。
在本发明中，所述含有芘系化合物的烃油原料可以为本领域常规的烃油原料，只要其中含有预定量的芘系化合物即可。在一种实施方式中，所述含有芘系化合物的烃油原料是初馏点为130-220℃(优选为160-200℃)的重馏分油。在优选情况下，所述含有芘系化合物的烃油原料是初馏点为130-220℃且终馏点为300-400℃的重馏分油。进一步优选地，所述含有芘系化合物的烃油原料为柴油馏分，且初馏点为160-200℃，终馏点为300-350℃。
根据本发明的一种优选实施方式，所述含有芘系化合物的烃油原料按照包括以下步骤的方法制备得到：
(1)将煤焦油原料进行预处理；
(2)将经过预处理的煤焦油原料进行加氢精制反应；
(3)将所得反应流出物可选地进行分离，然后进行加氢裂化反应；
(4)将加氢裂化反应所得反应流出物进行气液分离，然后对分离出的液相进行分馏，将分馏出的重馏分作为所述含有芘系化合物的烃油原料。
所述的煤焦油可以为低温煤焦油、中温煤焦油或高温煤焦油中的至少一种，也可以是煤焦油提取出萘、蒽、菲、咔唑和荧蒽中的至少一种后的剩余馏分。煤焦油的芳烃含量一般为20～100wt％，20℃密度一般为1.023～1.235g/cm3。煤焦油的馏程为200～700℃范围内的任意范围，一般初馏点至终馏点的温度差在100～400℃之间。优选情况下，所述煤焦油原料为高温煤焦油或者高温煤焦油提取出蒽、菲、咔唑和荧蒽中的至少一种后的剩余馏分。
步骤(1)所述的预处理通常包括机械脱杂质、脱水、电脱盐等操作，还可以任选地包括抽提脱除蒽、菲等操作。
步骤(2)中所述的加氢精制反应过程中所使用的催化剂可以为本领域常规的加氢精制催化剂，如可以为柴油加氢精制催化剂或加氢裂化预处理催化剂。加氢精制催化剂一般以ⅥB族和/或第Ⅷ族金属为活性组分，以氧化铝或含硅氧化铝为载体，第ⅥB族金属一般为Mo和/或W，第Ⅷ族金属一般为Co和/或Ni。以催化剂的重量计，第ⅥB族金属含量以氧化物计为10～50重量％，第Ⅷ族金属含量以氧化物计为3～15重量％；其性质如下：比表面积为100～350m2/g，孔容为0.15～0.6mL/g。可选择的商品催化剂有中国石油化工股份有限公司抚顺石油化工研究院研制开发的3936、3996、FF-16、FF-26、FF-36、FF-46、FF-56、FF-66等加氢精制催化剂，也可以是UOP公司的HC-K、HC-P催化剂，Topsoe公司的TK-555、TK-565催化剂，以及AKZO公司的KF-847、KF-848等。
步骤(2)中所述的加氢精制反应的工艺条件一般为：氢分压3～19MPa，平均反应温度为260～440℃，液时体积空速为0.1～4h-1，氢油体积比为300:1～3000:1。步骤(2)所述的加氢精制可以选择固定床、沸腾床等本领域的常规反应器形式。固定床反应器可以为上流式(并流)反应器、下流式(并流)反应器或气液逆流反应器形式。
步骤(3)中所述的加氢裂化反应过程中使用的催化剂可以为本领域常规的加氢裂化催化剂，如可以为轻油型加氢裂化催化剂、灵活型加氢裂化催化剂和(高)中油型加氢裂化催化剂。加氢裂化催化剂一般以第ⅥB族和/或第Ⅷ族金属为活性组分，第ⅥB族金属一般为Mo和/或W，第Ⅷ族金属一般为Co和/或Ni。催化剂的载体为氧化铝、含硅氧化铝和分子筛中的一种或多种。以催化剂的重量计，第ⅥB族金属含量以氧化物计为10～35重量％，第Ⅷ族金属含量以氧化物计为3～15重量％；分子筛含量为5～40重量％，氧化铝含量为10～80重量％；其比表面积为100～650m2/g，孔容为0.15～0.50mL/g。可供选择的商品催化剂有中国石油化工股份有限公司抚顺石油化工研究院研制开发的FC-26、FC-28、FC-14、ZHC-01、ZHC-02、ZHC-04等单段加氢裂化催化剂，也可以选择UOP公司的DHC39、DHC-8，CHERON公司的ICR126等加氢裂化催化剂。其中ZHC-02、ICR126为以无定形硅铝和Y型分子筛为裂化组分的加氢裂化催化剂，更适宜于本发明的加氢裂化反应过程。
