WO2009067885A1 - Système et procédé de fabrication d'essence de qualité élevée par recombinaison catalytique d'hydrocarbures - Google Patents

Système et procédé de fabrication d'essence de qualité élevée par recombinaison catalytique d'hydrocarbures Download PDF

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Publication number
WO2009067885A1
WO2009067885A1 PCT/CN2008/072943 CN2008072943W WO2009067885A1 WO 2009067885 A1 WO2009067885 A1 WO 2009067885A1 CN 2008072943 W CN2008072943 W CN 2008072943W WO 2009067885 A1 WO2009067885 A1 WO 2009067885A1
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Prior art keywords
gasoline
hydrogenation
light
hydrogenation unit
heavy
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PCT/CN2008/072943
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English (en)
French (fr)
Inventor
Ranfeng Ding
Original Assignee
Ranfeng Ding
Priority date (The priority date is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the date listed.)
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Publication date
Priority claimed from CN 200710169942 external-priority patent/CN101429442B/zh
Priority claimed from CN2007101699475A external-priority patent/CN101429447B/zh
Priority claimed from CN200710169949A external-priority patent/CN101429449B/zh
Application filed by Ranfeng Ding filed Critical Ranfeng Ding
Priority to JP2010532411A priority Critical patent/JP2011503264A/ja
Priority to EP08855052A priority patent/EP2236583A4/en
Priority to US12/682,034 priority patent/US8524043B2/en
Priority to EA201070499A priority patent/EA201070499A1/ru
Priority to CA2705034A priority patent/CA2705034C/en
Publication of WO2009067885A1 publication Critical patent/WO2009067885A1/zh

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Classifications

    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G21/00—Refining of hydrocarbon oils, in the absence of hydrogen, by extraction with selective solvents
    • 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
    • C10G55/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one refining process and at least one cracking process
    • C10G55/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one refining process and at least one cracking process plural serial stages only
    • C10G55/06—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one refining process and at least one cracking process plural serial stages only including at least one catalytic cracking step
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G67/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G67/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only
    • C10G67/16—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only plural parallel stages only
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
    • 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
    • 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
    • C10G2400/30—Aromatics

Definitions

  • the present invention relates to a system for preparing high quality gasoline and a method thereof, and more particularly to a system and method for preparing a high quality gasoline by catalytic hydrocarbon recombination.
  • Catalytic cracking, catalytic cracking and heavy oil catalytic cracking technology are the core technologies of refining.
  • Catalytic cracking is divided into wax oil catalytic cracking and heavy oil catalytic cracking.
  • the oil produced from these processes is called catalytic hydrocarbons, and the obtained catalytic hydrocarbons are processed.
  • Fractionation tower fractionation can be divided into dry steam, liquefied petroleum gas, gasoline, diesel, heavy oil and other products, of which gasoline and diesel account for more than 70% of the total gasoline and diesel supply on the market.
  • the Chinese invention patent of the "catalytic hydrocarbon recombination treatment method" of No. 03148181.7 provides a catalytic hydrocarbon recombination treatment method
  • the Chinese invention patents with the patent numbers 200310103541.9 and 200310103540.4 disclose the improved patent, which relates to Water washing systems and solvent recovery, but none of these published patents address the issue of how to reduce sulfur and reduce olefins.
  • the current GB17930 gasoline standard requires a sulfur content of not more than 0.05% (wt), an olefin content of not more than 35% (v), and a benzene content of not more than 2.5% (v), and most refineries can guarantee the quality of gasoline.
  • the national III gasoline standard to be implemented in 2010 requires: sulfur content not greater than 0.015% (wt), olefin content not greater than 30% (v), benzene content not greater than 1% (v).
  • higher national IV gasoline standards must be met: sulfur content not greater than 0.005% (wt), olefins not greater than 25% (V) or lower.
  • the gasoline quality solution must consider the transition from the national III gasoline standard to the national IV gasoline standard.
  • the better planning plan should be One-time follow the national IV gasoline standard planning program.
  • catalytic gasoline Since the proportion of various blending components in China's gasoline products is very different from that in developed countries, catalytic cracking gasoline (hereinafter referred to as catalytic gasoline) occupies a high proportion, and the proportion of reformed gasoline and alkylated gasoline is small, and this The situation will exist for a long time. Therefore, the problem of sulfur reduction and olefin reduction to be solved by upgrading the quality of gasoline mainly involves the problem of catalytic gasoline.
  • catalytic cracking feedstock hydrodesulfurization is not likely to be applied on a large scale due to large investment, high operating cost, limited existing refinery conditions, and low processing Sulphur crude oil refineries are not suitable.
  • excessive reduction of olefins in catalytic cracking units can also exacerbate the loss of light products and gasoline octane number (RON).
  • One object of the present invention is to provide a catalytic hydrocarbon recombination system for producing gasoline having a low sulfur content, a low olefin content, and an increased octane number (RON) at a low cost.
  • the present invention adopts the following technical solutions:
  • a system for catalyzing the recombination of hydrocarbons to produce high quality gasoline comprising a distillation device; characterized in that: the upper portion of the distillation device is connected to a light gasoline hydrogenation device through a light gasoline line; the lower portion of the distillation device is passed through a heavy gasoline pipeline and extracted The system is connected; the upper part of the extraction system directly extracts the product through the pipeline, and the lower part of the extraction system is connected to the light gasoline pipeline after the light gasoline hydrogenation unit through the pipeline.
  • a preferred technical solution is characterized in that: the lower part of the distillation apparatus is first connected to a heavy gasoline hydrogenation unit through a heavy gasoline line; and the heavy gasoline hydrogenation unit is connected to the extraction system through a pipeline.
  • a preferred technical solution is characterized in that: the lower part of the extraction system is first connected to an aromatic hydrocarbon hydrogenation unit through a pipeline, and the aromatic hydrocarbon hydrogenation unit is connected to a light gasoline pipeline after a light gasoline hydrogenation unit through a pipeline.
  • Another object of the present invention is to provide a process for the above-described catalytic hydrocarbon recombination to produce high quality gasoline.
  • a method for preparing high-quality gasoline by catalytic hydrocarbon recombination the steps are as follows: adding stable gasoline to a distillation device for fractionation, cutting and fractionating light gasoline and heavy gasoline; and the light gasoline is introduced into the light gasoline hydrogenation device through the upper portion of the distillation device Hydrogen; the heavy gasoline is subjected to extraction and separation in an extraction system to separate aromatic hydrocarbons and raffinate oil; the aromatic hydrocarbons are used by blending with a hydrotreated light gasoline through a pipeline, and the raffinate oil is directly produced as a chemical light oil.
  • a preferred embodiment is characterized in that the heavy gasoline is hydrotreated in a heavy gasoline hydrogenation unit prior to being subjected to extraction separation in the extraction system.
  • a preferred embodiment is characterized in that: the aromatic hydrocarbon is first subjected to hydrotreatment and then used in combination with a hydrotreated light gasoline via a pipeline.
  • the bottom pressure is 0. 12 ⁇ 0.
  • 30MPa absolute
  • the distillation range of the light gasoline is controlled at 30 ° C ⁇ 100 ° C
  • the distillation range of the heavy gasoline is controlled at 100 ° C ⁇ 205 °C.
  • the singularity of the top of the distillation apparatus is 0. 11 ⁇ 0. 28MPa (absolute), The bottom pressure is 0. 12 ⁇ 0. 30MPa (absolute); the distillation range of the light gasoline is controlled at 30 ° C ⁇ 100 ° C; the distillation range of the heavy gasoline is controlled at 100 ° C ⁇ 205 ° C.
