WO2010083642A1 - 催化烃重组后加氢制备高质量汽油的系统和方法 - Google Patents

催化烃重组后加氢制备高质量汽油的系统和方法 Download PDF

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Publication number
WO2010083642A1
WO2010083642A1 PCT/CN2009/070238 CN2009070238W WO2010083642A1 WO 2010083642 A1 WO2010083642 A1 WO 2010083642A1 CN 2009070238 W CN2009070238 W CN 2009070238W WO 2010083642 A1 WO2010083642 A1 WO 2010083642A1
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Prior art keywords
gasoline
oil
hydrogenation
extraction
temperature
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PCT/CN2009/070238
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English (en)
French (fr)
Inventor
丁冉峰
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北京金伟晖工程技术有限公司
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Priority to EP09838609.7A priority Critical patent/EP2390303B1/en
Priority to CA2717982A priority patent/CA2717982C/en
Priority to EA201071241A priority patent/EA019489B1/ru
Priority to BRPI0909889A priority patent/BRPI0909889A2/pt
Priority to US12/918,636 priority patent/US8419930B2/en
Priority to JP2011503329A priority patent/JP5543957B2/ja
Priority to PCT/CN2009/070238 priority patent/WO2010083642A1/zh
Publication of WO2010083642A1 publication Critical patent/WO2010083642A1/zh

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Classifications

    • BPERFORMING OPERATIONS; TRANSPORTING
    • B01PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01DSEPARATION
    • B01D3/00Distillation or related exchange processes in which liquids are contacted with gaseous media, e.g. stripping
    • B01D3/009Distillation or related exchange processes in which liquids are contacted with gaseous media, e.g. stripping in combination with chemical reactions
    • BPERFORMING OPERATIONS; TRANSPORTING
    • B01PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01DSEPARATION
    • B01D3/00Distillation or related exchange processes in which liquids are contacted with gaseous media, e.g. stripping
    • B01D3/14Fractional distillation or use of a fractionation or rectification column
    • B01D3/143Fractional distillation or use of a fractionation or rectification column by two or more of a fractionation, separation or rectification step
    • CCHEMISTRY; METALLURGY
    • C10PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10GCRACKING 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
    • C10G67/00Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only
    • CCHEMISTRY; METALLURGY
    • C10PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10GCRACKING 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
    • C10G67/00Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only
    • C10G67/02Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only plural serial stages only
    • C10G67/04Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only plural serial stages only including solvent extraction as the refining step in the absence of hydrogen
    • C10G67/0409Extraction of unsaturated hydrocarbons
    • CCHEMISTRY; METALLURGY
    • C10PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10GCRACKING 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
    • C10G2400/00Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/02Gasoline
    • CCHEMISTRY; METALLURGY
    • C10PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10GCRACKING 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
    • C10G2400/00Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/04Diesel oil

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 high quality gasoline by recombination of component refinery hydrocarbons.
  • Catalytic cracking, catalytic cracking and 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, usually fractionated.
  • the tower fractionation can distill off dry gas, liquefied gas, gasoline, diesel, heavy oil and other products. Among them, 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 Patent 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 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). 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), and 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 ⁇ gasoline standard to the national IV gasoline standard.
  • a better planning plan should be a one-time national IV gasoline standard planning plan.
  • the present invention adopts the following technical solutions:
  • a system for recombining a refinery hydrocarbon to produce high quality gasoline after recombination comprising: an extraction system, a distillation system and a hydrogenation unit, wherein an upper portion of the extraction system is connected to the distillation system through a pipeline; The lower part of the extraction system is connected to the hydrogenation unit through a pipeline, and the hydrogenation unit is connected to a pipeline at the upper part of the distillation system through a pipeline; the upper part of the distillation system directly extracts a product through a pipeline, the distillation system The middle part is connected to another hydrogenation unit through a pipeline; the lower part of the distillation system is directly produced through a pipeline
  • a system for recombining a refinery hydrocarbon to produce high quality gasoline after recombination comprising: an extraction system, a distillation system and a hydrogenation unit, wherein an upper portion of the extraction system is connected to the distillation system through a pipeline;
  • the lower part of the extraction system is connected to the oil extraction unit via a pipeline;
  • the upper part of the distillation system is used to produce light gasoline through a pipeline;
  • the lower part of the distillation system is connected to a heavy gasoline hydrogenation unit via a pipeline, the heavy gasoline hydrogenation unit
  • the lower part produces a reformate or a vinyl material through a pipeline.