本发明中优选使用(高)中油型加氢裂化催化剂。为了提高产品的收率和选择性，本发明中特别选择中油型加氢裂化催化剂(如FC-26催化剂)。该催化剂在加氢条件下对烷烃和带侧链芳烃具有较好的断链功能，能够将原料中环状烃(包括环烷烃、侧链的环烷烃、芳烃、带侧链的芳烃)中的侧链烷烃断开。同时该催化剂具有适宜的稠环芳烃 (不带侧链)饱和功能，且几乎不具有开环作用。加氢裂化所得生成油再经过分馏，可以将含有目的产品前驱物的组分尽量富集在适宜的窄馏分中。因此，使用中油型加氢裂化催化剂能够最大量的保持产品中环状烃的含量，而有助于提高最终目的产品的收率。
加氢裂化系统所用的反应器为常规固定床加氢反应器，更优选为下流式固定床反应器。步骤(3)中加氢裂化的工艺条件一般为：氢分压3～19MPa，平均反应温度为260～440℃，液时体积空速为0.3～4h-1，氢油体积比300:1～5000:1。
在步骤(3)中，“可选的”是指可以经过分离(如气液分离)过程，也可以不经过分离(如气液分离)过程。
步骤(4)所述的分馏操作可以选择本领域的常规技术。分馏所得重馏分的初馏点可以为130-220℃。优选为160-200℃。在优选情况下，步骤(4)分馏所得重馏分为柴油馏分，进一步优选地，柴油馏分的初馏点为130～220℃，更优选为160～200℃；终馏点为280～400℃，更优选为300～350℃。
在本发明中，所述方法还可以包括从步骤(4)所得到重馏分中去除石脑油馏分，并将得到的剩余液体馏分作为所述含有芘系化合物的烃油原料。
在本发明中，所述方法还可以包括将加氢反应所得反应流出物进行分离和分馏，得到富含十六氢芘组分和重组分，将所述重组分的至少部分循环回上述步骤(3)中进行加氢裂化反应。此处的分馏过程可以采用本领域的常规分馏技术。分馏所得产品除了富含十六氢芘组分(中间组分)和重组分外，还包括液体轻组分。其中所述液体轻组分和中间组分(富含十六氢芘组分)的切割温度为130～280℃，优选为200～260℃。所述中间组分和重组分的切割温度为300～360℃，优选为320～340℃。液体中间组分经过冷却降温后，并经过滤、真空抽提、可选择的离心分离操作后得到固体即为高纯度十六氢芘产品，经过分析其纯度可达95wt％以上。而所得到高于切割温度的液体重组分由于含有二苯并芘、茚并芘等五环以上烃类，可以通过循环加氢裂化后将其转化为芘，从而可以提高目的产品的收率。
在本发明中，为了获得高纯度的十六氢芘产品，所述方法还包括将加氢反应所得反应流出物进行分离和分馏，得到富含十六氢芘组分和重组分，将所述富含十六氢芘组分进行冷却降温，然后经过滤、抽提得到固体的十六氢芘。此处的分馏过程可以采用本领域的常规分馏技术。通过分馏所得富含十六氢芘液体馏分的初馏点一般为220～300℃，优选为260～280℃；终馏点一般为＞300～360℃(大于300℃且小于等于360℃)， 优选为320～340℃。将该液体馏分进行冷却降温，生成的十六氢芘将在液体中结晶析出，再经过滤、可选择的离心分离操作，即得到高纯度的十六氢芘产品。
根据本发明的一种具体实施方式，所述十六氢芘的制备方法包括：
(1)将煤焦油原料进行预处理；
(2)步骤(1)得到的煤焦油与氢气混合后，进入加氢精制反应区，与加氢精制催化剂接触进行反应；
(3)步骤(2)所得加氢精制反应流出物，可选地经过分离后，与氢气一起进入加氢裂化反应区，与加氢裂化催化剂接触进行反应；
(4)加氢裂化流出物进行气液分离，液体经过分馏，得到重馏分，重馏分的初馏点为130～220℃；