  • a preferred embodiment is characterized in that: the catalyst in the light gasoline hydrogenation unit is a selective hydrogenation catalyst GHT-20; the light gasoline hydrogenation unit has a volumetric space velocity ratio of 2-4; hydrogen/oil volume 6 ⁇ 1. 6MPa ( ⁇ ).
  • the operating temperature is 240 ⁇ 260 ° C, the operating pressure is 1. 4 ⁇ 1. 6MPa (absolute).
  • a preferred embodiment is characterized in that the physicochemical properties of the selective hydrogenation catalyst GHT-20 in the light gasoline hydrogenation unit are as follows: Strength N/cm 170 Bulk density g/ml 0. 70
  • the catalyst in the heavy gasoline hydrogenation unit is a total hydrogenation catalyst GHT-22; the volume ratio of the heavy gasoline hydrogenation unit is 2 to 4 The hydrogen/oil volume ratio is 250 to 350; the operating temperature is 290 to 330 ° C, and the operating pressure is 1.2 to 3 MPa (absolute).
  • a preferred embodiment is characterized in that the physical and chemical properties of all the hydrogenation catalysts GHT-22 in the heavy gasoline hydrogenation unit are shown in the following table.
  • a preferred embodiment is characterized in that: the catalyst in the aromatic hydrocarbon hydrogenation unit is the entire hydrogenation catalyst GHT-22; the volume ratio of the aromatic hydrocarbon hydrogenation unit is 2 to 3; the hydrogen/oil volume ratio is 250. ⁇ 300; Operating temperature is 285 ⁇ 325 °C, operating pressure is 1.5 ⁇ 2.5MPa (absolute).
  • a preferred embodiment is characterized in that the physicochemical properties of all the hydrogenation catalysts GHT-22 in the aromatic hydrocarbon hydrogenation unit are shown in the following table.
  • the distillation apparatus used in the present invention is a distillation system disclosed in the Chinese invention patent of the "catalytic hydrocarbon recombination treatment method" of Patent No. 03148181.7.
  • the extraction system uses an extraction system disclosed in Patent Nos. 200310103541.9 and 2003 10103540.4, including solvent recovery and water washing systems.
  • the hydrogenation unit used in the present invention is an existing hydrogenation unit including a heating furnace, a heat exchanger, a high pressure separator, an air condenser, a water condenser and the like.
  • Embodiment 1 is a schematic flow chart of Embodiment 1 of the present invention.
  • Figure 2 is a flow chart showing Embodiments 2 and 3 of the present invention.
  • Figure 3 is a flow chart showing the fourth and fifth embodiments of the present invention.
  • FIG. 1 it is a schematic flowchart of an embodiment of the present invention.
  • the distillation range is 30-205 ° C, the sulfur content is 85 ppm, the mercaptan content is 15 ppm, the olefin content is 25% (v), the diolefin content is 0.1% (v), and the aromatic content is 13% (v). ), a stabilized gasoline (catalytic gasoline) having an octane number (RON) of 87 and a density of 728 kg/m 3 is cut and fractionated in a distillation column 1 at a flow rate of 100,000 tons/year, and the temperature at the top of the distillation column 1 is 105°C, the bottom temperature is 216 ° C, the top pressure is 0.
  • RON octane number
  • the bottom pressure is 0. 25MPa (absolute), respectively, light gasoline and heavy gasoline, the light gasoline (the distillation range is 30-10CTC) is distilled out through the upper part of the distillation column 1, and the total amount of steam is 50,000 tons/year, and then enters the light gasoline hydrogenation unit 3 for hydrotreating; the catalyst in the light gasoline hydrogenation unit 3 is selected Hydrogenation catalyst GHT-20, the volume of the light gasoline hydrogenation unit 3 The airspeed ratio is 4; the hydrogen/oil volume ratio is 300; the operating temperature is 250 ° C, the operating pressure is 1.5 MPa (absolute) (selective hydrogenation); the heavy gasoline (distillation range is 100-205 ° C) The 50,000 ton/year flow rate is introduced into the heavy gasoline extraction system 2 for extraction and separation, and the aromatic hydrocarbon and the raffinate oil are separated; the solvent used in the extraction system 2 is sulfolane, the extraction temperature is 120 ° C, and the solvent ratio (solvent) /feed) is
  • the obtained blended gasoline has a distillation range of 30-205 ° C, a sulfur content of 102.8 ppm, and a mercaptan content of
  • the obtained chemical light oil has a distillation range of 100-205 ° C, a sulfur content of 29.0 ppm, a mercaptan content of 1.0 ppm, an olefin content of 28.2% (v), a diene content of less than 0.01% (v), and an aromatic content. It is 3.0% (v), has an octane number (RON) of 78.5, a density of 775.5 kg/ m3 , and a recovery of 39,000 tons/year.
  • RON octane number
  • the measurement method used in the present invention is:
  • Sulfur content SH/T0689-2000 Determination of total sulfur content of light hydrocarbons and engine fuels and other oils (UV fluorescence method);
  • Mercaptan sulfur GB/T1792-1988 Determination of mercaptan sulfur in distillate fuel oil (potentiometric titration); 4, olefins: GB/T11132-2002 liquid petroleum products hydrocarbons determination method (fluorescent indicator adsorption method);
  • aromatic hydrocarbons GB/T11132-2002 liquid petroleum products hydrocarbons determination method (fluorescent indicator adsorption method);
  • Density GB/T1884-2000 laboratory method for determination of density of crude oil and liquid petroleum products (density meter method);
  • FIG. 2 it is a schematic flowchart of the embodiment.
  • the distillation range is 30-205 ° C, the sulfur content is 100 ppm, the mercaptan content is 5 ppm, the olefin content is 30% (v), the diolefin content is 0.1% (v), and the aromatic content is 15% (v). ), octane number (RON) (RON) 89, low sulfur content stabilized gasoline (catalytic gasoline) with a density of 728 kg/ m3 , cutting fractionation in distillation column 1 at a flow rate of 100,000 tons/year, distillation column The top temperature of the column is 82 ° C, the temperature at the bottom of the column is 186 ° C, the pressure at the top of the column is 0.
  • the pressure at the bottom of the column is 0. 25 MPa (absolute), respectively, light gasoline and heavy gasoline are obtained, Light gasoline (distillation range of 30-80 ° C) is distilled off through the upper part of the distillation column 1, the total amount of steaming is 40,000 tons / year, and then into the light gasoline hydrogenation unit 3-1 hydrotreating;
  • the catalyst of the gasoline hydrogenation unit 3-1 is a selective hydrogenation catalyst GHT-20; the light gasoline hydrogenation unit 3-1 has a volumetric space velocity ratio of 2; a hydrogen/oil volume ratio of 150; and an operating temperature of 220°. C, Operating pressure is 0.
  • the heavy gasoline (the distillation range is 80-205 ° C) enters the heavy gasoline hydrogenation unit 3-2 hydrotreating at a flow rate of 60,000 tons/year; the heavy gasoline hydrogenation unit 3-2
  • the catalyst in the catalyst is the total hydrogenation catalyst GHT-22; the volumetric space velocity ratio of the heavy gasoline hydrogenation unit 3-2 is 2; the hydrogen/oil volume ratio is 250; the operating temperature is 290 ° C, and the operating pressure is 1.2 MPa.