  • Another object of the present invention is to provide a method of preparing high quality gasoline as described above.
  • a method for preparing high-quality gasoline after recombination of refinery hydrocarbons the steps of which are as follows: the raw materials enter the extraction system for extraction and separation, and the raffinate oil and the extracted oil are separated; the raffinate oil of the extraction system enters the distillation
  • the system performs cutting fractionation, the light gasoline is distilled out from the upper part of the distillation system, and is extracted as a blended gasoline; the chemical light oil is extracted from the middle side line of the distillation system; the chemical light oil enters the chemical light oil hydrogenation unit for hydrotreating;
  • the hydrogenated chemical light oil is produced as high quality ethylene or reforming material;
  • the extracted oil in the extraction system enters the extraction oil hydrogenation unit for hydrotreating; the hydrogenated oil is extracted and the light is Gasoline blended as a blending steam Oil is produced; the diesel oil after the distillation system is directly produced.
  • the bottom pressure is 0. 20 ⁇ 0. 30MPa (absolute); the distillation range of the light gasoline is controlled at 30 ° C ⁇ 110 ° C; the distillation range of the chemical light oil is 110 ⁇ 160 ° C; the distillation range of the diesel oil is 160 to 205 ° C.
  • a preferred embodiment is characterized in that: the solvent used in the extraction system is sulfolane, the extraction temperature is 120 ° C, the solvent ratio (solvent / feed) is 3.5 (mass), and the raffinate washing ratio is 0.2 (mass) lMPa ( ⁇ ).
  • the solvent recovery temperature is 165 ° C, the solvent recovery pressure is O. lMPa (absolute).
  • a preferred embodiment is characterized in that: the solvent used in the extraction system is N-methylpyrrolidone, the extraction temperature is 130 ° C, the solvent ratio (solvent / feed) is 2.5 (mass), and the residual oil is washed. The ratio is 0.25 (mass;), the solvent recovery temperature is 177 ° C, and the solvent recovery pressure is 0.15 MPa (absolute).
  • a preferred embodiment is characterized in that: the catalyst in the chemical light oil hydrogenation unit is all hydrogenation catalyst GHT-22; the volumetric space velocity ratio of the chemical light oil hydrogenation unit is l ⁇ h; hydrogen/oil The volume ratio is 250-500; the operating temperature is 250 ⁇ 320 °C, and the operating pressure is l ⁇ 4MPa (absolute).
  • a preferred embodiment is characterized in that: the chemical light oil hydrogenation unit has a volumetric space velocity ratio of 2.5 h - a hydrogen/oil volume ratio of 300; an operating temperature of 285 ° C and an operating pressure of 2.5 MPa (absolute).
  • a preferred embodiment is characterized in that: the volumetric airspeed ratio of the extracted oil hydrogenation unit is 2.51 ⁇ ; the hydrogen/oil volume ratio is 300; the operating temperature is 270 ° C, and the operating pressure is 2.5 MPa (absolute).
  • a method for preparing high-quality gasoline after recombination of refinery hydrocarbons the steps of which are as follows:
  • the raw materials enter the extraction system for extraction and separation, and the raffinate oil and the extracted oil are separated; the raffinate oil separated by the extraction system Entering the distillation system for cutting fractionation, the upper part of the distillation system is steamed out of light gasoline, and is produced as a blended gasoline; the lower part of the distillation system is steamed out of heavy gasoline, and enters a heavy gasoline hydrogenation unit for hydrotreating; the heavy gasoline
  • the hydrotreated heavy gasoline is produced as a vinyl material or a reforming material; the extracted oil separated by the extraction system enters a pumping oil hydrogenation unit for hydrotreating; and the hydrogenated oil is mixed with the light gasoline. After the recovery as a blend of gasoline.