(5)步骤(4)所得重馏分与氢气混合后，进入本发明所述的加氢催化剂接触进行加氢反应(即补充加氢精制反应)；
(6)步骤(5)得到反应流出物经过分离、分馏后，得到富含十六氢芘组分和重组分，富含十六氢芘组分经冷却降温后，并经过滤、真空抽提，所得固体即为十六氢芘产品。
进一步优选地，上述步骤(5)的加氢反应过程包括依次进行的两个反应阶段，与之相对应，第一反应阶段使用的催化剂为加氢催化剂A，第二反应应阶段使用的催化剂为加氢催化剂B。所述加氢催化剂A中的活性金属组分的含量百分比x1低于所述加氢催化剂B中的活性金属组分的含量百分比x2，优选地x1比x2低0.1-1.5个百分点，更优选地x1比x2低0.3-1.5个百分点。所述加氢催化剂A中的小晶粒Y型分子筛的含量百分比y1高于所述加氢催化剂B中的小晶粒Y型分子筛的含量百分比y2，优选地y1比y2高5-35个百分点，更优选地y1比y2高10-35个百分点。
在所述第一反应阶段和所述第二反应阶段中，所使用的加氢催化剂由于活性金属组分含量和小晶粒Y型分子筛含量的不同，而分别具有不同的性质。所述加氢催化剂A具有相对较低的活性金属组分含量和更高的Y型分子筛含量，因而催化剂的裂化性能较高。煤焦油加氢裂化所得柴油馏分中仍然含有的链烷烃和带侧链的多环芳烃进一步与所述加氢催化剂A接触反应，多环芳烃上的侧链经过反应后几乎全部从芳环上剥离，同时多环芳烃会发生部分双键饱和反应，如粗芘经加氢后可以生成六氢芘。所述加氢催化剂B则因为具有相对较高的活性金属组分含量和相对较低的小晶粒Y型分子筛含量，而加氢性能较高，裂解活性偏弱。所述第一反应阶段的加氢产物再与所述加氢催化剂B接触 反应时，其中经过部分加氢生成的非全氢芘(如六氢芘)在更低的反应温度下，由于催化剂的裂化活性受到限制，而饱和功能强，进一步进行加氢，全部的碳碳双键均得到饱和从而得到十六氢芘(全氢芘)产品，从而提高了本发明方法生产十六氢芘的收率。因此，在加氢反应过程中按照上述优选方式对加氢催化剂进行级配组合，很好地实现了柴油馏分中稠环芳烃尤其是粗芘的加氢饱和，从而使得加氢法能够用于直接生产高纯度的十六氢芘产品。
而且，与现有技术相比，本文中上述具体实施方式的工艺方法具有以下特点：
(1)按照本发明所述的方法，使用煤焦油原料作为初始原料，并选择适宜的工艺流程，通过加氢工艺能够制备出高纯度的十六氢芘，同时还能够得到性能优异的溶剂油产品。本发明所述的方法极大地扩宽了煤焦油生产高附加值产品的潜力。既为低附加值的煤焦油提供了一种提高其经济性的加工方法，也为十六氢芘产品开发了一种新原料和全新的工艺路线。
(2)本发明首先通过加氢精制、加氢裂化和分馏过程，针对煤焦油加氢裂化柴油中芳烃含量高，在加氢裂化过程中生成较多易溶解十六氢芘的轻重组分的情况，将加氢裂化生成油分馏出富含十六氢芘的窄馏分，从而将三环以上稠环芳烃组分富集到柴油馏分中，并实现了对相溶目的产品的组分的分隔(分离)，降低易溶组分对后续补充精制反应的影响；之后通过补充加氢过程，利用含有小晶粒Y型分子筛的贵金属催化剂对稠环芳烃的选择性裂解和加氢能力，将四环芳烃——粗芘进行全加氢，从而得到富含全氢芘(即十六氢芘)的馏分油；最后再通过分馏过程对全氢芘进行富集，经过冷却降温后十六氢芘从馏分油中析出结晶。
(3)本发明中使用的加氢催化剂，采用小晶粒Y型分子筛作为酸性组分，该Y型分子筛具有高硅铝比、高结晶度、二次孔多、大比表面积的特点，其与无定形硅铝以及加氢活性金属组分Pt和Pd相互配合，不仅促进了芳烃加氢饱和活性的发挥，更有利于芳烃的选择性开环和断链，而且有利于反应产物的扩散，同时容炭能力也大为增强，从而提高了催化剂的活性、选择性和稳定性。该催化剂特别适宜作为环烷基原料，尤其是粘度高、稠环芳烃含量高的环烷基原料的加氢脱芳烃反应中。