  • the solvent used in the extraction system 2 is N-methylpyrrolidone
  • the extraction temperature is 115 ° C
  • the solvent The ratio (solvent/feed) is 3.5 (mass)
  • the raffinate wash ratio is 0.2 (mass)
  • the solvent recovery temperature is 151 ° C
  • the solvent recovery pressure is 0.112 MPa (absolute);
  • the aromatics are 15,000 tons/
  • the annual flow rate is blended with the hydrogenated light gasoline, which is produced as a high quality ethylene material at a flow rate of 45,000 tons/year.
  • the obtained blended gasoline has a distillation range of 30-205 ° C, a sulfur content of 5.27 ppm, a mercaptan content of less than 1 ppm, an olefin content of 17.8% (v), a diene content of less than 0.01% (v), and an aromatic content of 25.6% (v), octane number (RON) of 94.1, density of 703.8 kg / m3 , and production volume of 55,000 tons / year.
  • the obtained high-quality vinyl material has a distillation range of 80-205 ° C, a sulfur content of 2.0 ppm, a mercaptan content of less than 1 ppm, an olefin content of less than 0.1% (v), a diene content of less than 0.01% (v), an aromatic hydrocarbon.
  • the content is 3.0% (v)
  • the octane number (RON) is 81.0
  • the density is 760.0 kg/ m3
  • the amount of production is 45,000 tons/year.
  • the physicochemical properties of the selective hydrogenation catalyst GHT-20 are shown in the following table.
  • Carrier m% 82. 4 The measurement method used in the present invention is (the same below):
  • Sulfur content SH/T0689-2000 Determination of total sulfur content of light hydrocarbons and engine fuels and other oils (UV fluorescence method);
  • olefins GB/T11132-2002 liquid petroleum products hydrocarbons determination method (fluorescent indicator adsorption method);
  • aromatic hydrocarbons GB/T11132-2002 liquid petroleum products hydrocarbons determination method (fluorescent indicator adsorption method);
  • Density GB/T1884-2000 laboratory method for determination of density of crude oil and liquid petroleum products (density meter method);
  • FIG. 2 it is a schematic flowchart of the embodiment.
  • the distillation range is 30-205 ° C, the sulfur content is 2000 ppm, the mercaptan content is 50 ppm, the olefin content is 40% (v), the diolefin content is 1.0% (v), and the aromatic content is 19% (v).
  • High sulphur content stabilized gasoline catalytic gasoline having an octane number (RON) (RON) of 91 and a density of 728 kg/ m3 .
  • the distillation range is 30-90 ° C) is distilled off through the upper part of the distillation column 1, the total amount of steaming is 43,000 tons / year, and then enters the light gasoline hydrogenation unit 3-1 hydrotreating; the light gasoline hydrogenation
  • the catalyst of the device 3-1 is a selective hydrogenation catalyst GHT-20; the volumetric space velocity ratio of the light gasoline hydrogenation unit 3-1 is 4; the hydrogen/oil volume ratio is 300; the operating temperature is 280 ° C, operation
  • the pressure is 2.
  • the heavy gasoline (distillation range is 90-205) enters the heavy gasoline hydrogenation unit 3-2 hydrotreating at a flow rate of 57,000 tons/year; the heavy gasoline hydrogenation unit 3 -2
  • the chemical agent is a selective hydrogenation catalyst GHT-20; the heavy gasoline hydrogenation unit 3-2 has a volumetric space velocity ratio of 4; a hydrogen/oil volume ratio of 350; an operating temperature of 330 ° C, and an operating pressure of 3.
  • OMPa (absolute) then enters the heavy gasoline extraction system 2 through the pipeline for separation and separation, and separates the aromatic hydrocarbon and the raffinate oil; the solvent used in the extraction system 2 is N-methylpyrrolidone, and the extraction temperature is 115 ° C.
  • the solvent ratio (solvent/feed) is 3.5 (mass), the raffinate wash ratio is 0.2 (mass), the solvent recovery temperature is 151 ° C, the solvent recovery pressure is 0.112 MPa (absolute); the aromatic hydrocarbon is 19,000 tons
  • the flow rate per year is adjusted with the hydrogenated light gasoline, and the raffinate oil is produced as a high-quality ethylene material at a flow rate of 38,000 tons/year.
  • the resulting blended gasoline has a distillation range of 30-205 ° C, a sulfur content of 7.52 ppm, a mercaptan content of less than 1 ppm, an olefin content of 17.99% (v), a diene content of less than 0.01% (v), and an aromatic content of 29.1% (v), octane number (RON) of 95.2, density of 720.1 kg/ m3 , yield It is 62,000 tons/year.
  • the olefin content is less than 1 ppm, the olefin content is 6% (v), and the diene content is less than 0.01% (v).
  • the aromatics content is 3.0 % (v)
  • the octane number (RON) is 81. 5
  • the density is 740. 0 kg / m3
  • the amount of production is 38,800 tons / year.
  • the distillation range is 30-205 ° C, the sulfur content is lOOppm, the mercaptan content is 5 ppm, the olefin content is 30% (v), the diolefin content is 0.1% (v), and the aromatic content is 15% (v). ), a low sulfur content stabilized gasoline (catalytic gasoline) having an octane number (RON) of 89 and a density of 728 kg/ m3 .
  • the fractionation in the distillation column 1 is carried out at a flow rate of 100,000 tons/year, and the column of the distillation column 1
  • the top temperature is 86 ° C, the bottom temperature is 192 ° C, the top pressure is 0.
  • the bottom pressure is 0. 25MPa (absolute), fractional distillation to obtain light gasoline and heavy gasoline, respectively, the light gasoline (the distillation range is 30-80 ° C) is steamed in the upper part of the distillation column 1 and then hydrotreated in the light gasoline hydrogenation unit 3-1, and the total steaming amount of the light gasoline after hydrogenation is 40,000 tons/year;
  • the catalyst in the light gasoline hydrogenation unit 3-1 is a selective hydrogenation catalyst GHT-20; the light gasoline hydrogenation unit 3-1 has a volumetric space velocity ratio of 2; a hydrogen/oil volume ratio of 150;
  • the temperature is 230 ° C, the operating pressure is 1.
  • the heavy gasoline (the distillation range is 80-205 ° C) with a flow rate of 60,000 tons / year through the pipeline into the heavy gasoline extraction system 2 extraction and separation , separation of aromatics
  • the solvent used in the extraction system 2 is N-formylmorpholine, the extraction temperature is 115 ° C, the solvent ratio (solvent / feed) is 3.5 (mass), and the raffinate oil wash ratio is 0. 2 (mass), the solvent recovery temperature is 151 V, the solvent recovery pressure is 0. 112 MPa (absolute); the aromatic hydrocarbon enters the aromatic hydrocarbon hydrogenation unit 3-2 hydrogenation at a flow rate of 15,000 tons/year.
  • the catalyst in the aromatic hydrocarbon hydrogenation unit 3-2 is the entire hydrogenation catalyst GHT-22; the aromatic hydrocarbon hydrogenation unit 3-2 has a volumetric space velocity ratio of 2; the hydrogen/oil volume ratio is 250; and the operating temperature is 295°. C, the operating pressure is 2. OMPa (absolute); and then blended with the hydrogenated light gasoline, the raffinate oil is used as a chemical light oil at a flow rate of 15,000 tons / year.