  • the singularity of the bottom of the column is 0.25 ° C.
  • the pressure at the bottom of the column is 0. 20 MPa (absolute), the pressure at the bottom of the column is 0. 25 MPa (absolutely ).
  • a preferred embodiment is characterized in that: the solvent used in the extraction system is sulfolane, and the extraction temperature is
  • a preferred embodiment is characterized in that: the solvent used in the extraction system is N-methylpyrrolidone, the extraction temperature is 130 ° C, the solvent ratio (solvent / feed) is 2.5 (mass), and the raffinate oil wash ratio The temperature was 0.25 (mass), the solvent recovery temperature was 177 ° C, and the solvent recovery pressure was 0.15 MPa (absolute).
  • a preferred embodiment is characterized in that: the solvent used in the extraction system is N-formylmorpholine, the extraction temperature is 150 ° C, the solvent ratio (solvent / feed) is 6.0 (mass), and the residual oil is washed. The ratio was 0.3 (mass), the solvent recovery temperature was 185 ° C, and the solvent recovery pressure was 0.2 MPa (absolute).
  • the heavy gasoline hydrogenation unit and the catalyst in the oil extraction unit are all hydrogenation catalysts GHT-22; the volumetric space velocity ratio is 1.0 to 4. Oh" 1 ; / oil volume ratio is 250 ⁇ 500; operating temperature is 250 ⁇ 290 ° C, operating pressure is 1.0 ⁇ 4 ⁇ OMPa (absolute).
  • the sufficiency of the operating temperature is 270 ° C, the operating pressure is 2. 50 MPa (absolutely ).
  • the tempering ratio is 2.50 MPa; the operating temperature is 285 ° C, the operating pressure is 2. 50 MPa (absolutely ).
  • a preferred embodiment is characterized in that the physical and chemical properties of the entire hydrogenation catalyst GHT-22 in the heavy gasoline hydrogenation unit and the extracted oil hydrogenation unit are as shown in the following table.
  • the solvent of the present invention may also be another solvent or a mixture of two or more of these solvents in any ratio.
  • the naphtha, stabilized gasoline and hydrocoking gasoline of the present invention may be in any ratio.
  • the cutting point (distillation range) of the light gasoline, the chemical light oil, and the diesel oil of the present invention can be adjusted.
  • the distillation range of light gasoline is controlled at 30 ° C to 70 ° C; the distillation range of the chemical light oil is 70 to 160 ° C; the distillation range of the diesel oil is 160 to 205 ° C; The distillation range of the chemical light oil is from 90 to 160 ° C; the distillation range of the diesel oil is from 160 to 205 ° C.
  • 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.
  • Embodiment 4 of the present invention is a schematic flow chart of Embodiment 4 of the present invention.
  • Figure 4 is a flow chart showing the embodiments 5 and 6 of the present invention.
  • FIG. 1 is a schematic flowchart of Embodiment 1 of the present invention.
  • the distillation range is 35-205 ° 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)
  • the aromatic content is 15% (v).
  • the raffinate oil from the upper part of the extraction tower 1 enters the distillation column 2 at a flow rate of 48,000 tons/year for cutting fractionation, and the temperature of the top of the distillation column 2 is 77 ° C, the bottom temperature is 173 ° C, the top pressure is 0.15 MPa (absolute), the bottom pressure is 0.20 MPa (absolute), respectively, light gasoline, chemical light oil and diesel.
  • the light gasoline (distillation range 30-110 ° C) was distilled off through the upper portion of the distillation column 2, and was produced as a blended gasoline at a flow rate of 19,500 tons/year.
  • the chemical light oil (distillation range 110-160 ° C) is withdrawn through the side line of the distillation column 2, and the total amount of steam distillation is 22,500 tons / year; then enters the hydrogenation unit 3-1 at a flow rate of 22,500 tons / year
  • the hydrotreating is carried out, the catalyst in the hydrogenation unit 3-1 is the entire hydrogenation catalyst GHT-22; the volumetric space velocity ratio of the hydrogenation unit 3-1 is 41; the hydrogen/oil volume ratio is 500;
  • the operating temperature is 320 ° C, and the operating pressure is 4 MPa (absolute); the hydrogenated chemical light oil is produced as a high-quality vinyl or reforming material at a flow rate of 22,500 tons/year.