(4)本发明中，在补充加氢反应区内，优选采用两种不同的加氢催化剂。第一反应阶段内使用的加氢催化剂A具有相对较高的Y型分子筛含量和相对较低的金属含量，因而表现为偏裂解活性。对于富集了三环以上稠环芳烃的加氢裂化柴油馏分而言，加氢催化剂A的偏裂解性能能够对带侧链烷烃的稠环芳烃进行有效的断链反应，进一步将侧 链从芳环上剥离。而第二反应阶段内使用的加氢催化剂B，则因为具有相对较高的金属含量和较低的Y型分子筛含量，而加氢性能偏强，而其适宜的裂解活性亦对稠环芳烃的加氢饱和具有重要的催化作用。因此，在第一反应阶段内已经发生部分加氢而生成的非全氢芘(如六氢芘)可以在更低的反应温度下，在第二反应阶段完成全部芳环的加氢饱和，从而得到十六氢芘(全氢芘)产品。
(5)本发明对加氢裂化所得柴油馏分采用缓和的逐步饱和各种杂环芳烃的加工方案用于生产十六氢芘产品和低芳溶剂油，可以最大程度地避免单段工艺中在高温下煤焦油加氢裂化柴油馏分中芳烃缩合积炭和裂解反应，而严重影响催化剂使用寿命问题的发生。
以下结合附图和实施例对本发明所述的方法做更详细的说明。
如图1所示，本发明的一种工艺流程为：经过预处理(图中省略了预处理单元)后的煤焦油经过管线1，与经过管线2的氢气混合后进入加氢精制反应器3，进行脱除硫、氮、氧、金属等加氢反应，精制反应流出物经过管线4进入加氢裂化反应器5进行裂化反应，加氢裂化反应流出物经过管线6进入气液分离器7(气液分离器7通常包括高压分离器与低压分离器)，所得富氢气体经过管线10和任选地经过脱硫化氢处理后，与管线9进入的新鲜氢气混合，得到循环氢。气液分离器所得液体经过管线8，进入分馏塔11进行分离，得到气体产物、轻馏分油和重馏分油分别经管线12、管线13和管线15排出，所得柴油馏分经管线14，与来自管线17的氢气混合后，进入第一补充加氢精制反应器16，与低活性加氢催化剂A接触，进行加氢反应；所得反应流出物经管线18进入第二补充加氢精制反应器19，在氢气存在下，与高活性加氢催化剂B接触，对四环和少量三环芳烃进行饱和，同时保持多环芳烃饱和后环烷烃环形的完整，变成带三环或四环环烷烃。补充加氢精制后所得流出物经过管线20进入气液分离器21(气液分离器21通常包括高压分离器与低压分离器)进行分离，分离后得到的富氢气体经过管线22，与管线23引入的新鲜氢气进行混合得到循环氢；分离后得到的液体经过可选的气提处理后(图中省略)，经过管线24进入分馏塔25进行分馏，少量气体经管线26排出，所得富含十六氢芘液体经过管线28进入冷却降温、过滤和真空抽提单元29，所得固体产品十六氢芘经管线30排出；分馏所得低沸点溶剂油经管线27、高沸点溶剂油经管线32，与抽提后经过管线31所得液体混合后，作为低芳溶剂油产品。
如图2所示，本发明的另一种工艺流程为：经过预处理(图中省略了预处理单元)后的煤焦油经过管线1，与经过管线2的氢气混合后进入加氢精制反应器3，进行脱除 硫、氮、氧、金属等加氢反应，精制反应流出物经过管线4进入加氢裂化反应器5进行裂化反应，加氢裂化反应流出物经过管线6进入气液分离器7(气液分离器7通常包括高压分离器与低压分离器)，所得富氢气体经过管线10和任选地经过脱硫化氢处理后，与管线9进入的新鲜氢气混合，得到循环氢。气液分离器所得液体经过管线8，进入分馏塔11进行分离，所得气体产物和轻馏分油分别经管线12和13排出，所得重馏分油经管线14，与来自管线16的氢气混合后，进入第一补充加氢精制反应器15，与低活性加氢催化剂A接触，进行加氢反应；所得反应流出物经管线17进入第二补充加氢精制反应器18，在氢气存在下，与高活性加氢催化剂B接触，对四环和少量三环芳烃进行饱和，同时保持多环芳烃饱和后环烷烃环形的完整，变成带三环或四环环烷烃。补充加氢精制后所得流出物经过管线19进入气液分离器20(气液分离器20通常包括高压分离器与低压分离器)进行分离，分离后得到的富氢气体经过管线21，与管线22引入的新鲜氢气进行混合得到循环氢；分离后得到的液体经过可选择的气提处理后(图中省略)，经过管线23进入分馏塔24进行分馏，少量气体经管线25排出，所得富含十六氢芘液体经过管线27进入冷却降温、过滤和真空抽提单元28，所得固体产品十六氢芘经管线29排出；分馏所得低沸点溶剂油经管线26与经管线30排除的抽提所得液体混合后，作为低芳溶剂油产品，高沸点溶剂油经管线31返回到加氢裂化反应器5进行裂化反应，以获取更多的十六氢芘组分。