  • the diene content is less than 0.01% (v)
  • the olefin content is less than 0.01% (v)
  • the diene content is less than 0.01% (v)
  • the aromatics content is 25.6% (v)
  • the octane number (RON) is 94.1
  • the density is 703. 8 kg/ m3
  • the recovery is 55,000 tons/year.
  • the thiol content is 0. 0ppm
  • the mercaptan content is O.
  • the density of the olefin is 0. 01% (v)
  • the aromatic content is 3.0% (v)
  • the octane number (RON) is 81. 0, the density. It is 760. 0 kg/ m3 and the production volume is 45,000 tons/year.
  • the physicochemical properties of the selective hydrogenation catalyst GHT-20 are shown in the following table.
  • FIG. 3 it is a schematic flowchart of this embodiment.
  • the distillation range is 30-205 ° C, the sulfur content is 2000 ppm, the mercaptan content is 50 ppm, the olefin content is 40% (v), the diene content is 1.0% (v), and the aromatic content is 19% (v). ), high sulphur content stabilized gasoline (catalytic gasoline) having an octane number (RON) of 91 and a density of 728 kg/ m3 .
  • Cutting fractionation in the distillation column 1 at a flow rate of 100,000 tons/year, the tower of the distillation column 1 The top temperature is 86 ° C, the bottom temperature is 192 ° C, the top pressure is 0.
  • the bottom pressure is 0. 25MPa (absolute), fractional distillation to obtain light gasoline and heavy gasoline, respectively, the light gasoline (the distillation range is 30-90 ° C) is steamed in the upper part of the distillation column 1 and then hydrotreated in the light gasoline hydrogenation unit 3-1, and the total steaming amount of the light gasoline after hydrogenation is 43,000 tons/year;
  • the catalyst in the light gasoline hydrogenation unit 3-1 is a selective hydrogenation catalyst GHT-20; the light gasoline hydrogenation unit 3-1 has a volumetric space velocity ratio of 4; a hydrogen/oil volume ratio of 300;
  • the temperature is 250 ° C, the operating pressure is 1.
  • the heavy gasoline (; distillation range is 90-205 ° C) with a flow rate of 570,000 tons / year through the pipeline into the heavy gasoline extraction system 2
  • extraction separation The aromatic solvent and the raffinate oil; the solvent used in the extraction system 2 is N-formylmorpholine, the extraction temperature is 115 ° C, the solvent ratio (solvent / feed) is 3.5 (mass), the raffinate oil wash The ratio is 0.2 (mass), the solvent recovery temperature is 151 ° C, the solvent recovery pressure is 0.11 MPa (absolute); the aromatic hydrocarbon enters the aromatic hydrocarbon hydrogenation unit 3-2 at a flow rate of 15,000 tons/year.
  • the catalyst in the aromatic hydrocarbon hydrogenation unit 3-2 is all 5MPa ( ⁇ ); then the operating temperature is 325 ° C, the operating pressure is 2. 5MPa (absolute); and then the volume of the air-to-air ratio is 3; Then, it is blended with the hydrogenated light gasoline, and the raffinate oil is produced as a chemical light oil at a flow rate of 15,000 tons/year.
  • the olefin content is less than 0. 01% (v), the olefin content is less than 0.01% (v), the diene content is less than 0.01% (v)
  • the aromatics content is 28.2% (v), the octane number (RON) is 94. 05, the density is 721. 4 kg / m3 , and the amount of extraction is 640,000 tons/year.