  • the extracted oil from the lower portion of the extraction column 1 is hydrotreated at a flow rate of 12,000 tons/year into the hydrogenation unit 3-2, and the catalyst in the hydrogenation unit 3-2 is the entire hydrogenation catalyst GHT-22.
  • the hydrogenation device 3-2 has a volumetric space velocity ratio of 41 ⁇ ; a hydrogen/oil volume ratio of 500; an operating temperature of 290 ° C, an operating pressure of 4 MPa (absolute); and a hydrogenated oil of 1.20 million
  • the flow rate per ton/year is mixed with the light gasoline and then produced as blended gasoline.
  • the diesel oil cut by the distillation column 2 (distillation range 160-205 ° C) was directly produced as a diesel product at 0.6 million tons/year.
  • the obtained blended gasoline has a distillation range of 30-205 ° C, a sulfur content of 3.1 ppm, a mercaptan content of less than 1.0 ppm (a trace amount, not detected), an olefin content of 24.1% (v), a diene content of 0.05. % (v); aromatic content is 26.7% (v), the octane number (RON) is 95.3, the density is 727.6 kg/ m3 , and the amount of production is 31,500 tons/year.
  • the distillation range of the obtained high-quality ethylene or reforming material is 110-160 ° C, the sulfur content is trace amount, the detection is not found, the mercaptan content is less than 1.0 ppm (the trace amount, the detection is not obtained;), the olefin content is Traces, undetectable, bromine index (bromine number) of 39 (0.039), aromatics content of 1.3% (v), octane number (RON) of 83.0, density of 729.0 kg/ m3 , yield of 2.25 10,000 tons / year.
  • the obtained diesel oil has a distillation range of 160-205 ° C, a sulfur content of 26.3 ppm, a mercaptan content of 1.55 ppm, an olefin content of 27.8% (v), a diene content of 0.04% (v), and an aromatic content of 5.6% (v). ), the cetane number is 45.8, the density is 751.7 kg/ m3 , and the recovery is 0.6 million tons/year.
  • 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 (violet fluorescence method);
  • mercaptan sulfur GB/T1792-1988 distillate fuel oil in the determination of mercaptan sulfur (potentiometric titration);
  • 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);
  • octane number GB/T5487 gasoline octane number determination method (research method); 7.
  • Density GB/T1884-2000 laboratory method for density determination of crude oil and liquid petroleum products (densitometer method);
  • FIG. 2 is a schematic flowchart of Embodiment 2 of the present invention.
  • the distillation range is 35-205 ° 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).
  • Stable gasoline catalytic gasoline
  • RON octane number
  • a density of 728 kg/ m3 is extracted and extracted in the extraction column 1 at a flow rate of 60,000 tons/year; at the same time, the distillation range is 30-205.
  • sulfur content is 200ppm, mercaptan content is lppm, olefin content is less than 0.1% (v) (trace, not detected), diolefin content is less than 0.01% (v) (; trace, not detected; ), the aromatics content is 8% (v), the octane number (RON) is 82, and the naphtha having a density of 732 kg/ m3 is extracted and separated in the extraction tower 1 at a flow rate of 20,000 tons/year;
  • the distillation range is 30-205 ° C
  • the sulfur content is 150 ppm
  • the mercaptan content is 1 ppm
  • the olefin content is 6% (v)
  • the diene content is less than 0.01% (v) (trace, not detected)
  • the aromatic hydrocarbon content is 10% (v)
  • the octane number (RON) is 79, and the hydrogenated coking gasoline having a density of 721 kg/ m3 is extracted and separated
  • the temperature at the top of the distillation column 2 is 87 ° C
  • the temperature at the bottom of the column is 184 ° C
  • the pressure at the top of the column is 0.2 MPa (absolute)
  • the pressure at the bottom of the column is 0.25 MPa (absolute), respectively.
  • Light gasoline, chemical gasoline and diesel The light gasoline (distillation range of 30-110 ° C) was distilled off through the upper portion of the distillation column 2, and was produced as a blended gasoline at a flow rate of 33,000 tons/year.