本发明中，十六氢芘纯度采用GC-MS气相色谱-质谱小分子定性分析，溶剂油的赛波特颜色采用GB/T3555-1992的方法检测，溶剂油的芳烃含量采用GB/T 17474的方法检测，煤焦油加氢裂化得到产物中芘系化合物的含量采用ISO13877-1998方法进行测定。
以下实施例和对比例中使用的高温煤焦油原料如下表1所示。所述高温煤焦油由产自河南安阳的煤炭经过1000℃干馏，并脱除萘后所得煤焦油馏分。
表1
  高温煤焦油
密度(20℃)/kg.m-3 1023.1
馏程，℃ 320～550
芳烃含量，wt％ 52.3
三环以上芳烃，wt％ 43
十六氢芘，wt％ 0
凝点，℃ 32
硫，μg/g 3000
氮，μg/g 15000
以下实施例中使用的补充加氢精制催化剂均根据CN104588073A公开的方法进行制备。具体地，所使用的补充加氢精制催化剂的性质如下表2所示。
表2
项目 催化剂A 催化剂B
活性金属，wt％    
Pt/Pd 0.12/0.28 0.25/0.55
载体 无定形硅铝+氧化铝+Y分子筛 无定形硅铝+氧化铝+Y分子筛
Y分子筛，wt％ 30 16
Y分子筛性质    
晶粒平均直径，nm 370 370
相对结晶度，％ 110 110
二次孔(1.7-10nm)，％ 62 62
SiO2/Al2O3摩尔比 85 85
晶胞常数，nm 2.432 2.432
比表面积，m2/g 990 990
孔容，mL/g 0.59 0.59
以下对比例中使用的补充加氢精制催化剂根据CN104588073A的方法制备，其中，小晶粒Y型分子筛的制备参照CN104588073A中的对比例1和2，具体地，所使用的补充加氢精制催化剂的性质如下表3所示。
表3
项目 催化剂C 催化剂D
活性金属，wt％    
Pt/Pd 0.12/0.28 0.25/0.55
载体 无定形硅铝+氧化铝+Y分子筛 无定形硅铝+氧化铝+Y分子筛
Y分子筛，wt％ 30 16
Y分子筛性质    
晶粒平均直径，nm 400 450
相对结晶度，％ 95 80
二次孔(1.7-10nm)，％ 37.1 27.5
SiO2/Al2O3摩尔比 50 25
晶胞常数，nm 2.441 2.450
比表面积，m2/g 892 780
孔容，mL/g 0.33 0.32
在以下实施例1-4和对比例1-3中，加氢精制反应和加氢裂化反应的操作条件如下表4所示。
表4
  加氢精制 加氢裂化
催化剂 FF-36 FC-26
氢分压，MPa 15.0 15.0
液时空速，h-1 0.6 0.6
氢油体积比，v/v 1500 1500
反应温度，℃ 340 360
氮含量，μg/g 12 5
在以下实施例1-4和对比例1-3中，从加氢裂化反应流出物中分离出的柴油馏分的馏程为160～340℃，加氢裂化反应所得产物的产品分布包括：<160℃馏分8.3wt％，160-340℃馏分55.5wt％，>340℃馏分36.2wt％，且160-340℃馏分中芘含量为1.5wt％；从补充加氢精制反应流出物中分离出的富含十六氢芘液体馏分的馏程为280～320℃。
实施例1
如图1所示流程，补充加氢精制反应区内催化剂不分段，而仅采用一种催化剂A，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表5所示。