  • the olefin content is 0. 01% (v), the olefin content is 0. 01% (v), the olefin content is 0. 01% (%) v), an aromatic content of 3.0 % (v), an octane number (RON) of 82.0, a density of 740. 0 kg / m3 , a production of 36,000 tons / year.
  • the amount of catalyst used is greatly reduced by hydrotreating specifically for light gasoline, and/or heavy gasoline, and/or aromatics; Gasoline, and / or heavy gasoline, and / or aromatics, the amount of catalyst used is much smaller, in addition, can reduce the content of monoolefins and diene, not only reduce the content of mercaptans, but also reduce the total sulfur content; , Selective hydrogenation, using specific catalysts and parameters, mainly desulfurization, to solve the problem of olefins and diene, the effect is obvious.

Description

一种催化烃重组制备高质量汽油的系统及其方法 技术领域
本发明涉及一种制备高质量汽油的系统及其方法， 特别涉及一种 催化烃重组制备高质量汽油的系统及其方法。
背景技术
催化裂化、催化裂解及重油催化裂解技术是炼油的核心技术， 催化裂 化分为蜡油催化裂化、重油催化裂化； 从这些工艺生产的生成油统称为催 化烃， 所得催化烃经过加工处理， 一般是分馏塔分馏， 可以分馏出干汽、 液化汽、 汽油、 柴油、 重油等产品， 其中汽油、 柴油占据市场上汽油、 柴 油供应总量的 70%以上。
随着环保要求的越来越严格， 汽油、柴油的标准不断提高， 现有的催 化烃经过分馏塔分馏的加工处理方法显出以下不足：一个是该处理方法所 生产的汽油和柴油的质量有待提高：汽油的烯烃含量偏高，辛烷值（RON) 偏低， 柴油的十六烷值偏低， 安定性不符合要求； 二是上述处理方法不能 同时生产多种标号的汽油， 而且产品品种单一； 三是所生产的柴油、汽油 的比例与市场的需求不匹配， 柴油不能满足需求， 而汽油供大于求。
为了解决上述问题， 专利号为 03148181.7的 "催化烃重组处理方法" 的中国发明专利提供了一种催化烃重组处理方法， 并且专利号分别为 200310103541.9和 200310103540.4的中国发明专利公开了其改进专利， 涉及水洗系统及溶剂回收，但这些公开的专利中均未涉及如何降硫和降烯 烃的问题。
目前的 GB17930汽油标准要求硫含量不大于 0. 05% (wt), 烯烃含量 不大于 35% (v)、 苯含量不大于 2. 5% (v)， 绝大部分炼油厂可以保证汽油 质量。 但是， 即将于 2010年实施的国家 III汽油标准要求： 硫含量不大 于 0. 015% (wt), 烯烃含量不大于 30% (v)， 苯含量不大于 1% (v)。 对大 多数炼油厂而言， 必须面对更高的国家 IV汽油标准要求： 硫含量不大于 0. 005% (wt ), 烯烃不大于 25% (V) 或更低。 汽油质量解决方案必须考虑 从国家 III汽油标准到国家 IV汽油标准的过渡， 较好的规划方案应该是 一次性按照国家 IV汽油标准规划方案。
由于我国汽油产品中各调和组分的比例与发达国家差别很大，催化裂 化汽油 （以下简称催化汽油） 占有很高的比例， 重整汽油、 烷基化汽油所 占比例较小， 而且， 这种状况将长期存在。 因此， 汽油质量升级所要解决 的降硫和降烯烃的问题主要涉及催化汽油的问题。
一般认为， 催化裂化原料中总硫的 5-10%将进入汽油馏分， 根据我国 炼油厂催化原料加氢精制能力很小、二次加工催化裂化能力较大并有渣油 焦化的特点， 加工低硫 （含硫 0. 3% ) 原油的炼油厂催化汽油硫含量约 200ppm， 加工含硫 0. 8%的原油， 催化汽油中硫含量约 900ppm， 因此， 汽 油质量升级的难点从烯烃转变为硫的问题。催化裂化工艺或催化剂的改进 不可能从根本上解决硫的问题， 催化裂化原料加氢脱硫由于投资大、运行 费用高、现有炼油厂条件有限而不可能大规模应用， 而且对于加工较低含 硫原油的炼油厂并不适用， 同时， 催化裂化装置过度降低烯烃还会加剧轻 质产品及汽油辛烷值 （RON) 的损失。
因此，提供一种低成本制备低硫含量、低烯烃含量并且辛烷值（RON) 高的调和汽油的处理系统及其方法就成为该技术领域急需解决的技术难 题。
发明内容
本发明的目的之一是提供一种低成本制备低硫含量、低烯烃含量并且 提高辛烷值 （RON) 的汽油的催化烃重组系统。
为实现上述目的， 本发明采取以下技术方案：
方案之一：
一种催化烃重组制备高质量汽油的系统，包括蒸馏装置；其特征在于: 所述蒸馏装置上部通过轻汽油管线与轻汽油加氢装置相连接；所述蒸馏装 置下部通过重汽油管线与抽提系统相连接；所述抽提系统上部通过管线直 接采出产品，所述抽提系统下部通过管线与轻汽油加氢装置后的轻汽油管 线相连接。
一种优选技术方案， 其特征在于： 所述蒸馏装置下部通过重汽油管线 先与重汽油加氢装置相连接；所述重汽油加氢装置通过管线再与所述抽提 系统相连接。 一种优选技术方案， 其特征在于： 所述抽提系统的下部先通过管线与 芳烃加氢装置相连接，所述芳烃加氢装置再通过管线与轻汽油加氢装置后 的轻汽油管线相连接。
本发明的另一目的是提供上述催化烃重组制备高质量汽油的方法。 —种催化烃重组制备高质量汽油的方法， 其步骤如下： 将稳定汽油加 入蒸馏装置进行分馏， 切割分馏出轻汽油和重汽油； 所述轻汽油通过蒸馏 装置上部进入轻汽油加氢装置进行加氢；所述重汽油在抽提系统进行萃取 分离， 分离出芳烃和抽余油； 所述芳烃通过管线与经过加氢处理的轻汽油 调和使用， 所述抽余油作为化工轻油直接采出。
一种优选方案， 其特征在于： 所述重汽油在进入所述抽提系统中进行 萃取分离之前先在重汽油加氢装置中进行加氢处理。
一种优选方案， 其特征在于： 所述芳烃先经过加氢处理后再通过管线 与经过加氢处理的轻汽油调和使用。
一种优选方案，其特征在于：所述蒸馏装置的塔顶温度为 100〜110°C， 塔底温度为 206〜226°C；所述蒸馏装置的塔顶压力为 0. 11〜0. 28MPa(绝）， 塔底压力为 0. 12〜0. 30MPa (绝）； 所述轻汽油的馏程控制在 30°C〜100°C ; 所述重汽油的馏程控制在 100°C〜205°C。
一种优选方案， 其特征在于： 所述蒸馏装置的塔顶温度为 105°C， 塔 底温度为 216°C ; 所述蒸馏装置的塔顶压力为 0. 11〜0. 28MPa (绝）， 塔底 压力为 0. 12〜0. 30MPa (绝）； 所述轻汽油的馏程控制在 30°C〜100°C； 所 述重汽油的馏程控制在 100°C〜205°C。
一种优选方案， 其特征在于： 所述轻汽油加氢装置中的催化剂为选择 性加氢催化剂 GHT-20;所述轻汽油加氢装置的体积空速比为 2-4;氢 /油体 积比为 250〜350； 操作温度为 240〜260°C， 操作压力为 1. 4〜1. 6MPa (绝）。
一种优选方案， 其特征在于： 所述轻汽油加氢装置中的选择性加氢催 化剂 GHT-20的理化性质如下表所示：
Figure imgf000005_0001
强度 N/cm 170 堆密度 g/ml 0. 70
比表面 m /g 180
孔容 ml/g 0. 5-0. 6
wo3 m% 6. 6
NiO m% 2. 1