  • the chemical light oil (distillation range 110-160 ° C) is drawn through the side line of the distillation column 2, The total amount of steaming is 41,000 tons/year; then it is subjected to hydrotreating at a flow rate of 41,000 tons/year, and the catalyst in the hydrogenation unit 3-1 is the entire hydrogenation catalyst.
  • GHT-22; the hydrogenation device 3-1 has a volumetric space velocity ratio of 1.0H; a hydrogen/oil volume ratio of 250; an operating temperature of 250V, and an operating pressure of 1.0 MPa (absolute); It is produced as a high-quality vinyl or reforming material at a flow rate of 41,000 tons/year.
  • the extracted oil in the extraction tower 1 enters the hydrogenation unit 3-2 at a flow rate of 16,000 tons/year for hydrotreating, and the catalyst in the hydrogenation unit 3-2 is the entire hydrogenation catalyst GHT-22;
  • the hydrogenation device 3-2 has a volumetric space velocity ratio of 1.Oh" 1 ; a hydrogen/oil volume ratio of 250; an operating temperature of 250 ° C, an operating pressure of 1.0 MPa (absolute); and a hydrogenation extraction
  • the oil is mixed with the light gasoline at a flow rate of 16,000 tons/year and then produced as a blended gasoline.
  • the diesel oil cut through the distillation column 2 (distillation range 160-205 ° C) is directly used as a diesel product of 1.0 million tons/year. Out.
  • the obtained blended gasoline has a distillation range of 30-205 ° C, a sulfur content of 3.4 ppm, a mercaptan content of less than 1.0 ppm (a trace amount, not detected), an olefin content of 16.6% (v), and a diene content of 0.05. % (v);
  • the aromatic content is 23.7% (v)
  • the octane number (RON) is 95.0
  • the density is 719.2 kg/ m3
  • the recovery is 49,000 tons/year.
  • the distillation range of the obtained high-quality ethylene or reforming material is 110-160 ° C, the sulfur content is trace amount, the detection is not found, the mercaptan content is less than 1.0 ppm (the trace amount, the detection is not obtained;), the olefin content is Traces, undetectable, bromine index (bromine number) of 32 (0.032), aromatics content of 1.3% (v), octane number (RON) of 76.2, density of 731.0 kg/ m3 , yield of 4.1 10,000 tons / year.
  • the obtained diesel oil has a distillation range of 160-205 ° C, a sulfur content of 46.5 ppm, a mercaptan content of 0.17 ppm (a trace amount, which is not detected), an olefin content of 15.0% (v), and a diene content of 0.04%.
  • the aromatics content was 3.6% (v)
  • the cetane number was 45.5
  • the density was 753.9 kg/ m3
  • the amount of production was 1.0 million tons/year.
  • the procedure is the same as in Embodiment 2.
  • the distillation range is 30-205 ° C, the sulfur content is 800 ppm, the mercaptan content is 9 ppm, the olefin content is 36% (v), the diolefin content is 0.9% (v), and the aromatic content is 17% (v).
  • Stable gasoline catalytic gasoline with an octane number (RON) of 91 and a density of 731 kg/ m3 is extracted and extracted in the extraction column 1 at a flow rate of 60,000 tons/year; at the same time, the distillation range is 30-205.
  • sulfur content is 200ppm
  • mercaptan content is lppm
  • olefin content is less than 0.1% (v) (trace, not detected)
  • diolefin content is less than 0.01% (v) (trace, not detected)
  • aromatic content 8% (v) an octane number (RON) of 82
  • a naphtha having a density of 732 kg/ m3 is extracted and separated in the extraction column 1 at a flow rate of 20,000 tons/year; at the same time,
  • the distillation range is 30-205
  • the sulfur content is 150ppm
  • the mercaptan content is lppm
  • the olefin content is 6% (v)
  • the diolefin content is less than 0.01% (V) (trace, not detected)
  • the aromatic content is 10 % (v), octane number (RON) of 79, a density of 721 kg / m 3 hydrotreated coker gasoline to 20,000 tons
  • the chemical light oil (distillation range 110-160 ° C) is withdrawn through the side line of the distillation column 2, and the total steaming amount is 41,000 tons/year; then, the hydrogenation unit is introduced at a flow rate of 41,000 tons/year.