对比例1
参照实施例1，所不同的是，补充加氢精制反应区内采用的催化剂为催化剂C，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表5所示。
实施例2
如图1所示流程，补充加氢精制反应区内催化剂不分段，而仅采用一种催化剂B，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表5所示。
对比例2
参照实施例2，所不同的是，补充加氢精制反应区内采用的催化剂为催化剂D，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表5所示。
实施例3
如图1所示流程，补充加氢精制反应区分为两个补充加氢精制反应阶段，并采用催化剂A和催化剂B的级配方案，第一补充加氢精制反应阶段使用催化剂A，第二补充加氢精制反应阶段使用催化剂B，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表5所示。
对比例3
参照实施例3，所不同的是，补充加氢精制反应区内采用的催化剂级配为催化剂C和催化剂D，第一补充加氢精制反应阶段使用催化剂C，第二补充加氢精制反应阶段使用催化剂D，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表5所示。
实施例4
参照实施例3，所不同的是，改变补充加氢精制催化剂的级配顺序，第一补充加氢精制反应阶段使用催化剂B，第二补充加氢精制反应阶段使用催化剂A，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表5所示。
表5
Figure PCTCN2017113834-appb-000001
[1]：以160-340℃馏分为基准。
从表5中所列数据可知，对原料加氢的现有技术方案来说，不同条件下得到十六氢芘产品的纯度、收率不同，实施例3的催化剂级配方法最优。
从实施例1-4的数据可以看出，本发明中煤焦油原料经过预处理、加氢精制、加氢裂化、补充加氢精制的工艺方法，均能够得到纯度较高的十六氢芘产品。而其中，补充精制反应区采用催化剂级配的方案，所得十六氢芘的收率和纯度更高，具有更理想的加氢效果。
在以下实施例5-8和对比例4-6中，加氢精制反应和加氢裂化反应的操作条件如下表6所示。
表6
  加氢精制 加氢裂化
催化剂 FF-36 FC-26
氢分压，MPa 15.0 15.0
液时空速，h-1 0.6 0.6
氢油体积比，v/v 2000 2000
反应温度，℃ 335 365
氮含量，μg/g 11 3
在以下实施例5-8和对比例4-6中，从加氢裂化反应流出物中分离出的重馏分的初馏点为160℃，加氢裂化反应所得产物的产品分布包括：<160℃馏分8.4wt％，160℃以上馏分91.6wt％，且160℃以上馏分中芘含量为1.2wt％；从补充加氢精制反应流出物中分离出的富含十六氢芘液体馏分的馏程为250～340℃。
实施例5
如图2所示流程，补充加氢精制反应区内催化剂不分段，而仅采用一种催化剂A，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表7所示。
对比例4
参照实施例5，所不同的是，补充加氢精制反应区内采用的催化剂为催化剂C，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表7所示。
实施例6
如图2所示流程，补充加氢精制反应区内催化剂不分段，而仅采用一种催化剂B，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表7所示。