CoO m% 0. 16 一种优选方案， 其特征在于： 所述重汽油加氢装置中的催化剂为全部 加氢催化剂 GHT-22;所述重汽油加氢装置的体积空速比为 2〜4；氢 /油体 积比为 250〜350；操作温度为 290〜330°C，操作压力为 1.2〜3MPa (绝）。
一种优选方案， 其特征在于： 所述重汽油加氢装置中的全部加氢催化 剂 GHT-22的理化性质如下表所示。
Figure imgf000006_0001
一种优选方案， 其特征在于： 所述芳烃加氢装置中的催化剂为全部加 氢催化剂 GHT-22;所述芳烃加氢装置的体积空速比为 2〜3；氢 /油体积比 为 250〜300； 操作温度为 285〜325 °C， 操作压力为 1.5〜2.5MPa (绝)。
一种优选方案， 其特征在于： 所述芳烃加氢装置中的全部加氢催化剂 GHT-22的理化性质如下表所示。
Figure imgf000006_0002
堆密度 g/ml 0. 73 比表面积 m /g 180
孔容 ml/g 0. 5-0. 6
wo3 m% 15
NiO m% 1. 7
CoO m% 0. 15
Ν¾0 m% <0. 09
Fe203 m% <0. 06
Si02 m% <0. 60
载体 m% 82. 4
本发明所用蒸馏装置为专利号为 03148181.7 的 "催化烃重组处理方 法" 的中国发明专利中公开的蒸馏系统。 所述抽提系统使用专利号为 200310103541.9和 2003 10103540.4中公开的抽提系统， 包括溶剂回收及 水洗系统。
本发明所用加氢装置为现有的加氢装置， 包括加热炉， 换热器， 高压 分离器， 空气冷凝器、 水冷凝器等。
下面通过附图和具体实施方式对本发明做进一步说明，但并不意味着 对本发明保护范围的限制。
附图说明
图 1为本发明实施例 1的流程示意图。
图 2为本发明实施例 2和 3的流程示意图。
图 3为本发明实施例 4和 5的流程示意图。
具体实施方式
实施例 1
如图 1所示， 为本发明实施例的流程示意图。将馏程为 30-205 °C， 含 硫量为 85ppm， 硫醇含量为 15ppm， 烯烃含量为 25% ( v)， 二烯烃含量为 0. 1% ( v)， 芳烃含量为 13% ( v)， 辛烷值（RON) 为 87， 密度为 728千克 / 米 3的稳定汽油 （催化汽油） 以 10万吨 /年的流量在蒸馏塔 1中进行切割 分馏， 蒸馏塔 1 的塔顶温度为 105 °C， 塔底温度为 216°C， 塔顶压力为 0. 2MPa (绝）， 塔底压力为 0. 25MPa (绝）， 分别得到轻汽油和重汽油， 所 述轻汽油 (馏程为 30-10CTC )通过蒸馏塔 1 上部蒸出， 其总的蒸出量为 5 万吨 /年， 然后进入轻汽油加氢装置 3加氢处理； 所述轻汽油加氢装置 3 中的催化剂为选择性加氢催化剂 GHT-20, 所述轻汽油加氢装置 3 的体积 空速比为 4; 氢 /油体积比为 300; 操作温度为 250°C， 操作压力为 1.5MPa (绝） （选择性加氢）； 所述重汽油（馏程为 100-205°C)以 5万吨 /年的流 量进入重汽油抽提系统 2中萃取分离， 分离出芳烃和抽余油； 所述抽提系 统 2中所用溶剂为环丁砜， 萃取温度为 120°C， 溶剂比 （溶剂 /进料） 为 4.0(质量），抽余油水洗比为 0.2 (质量），溶剂回收温度为 175°C，溶剂回收 压力为 0.065MPa (绝）； 所述芳烃以 1.1万吨 /年的流量与所述加氢后的 轻汽油调和， 所述抽余油以 3.9万吨 /年的流量作为化工轻油采出。
所得调和汽油的馏程为 30-205°C， 含硫量为 102.8ppm， 硫醇含量为
4.3ppm， 烯烃含量为 17.9% (v)， 二烯烃含量为 0.05% (v)， 芳烃含量为 20.0% (v)， 辛烷值 （RON) 为 91.8， 密度为 700.6千克 /米 3， 采出量为
6.1万吨 /年。
所得化工轻油的馏程为 100-205°C， 含硫量为 29.0ppm， 硫醇含量为 1.0ppm， 烯烃含量为 28.2% (v)， 二烯烃含量为小于 0.01% (v)， 芳烃含 量为 3.0% (v)， 辛烷值（RON) 为 78.5， 密度为 775.5千克 /米 3， 采出量 为 3.9万吨 /年。
所述选择性加氢催化剂 GHT-20的理化性质如下表所示：
Figure imgf000008_0001
本发明所用测定方法为：
1、 馏程： GB/T6536-1997石油产品蒸馏测定法；
2、 硫含量： SH/T0689-2000 轻质烃及发动机燃料和其他油品的总硫 含量测定法 （紫外荧光法） ；
3、 硫醇硫： GB/T1792-1988 馏分燃料油中硫醇硫测定法 （电位滴定 法）； 4、 烯烃： GB/T11132-2002液体石油产品烃类测定法 （荧光指示剂吸 附法） ；
5、 芳烃： GB/T11132-2002液体石油产品烃类测定法 （荧光指示剂吸 附法） ；
6、 辛烷值： GB/T5487 汽油辛烷值测定法 （研究法） ；
7、密度： GB/T1884-2000原油和液体石油产品密度实验室测定法（密 度计法）；
8、 双烯的测定： 滴定法。
9、 加氢催化剂分析方法:
Figure imgf000009_0001
实施例 2
如图 2所示， 为本实施例的流程示意图。将馏程为 30-205°C， 含硫量 为 100ppm，硫醇含量为 5ppm，烯烃含量为 30% (v)，二烯烃含量为 0. 1% (ν)， 芳烃含量为 15% (v)， 辛烷值 （RON) (RON) 为 89， 密度为 728千克 /米 3 的低硫含量稳定汽油 （催化汽油） 以 10万吨 /年的流量在蒸馏塔 1中进行 切割分馏， 蒸馏塔 1的塔顶温度为 82°C， 塔底温度为 186°C， 塔顶压力为 0. 2MPa (绝）， 塔底压力为 0. 25MPa (绝）， 分别得到轻汽油和重汽油， 所 述轻汽油 (馏程为 30-80°C)通过蒸馏塔 1上部蒸出， 其总的蒸出量为 4万 吨 /年， 然后进入轻汽油加氢装置 3-1加氢处理； 所述轻汽油加氢装置 3-1 的催化剂为选择性加氢催化剂 GHT-20; 所述轻汽油加氢装置 3-1 的体积 空速比为 2 ; 氢 /油体积比为 150; 操作温度为 220°C， 操作压力为 0. 6MPa (绝）； 所述重汽油 (；馏程为 80-205°C)以 6万吨 /年的流量进入重汽油加氢 装置 3-2加氢处理； 所述重汽油加氢装置 3-2中的催化剂为全部加氢催化 剂 GHT-22; 所述重汽油加氢装置 3-2的体积空速比为 2; 氢 /油体积比为 250； 操作温度为 290°C， 操作压力为 1.2MPa (绝）； 然后通过管线进入重 汽油抽提系统 2中萃取分离， 分离出芳烃和抽余油； 所述抽提系统 2中所 用溶剂为 N-甲基吡咯烷酮， 萃取温度为 115°C， 溶剂比 （溶剂 /进料） 为 3.5 (质量）， 抽余油水洗比为 0.2 (质量） ，溶剂回收温度为 151°C，溶剂 回收压力为 0.112MPa (绝）； 所述芳烃以 1.5万吨 /年的流量与所述加氢 后的轻汽油调和，所述抽余油以 4.5万吨 /年的流量作为优质乙烯料采出。
所得调和汽油的馏程为 30-205°C， 含硫量为 5.27ppm， 硫醇含量为小 于 lppm， 烯烃含量为 17.8% (v)， 二烯烃含量为小于 0.01% (v)， 芳烃含量 为 25.6%(v)， 辛烷值（RON) 为 94.1， 密度为 703.8千克 /米 3， 采出量为 5.5万吨 /年。
所得优质乙烯料的馏程为 80-205°C， 含硫量为 2.0ppm， 硫醇含量为 小于 lppm， 烯烃含量为小于 0.1%(ν)， 二烯烃含量为小于 0.01% (v)， 芳 烃含量为 3.0%(v)， 辛烷值 （RON) 为 81.0， 密度为 760.0千克 /米 3， 采 出量为 4.5万吨 /年。
所述选择性加氢催化剂 GHT-20的理化性质如下表所示。
Figure imgf000010_0001
所述全部加氢催化剂 GHT-22的理化性质如下表所示。
Figure imgf000010_0002
规格 mm Φ 1. 7
强度 N/cm 180
堆密度 g/ml 0. 73
2 1
比表面积 m /g 180
孔容 ml/g 0. 57
wo3 m% 15
NiO m% 1. 7
CoO m% 0. 15
Ν¾0 m% <0. 09
Fe203 m% <0. 06
Si02 m% <0. 60
载体 m% 82. 4 本发明所用测定方法为 （下同）：
1、 馏程： GB/T6536-1997石油产品蒸馏测定法；
2、 硫含量： SH/T0689-2000 轻质烃及发动机燃料和其他油品的总硫 含量测定法 （紫外荧光法） ；
3、 硫醇硫： GB/T1792-1988 馏分燃料油中硫醇硫测定法 （电位滴定 法）；
4、 烯烃： GB/T11132-2002液体石油产品烃类测定法 （荧光指示剂吸 附法） ；
5、 芳烃： GB/T11132-2002液体石油产品烃类测定法 （荧光指示剂吸 附法） ；
6、 辛烷值： GB/T5487 汽油辛烷值测定法 （研究法） ；
7、密度： GB/T1884-2000原油和液体石油产品密度实验室测定法（密 度计法）；
8、 双烯的测定： 滴定法。
9、 加氢催化剂分析方法:
Figure imgf000011_0001
表面积 低温氮吸附法 2400型吸附仪
孔容 压汞法 Auto Pore II 9200
强度 抗压碎强度测定 DLII型智能颗粒强度测定仪
法
堆密度 称量法 实施例 3
如图 2所示， 为本实施例的流程示意图。将馏程为 30-205°C， 含硫量 为 2000ppm， 硫醇含量为 50ppm， 烯烃含量为 40% (v)， 二烯烃含量为 1.0%(v)， 芳烃含量为 19%(v)， 辛烷值 （RON) (RON) 为 91， 密度为 728 千克 /米 3的高硫含量稳定汽油（催化汽油） 以 10万吨 /年的流量在蒸馏塔 1中进行切割分馏， 蒸馏塔 1的塔顶温度为 86°C， 塔底温度为 192°C， 塔 顶压力为 0.2MPa (绝）， 塔底压力为 0.25MPa (绝）， 分馏分别得到轻汽油 和重汽油， 所述轻汽油 (馏程为 30-90°C)通过蒸馏塔 1上部蒸出， 其总的 蒸出量为 4.3万吨 /年， 然后进入轻汽油加氢装置 3-1加氢处理； 所述轻汽 油加氢装置 3-1 的催化剂为选择性加氢催化剂 GHT-20; 所述轻汽油加氢 装置 3-1的体积空速比为 4; 氢 /油体积比为 300; 操作温度为 280°C， 操 作压力为 2. OMPa (绝）； 所述重汽油 (馏程为 90-205 )以 5.7万吨 /年的 流量进入重汽油加氢装置 3-2加氢处理； 所述重汽油加氢装置 3-2中的催 化剂为选择性加氢催化剂 GHT-20; 所述重汽油加氢装置 3-2 的体积空速 比为 4;氢 /油体积比为 350;操作温度为 330°C，操作压力为 3. OMPa (绝）； 然后通过管线进入重汽油抽提系统 2中萃取分离， 分离出芳烃和抽余油； 所述抽提系统 2中所用溶剂为 N-甲基吡咯烷酮， 萃取温度为 115°C， 溶剂 比 （溶剂 /进料） 为 3.5 (质量）， 抽余油水洗比为 0.2 (质量） ，溶剂回收 温度为 151°C，溶剂回收压力为 0.112MPa (绝）； 所述芳烃以 1.9万吨 /年 的流量与所述加氢后的轻汽油调和， 所述抽余油以 3.8万吨 /年的流量作 为优质乙烯料采出。
所得调和汽油的馏程为 30-205°C， 含硫量为 7.52ppm， 硫醇含量为小 于 lppm， 烯烃含量为 17.99%(v)， 二烯烃含量为小于 0.01% (v)， 芳烃含 量为 29.1%(ν)， 辛烷值（RON) 为 95.2, 密度为 720.1千克 /米 3， 采出量 为 6. 2万吨 /年。
所得优质乙烯料的馏程为 90-205 °C， 含硫量为 2. 0ppm， 硫醇含量为 小于 lppm， 烯烃含量为 6% (v)， 二烯烃含量为小于 0. 01% (v)， 芳烃含量 为 3. 0% (v)， 辛烷值 （RON) 为 81. 5， 密度为 740. 0千克 /米 3， 采出量为 3. 8万吨 /年。
实施例 4
如图 3所示， 是本实施例的流程示意图。将馏程为 30-205°C， 含硫量 为 lOOppm，硫醇含量为 5ppm，烯烃含量为 30% (v)，二烯烃含量为 0. 1% (ν)， 芳烃含量为 15% (v)， 辛烷值 （RON) 为 89， 密度为 728千克 /米 3的低硫 含量稳定汽油 （催化汽油） 以 10万吨 /年的流量在蒸馏塔 1中进行切割分 馏， 蒸馏塔 1 的塔顶温度为 86 °C， 塔底温度为 192 °C， 塔顶压力为 0. 2MPa (绝）， 塔底压力为 0. 25MPa (绝）， 分馏分别得到轻汽油和重汽油， 所述轻汽油 (馏程为 30-80°C)通过蒸馏塔 1上部蒸出后在轻汽油加氢装置 3-1中加氢处理， 加氢后的轻汽油总的蒸出量为 4万吨 /年； 所述轻汽油加 氢装置 3-1中的催化剂为选择性加氢催化剂 GHT-20;所述轻汽油加氢装置 3-1的体积空速比为 2 ; 氢 /油体积比为 150; 操作温度为 230°C， 操作压 力为 1. OMPa (绝）； 所述重汽油 (馏程为 80-205°C)以 6万吨 /年的流量通过 管线进入重汽油抽提系统 2中萃取分离， 分离出芳烃和抽余油； 所述抽提 系统 2中所用溶剂为 N-甲酰基吗啉， 萃取温度为 115°C， 溶剂比 （溶剂 / 进料）为 3. 5 (质量）， 抽余油水洗比为 0. 2 (质量），溶剂回收温度为 151 V，溶剂回收压力为 0. 112MPa (绝）； 所述芳烃以 1. 5万吨 /年的流量进入 芳烃加氢装置 3-2加氢， 所述芳烃加氢装置 3-2中的催化剂为全部加氢催 化剂 GHT-22 ; 所述芳烃加氢装置 3-2的体积空速比为 2 ; 氢 /油体积比为 250； 操作温度为 295°C， 操作压力为 2. OMPa (绝）； 然后再与所述加氢后 的轻汽油调和， 所述抽余油以 1. 5万吨 /年的流量作为化工轻油采出。
所得调和汽油的馏程为 30-205°C， 含硫量为 4. 2ppm， 硫醇含量为小 于 lppm， 烯烃含量为 17. 8% (v)， 二烯烃含量为小于 0. 01% (v)， 芳烃含量 为 25. 6% (v)， 辛烷值（RON) 为 94. 1， 密度为 703. 8千克 /米 3， 采出量为 5. 5万吨 /年。
所得化工轻油的馏程为 80-205 °C， 含硫量为 10. 0ppm， 硫醇含量为 l. Oppm, 烯烃含量为 35. 5% (v)， 二烯烃含量为 0. 01% (v)， 芳烃含量为 3. 0% (v)， 辛烷值（RON) 为 81. 0， 密度为 760. 0千克 /米 3， 采出量为 4. 5 万吨 /年。
所述选择性加氢催化剂 GHT-20的理化性质如下表所示。
Figure imgf000014_0001
实施例 5