对比例5
参照实施例6，所不同的是，补充加氢精制反应区内采用的催化剂为催化剂D，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表7所示。
实施例7
如图2所示流程，补充加氢精制反应区分为两个补充加氢精制反应阶段，并采用催化剂A和催化剂B的级配方案，第一补充加氢精制反应阶段使用催化剂A，第二补充加氢精制反应阶段使用催化剂B，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表7所示。
对比例6
参照实施例7，所不同的是，补充加氢精制反应区内采用的催化剂级配为催化剂C和催化剂D，第一补充加氢精制反应阶段使用催化剂C，第二补充加氢精制反应阶段使用催化剂D，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表7所示。
实施例8
参照实施例7，所不同的是，改变补充加氢精制催化剂的级配顺序，第一补充加氢精制反应阶段使用催化剂B，第二补充加氢精制反应阶段使用催化剂A，其中，补充加氢精制反应区的操作条件和所得十六氢芘的收率和纯度如表7所示。
表7
Figure PCTCN2017113834-appb-000002
Figure PCTCN2017113834-appb-000003
[2]：以>160℃馏分为基准。
从表7中所列数据可知，对原料加氢的现有技术方案来说，不同条件下得到十六氢芘产品的纯度、收率不同，实施例7的催化剂级配方法最优。
从实施例5-8的数据可以看出，本发明中煤焦油原料经过预处理、加氢精制、加氢裂化、补充加氢精制的工艺方法，均能够得到纯度较高的十六氢芘产品。而其中，补充精制反应区采用催化剂级配的方案，所得十六氢芘的收率和纯度更高，具有更理想的加氢效果。
以上详细描述了本发明的优选实施方式，但是，本发明并不限于此。在本发明的技术构思范围内，可以对本发明的技术方案进行多种简单变型，包括各个技术特征以任何其它的合适方式进行组合，这些简单变型和组合同样应当视为本发明所公开的内容，均属于本发明的保护范围。

Claims (22)

  1. 一种十六氢芘的制备方法，该方法包括：在加氢催化剂的存在下，将含有芘系化合物的烃油原料进行加氢反应，其中，所述芘系化合物选自芘及其不饱和加氢产物中的至少一种，所述加氢催化剂含有载体和负载于载体上的活性金属组分，所述活性金属组分为Pt和/或Pd，所述载体含有小晶粒Y型分子筛、氧化铝和无定形硅铝，
    所述小晶粒Y型分子筛的晶粒平均直径为200-700nm，SiO2/Al2O3摩尔比为40-120，相对结晶度≥95％，比表面积为900-1200m2/g，1.7-10nm的二次孔的孔容占总孔容的50％以上。
  2. 根据权利要求1所述的方法，其中，所述小晶粒Y型分子筛的晶粒平均直径为300-500nm，相对结晶度为95-120％，1.7-10nm的二次孔的孔容占总孔容的50-80％。
  3. 根据权利要求1或2所述的方法，其中，所述小晶粒Y型分子筛的晶胞常数为2.425-2.435nm，孔容为0.5-0.8mL/g。
  4. 根据权利要求1-3中任意一项所述的方法，其中，在所述加氢催化剂中，以所述加氢催化剂的总重量为基准，所述活性金属组分的含量为0.1-2重量％，所述载体的含量为98-99.9重量％。
  5. 根据权利要求1-4中任意一项所述的方法，其中，在所述载体中，以所述载体的总重量为基准，所述小晶粒Y型分子筛的含量为5-40重量％，所述氧化铝的含量为10-40重量％，所述无定形硅铝的含量为20-65重量％。
  6. 根据权利要求1-5中任意一项所述的方法，其中，所述加氢反应的过程包括依次进行的两个反应阶段，第一反应阶段使用的加氢催化剂中的活性金属组分的含量百分比x1低于第二反应阶段使用的加氢催化剂中的活性金属组分的含量百分比x2，第一反应阶段使用的加氢催化剂中的小晶粒Y型分子筛的含量百分比y1高于第二反应阶段使用的加氢催化剂中的小晶粒Y型分子筛的含量百分比y2。
  7. 根据权利要求6所述的方法，其中，x1比x2低0.1-1.5个百分点，y1比y2高5-35 个百分点。
  8. 根据权利要求7所述的方法，其中，x1比x2低0.3-1.5个百分点，y1比y2高10-35个百分点。
  9. 根据权利要求1-8中任意一项所述的方法，其中，所述加氢反应的工艺条件包括：氢分压为4-20MPa，液时体积空速为0.05-6h-1，氢油体积比为50-3000，平均反应温度为150-380℃。
  10. 根据权利要求6-8中任意一项所述的方法，其中，所述第二反应阶段的平均反应温度比所述第一反应阶段的平均反应温度低10-150℃。
  11. 根据权利要求10所述的方法，其中，所述第一反应阶段的平均反应温度为180-380℃，所述第二反应阶段的平均反应温度为150-350℃。
  12. 根据权利要求1-11中任意一项所述的方法，其中，所述含有芘系化合物的烃油原料中的芘系化合物的含量为0.5重量％以上。
  13. 根据权利要求1或12所述的方法，其中，所述含有芘系化合物的烃油原料是初馏点为130-220℃且终馏点为300-400℃的重馏分油。
  14. 根据权利要求13所述的方法，其中，所述含有芘系化合物的烃油原料为柴油馏分，且初馏点为160-200℃，终馏点为300-350℃。
  15. 根据权利要求1-14中任意一项所述的方法，其中，所述方法还包括按照以下步骤制备所述含有芘系化合物的烃油原料：
    (1)将煤焦油原料进行预处理；
    (2)将经过预处理的煤焦油原料进行加氢精制反应；
    (3)将所得反应流出物可选地进行分离，然后进行加氢裂化反应；
    (4)将加氢裂化反应所得反应流出物进行气液分离，然后对分离出的液相进行分馏，将分馏出的重馏分作为所述含有芘系化合物的烃油原料。
  16. 根据权利要求15所述的方法，其中，所述煤焦油原料的芳烃含量为20-100重量％，20℃密度为1.023-1.235g/cm3，馏程为200-700℃。
  17. 根据权利要求15或16所述的方法，其中，所述煤焦油原料为高温煤焦油或者高温煤焦油提取出蒽、菲、咔唑和荧蒽中的至少一种后的剩余馏分。
  18. 根据权利要求15-17中任意一项所述的方法，其中，步骤(1)所述的预处理包括机械脱杂质、脱水和电脱盐操作。
  19. 根据权利要求15-18中任意一项所述的方法，其中，所述加氢精制反应的工艺条件包括：氢分压为3-19MPa，平均反应温度为260-440℃，液时体积空速为0.1-4h-1，氢油体积比为300：1至3000：1。
  20. 根据权利要求15-19中任意一项所述的方法，其中，所述加氢裂化反应的工艺条件包括：氢分压为3-19MPa，平均反应温度为260-440℃，液时体积空速为0.3-4h-1，氢油体积比为300：1至5000：1。
  21. 根据权利要求15-20中任意一项所述的方法，其中，所述方法还包括将加氢反应所得反应流出物进行分离和分馏，得到富含十六氢芘组分和重组分，将所述重组分的至少部分循环回步骤(3)中进行加氢裂化反应。
  22. 根据权利要求1-20中任意一项所述的方法，其中，所述方法还包括将加氢反应所得反应流出物进行分离和分馏，得到富含十六氢芘组分和重组分，将所述富含十六氢芘组分进行冷却降温，然后经过滤、抽提得到固体的十六氢芘。
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