如图 3所示， 是本实施例的流程示意图。将馏程为 30-205°C， 含硫量 为 2000ppm， 硫醇含量为 50ppm， 烯烃含量为 40% (v)， 二烯烃含量为 1. 0% (v)， 芳烃含量为 19% (v)， 辛烷值 （RON) 为 91， 密度为 728千克 / 米 3的高硫含量稳定汽油 （催化汽油） 以 10万吨 /年的流量在蒸馏塔 1中 进行切割分馏， 蒸馏塔 1的塔顶温度为 86°C， 塔底温度为 192°C， 塔顶压 力为 0. 2MPa (绝）， 塔底压力为 0. 25MPa (绝）， 分馏分别得到轻汽油和重汽 油， 所述轻汽油 (馏程为 30-90°C)通过蒸馏塔 1上部蒸出后在轻汽油加氢 装置 3-1中加氢处理， 加氢后的轻汽油总的蒸出量为 4.3万吨 /年； 所述轻 汽油加氢装置 3-1中的催化剂为选择性加氢催化剂 GHT-20;所述轻汽油加 氢装置 3-1的体积空速比为 4; 氢 /油体积比为 300; 操作温度为 250°C， 操作压力为 1. OMPa (绝）； 所述重汽油 (；馏程为 90-205°C)以 5. 7万吨 /年的 流量通过管线进入重汽油抽提系统 2中萃取分离， 分离出芳烃和抽余油； 所述抽提系统 2中所用溶剂为 N-甲酰基吗啉， 萃取温度为 115°C， 溶剂比 (溶剂 /进料） 为 3. 5 (质量）， 抽余油水洗比为 0. 2 (质量） ，溶剂回收温 度为 151 °C，溶剂回收压力为 0. 112MPa (绝）；所述芳烃以 1. 5万吨 /年的流 量进入芳烃加氢装置 3-2加氢， 所述芳烃加氢装置 3-2中的催化剂为全部 加氢催化剂 GHT-22 ; 所述芳烃加氢装置 3-2的体积空速比为 3 ; 氢 /油体 积比为 300; 操作温度为 325°C， 操作压力为 2. 5MPa (绝）； 然后再与所述 加氢后的轻汽油调和， 所述抽余油以 1. 5万吨 /年的流量作为化工轻油采 出。
所得调和汽油的馏程为 30-205°C， 含硫量为 10. 0ppm， 硫醇含量为小 于 lppm， 烯烃含量为 17. 84% (v)， 二烯烃含量为小于 0. 01% (v)， 芳烃含 量为 28. 2% (v)， 辛烷值 （RON) 为 94. 05， 密度为 721. 4千克 /米 3， 采出 量为 6. 4万吨 /年。
所得化工轻油的馏程为 90-205 °C， 含硫量为 10. 0ppm， 硫醇含量为 1. 0ppm， 烯烃含量为 58. 3% (v)， 二烯烃含量为 0. 01% (v)， 芳烃含量为 3. 0% (v)， 辛烷值（RON) 为 82. 0， 密度为 740. 0千克 /米 3， 采出量为 3. 6 万吨 /年。
工业应用性
本发明的优点是：
与前加氢（在稳定汽油进入蒸馏装置之前进行加氢处理， 其所用的催 化剂的量大， 并且只能降低双烯和硫醇的量）相比； 本发明的制备低含硫 量和低烯烃含量汽油的催化烃重组处理系统及其方法的优点是： 首先， 由 于专门针对轻汽油， 和 /或重汽油， 和 /或芳烃进行加氢处理， 所用催化剂 的量大大减少； 其次， 针对轻汽油， 和 /或重汽油， 和 /或芳烃， 所用催化 剂的量小得多，另外，可以降低单烯和双烯的含量，不仅降低硫醇的含量， 还可以降低总的含硫量； 最后， 选择性加氢， 采用特定的催化剂和参数， 主要是脱硫醇， 解决烯烃和双烯问题， 效果明显。

Claims

权 利 要 求 书
1、 一种催化烃重组制备高质量汽油的系统， 包括蒸馏装置； 其特征 在于： 所述蒸馏装置上部通过轻汽油管线与轻汽油加氢装置相连接； 所述 蒸馏装置下部通过重汽油管线与抽提系统相连接；所述抽提系统上部通过 管线直接采出产品，所述抽提系统下部通过管线与轻汽油加氢装置后的轻 汽油管线相连接。
2、 根据权利要求 1所述的催化烃重组制备高质量汽油的系统， 其特 征在于： 所述蒸馏装置下部通过重汽油管线先与重汽油加氢装置相连接； 所述重汽油加氢装置通过管线再与所述抽提系统相连接。
3、 根据权利要求 1所述的催化烃重组制备高质量汽油的系统， 其特 征在于： 所述抽提系统的下部先通过管线与芳烃加氢装置相连接， 所述芳 烃加氢装置再通过管线与轻汽油加氢装置后的轻汽油管线相连接。
4、 一种催化烃重组制备高质量汽油的方法， 其步骤如下： 将稳定汽 油加入蒸馏装置进行分馏， 切割分馏出轻汽油和重汽油； 所述轻汽油通过 蒸馏装置上部进入轻汽油加氢装置进行加氢；所述重汽油在抽提系统进行 萃取分离， 分离出芳烃和抽余油； 所述芳烃通过管线与经过加氢处理的轻 汽油调和使用， 所述抽余油作为化工轻油直接采出。
5、 根据权利要求 4所述的催化烃重组制备高质量汽油的方法， 其特 征在于：所述重汽油在进入所述抽提系统中进行萃取分离之前先在重汽油 加氢装置中进行加氢处理。
6、 根据权利要求 4所述的催化烃重组制备高质量汽油的方法， 其特 征在于：所述芳烃经过芳烃加氢装置进行加氢处理后再通过管线与经过加 氢处理的轻汽油调和使用。
7、根据权利要求 4一 6中任一项所述的催化烃重组制备高质量汽油的 方法， 其特征在于： 所述蒸馏装置的塔顶温度为 100〜110°C， 塔底温度为 206〜226°C ; 所述蒸馏装置的塔顶压力为 0. l l〜0. 28MPa (绝）， 塔底压力 为 0. 12〜0. 30MPa (绝）； 所述轻汽油的馏程控制在 30°C〜100°C； 所述重 汽油的馏程控制在 100°C〜205 °C。
8、根据权利要求 4一 6中任一项所述的催化烃重组制备高质量汽油的 方法， 其特征在于： 所述蒸馏装置的塔顶温度为 105 °C， 塔底温度为 216 °C ;所述蒸馏装置的塔顶压力为 0. 11〜0. 28MPa (绝），塔底压力为 0. 12〜 0. 30MPa (绝）； 所述轻汽油的馏程控制在 30°C〜100°C ; 所述重汽油的馏 程控制在 100°C~205 °C。
9、 根据权利要求 4所述的催化烃重组制备高质量汽油的方法， 其特 征在于： 所述轻汽油加氢装置中的催化剂为选择性加氢催化剂 GHT-20; 所述轻汽油加氢装置的体积空速比为 2-4;氢 /油体积比为 250〜350；操作 温度为 240〜260°C， 操作压力为 1. 4〜1. 6MPa (绝）。
10、根据权利要求 9所述的催化烃重组制备高质量汽油的方法， 其特 征在于： 所述轻汽油加氢装置中的选择性加氢催化剂 GHT-20的理化性质 如下表所示：
Figure imgf000017_0001
11、根据权利要求 5所述的催化烃重组制备高质量汽油的方法， 其特 征在于： 所述重汽油加氢装置中的催化剂为全部加氢催化剂 GHT-22 ; 所 述重汽油加氢装置的体积空速比为 2〜4； 氢 /油体积比为 250〜350； 操作 温度为 290〜330°C， 操作压力为 1. 2〜3MPa (绝）。
12、 根据权利要求 11所述的催化烃重组制备高质量汽油的方法， 其 特征在于： 所述重汽油加氢装置中的全部加氢催化剂 GHT-22的理化性质 如下表所示。
Figure imgf000017_0002
规格 mm Φ 1. 5-2. 0
强度 N/cm 180
堆密度 g/ml 0. 73
比表面积 m /g 180
孔容 ml/g 0. 5-0. 6
wo3 m% 15
NiO m% 1. 7
CoO m% 0. 15
Ν¾0 m% <0. 09
Fe203 m% <0. 06
Si02 m% <0. 60
载体 m% 82. 4
13、根据权利要求 6所述的催化烃重组制备高质量汽油的方法， 其特 征在于： 所述芳烃加氢装置中的催化剂为全部加氢催化剂 GHT-22 ; 所述 芳烃加氢装置的体积空速比为 2〜3； 氢 /油体积比为 250〜300； 操作温度 为 285〜325 °C， 操作压力为 1 · 5〜2· 5MPa (绝）。
14、 根据权利要求 13所述的催化烃重组制备高质量汽油的方法， 其 特征在于： 所述芳烃加氢装置中的全部加氢催化剂 GHT-22的理化性质如 下表所示。
Figure imgf000018_0001
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CA2705034C (en) 2016-10-11
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US20100243522A1 (en) 2010-09-30
EP2236583A4 (en) 2013-01-30
JP2011503264A (ja) 2011-01-27
US8524043B2 (en) 2013-09-03
EP2236583A1 (en) 2010-10-06

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