  • Hydrotreating, the catalyst in the hydrogenation unit 3-1 is the entire hydrogenation catalyst GHT-22; the volumetric space velocity ratio of the hydrogenation unit 3-1 is SJh; the hydrogen/oil volume ratio is 300;
  • the temperature is 285 ° C, and the operating pressure is 2.5 MPa (absolute); the hydrogenated chemical light oil is produced as a high-quality vinyl or reforming material at a flow rate of 41,000 tons/year.
  • the extracted oil in the extraction tower 1 enters the hydrogenation unit 3-2 at a flow rate of 18,000 tons/year for hydrotreating, and the catalyst in the hydrogenation unit 3-2 is the entire hydrogenation catalyst GHT-22;
  • the hydrogenation device 3-2 has a volumetric space velocity ratio of 2.51 ⁇ ; a hydrogen/oil volume ratio of 300; an operating temperature of 270° C., an operating pressure of 2.5 MPa (absolute); and a hydrogenated oil extraction of 18,000
  • the flow rate per ton/year is mixed with the light gasoline and then produced as blended gasoline.
  • the diesel oil cut by the distillation column 2 (distillation range 160-205 ° C) was directly produced as a diesel product of 0.8 million tons/year.
  • the obtained blended gasoline has a distillation range of 30-205 ° C, a sulfur content of 3.2 ppm, a mercaptan content of less than 1.0 ppm (a trace amount, not detected), an olefin content of 19.5% (v), and a diene content of 0.05. % (v);
  • the aromatic content is 25.0% (v)
  • the octane number (RON) is 94.7
  • the density 723.3 kg/ m3
  • the recovery is 51,000 tons/year.
  • the obtained diesel oil has a distillation range of 160-205 ° C, a sulfur content of 478.1 ppm, a mercaptan content of 2.61 ppm, an olefin content of 17.8% (v), a diene content of 0.04% (v), and an aromatic content of 5.6% ( v), cetane number is 45.1, density is 759.1 kg/ m3 , and production volume is 0.8 million tons/year.
  • FIG. 3 is a schematic flowchart of Embodiment 4 of the present invention.
  • 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)
  • the aromatic content is 15% (v).
  • Stabilized gasoline having an octane number (RON) of 89 and a density of 728 kg/ m3 was extracted and separated in a stripping tower 1 at a flow rate of 60,000 tons/year to separate the extracted oil and the raffinate oil;
  • the solvent used in the extraction column 1 is sulfolane, the extraction temperature is 120 ° C, the solvent ratio (solvent / feed) is 3.5 (mass), the raffinate washing ratio is 0.2 (mass), and the solvent recovery temperature is 165 °. C, solvent recovery pressure is O.
  • the raffinate oil of the extraction tower 1 enters the distillation column 2 at a flow rate of 49,200 tons/year for cutting and fractionation, and the temperature at the top of the distillation column 2 is 77 ° C, and the temperature at the bottom of the column is At 173 ° C, the top pressure is 0.15 MPa (absolute), and the bottom pressure is 0.20 MPa (absolute), and light gasoline and heavy gasoline are obtained respectively.
  • the light gasoline (distillation range 30-110 ° C) was distilled off through the upper portion of the distillation column 2, and was produced as a blended gasoline at a flow rate of 22,100 tons/year.
  • the heavy gasoline enters the heavy gasoline hydrogenation unit 3-1 at a flow rate of 27,100 tons/year for hydrotreating, and the catalyst in the heavy gasoline hydrogenation unit 3-1 is all Hydrogenation catalyst GHT-22; the volumetric space velocity ratio of the heavy gasoline hydrogenation unit 3-1 is Oh; the hydrogen/oil volume ratio is 500; the operating temperature is 320 ° C, and the operating pressure is 4.0 MPa (absolute);
  • the extracted oil separated from the extraction column 1 is subjected to hydrotreating at a flow rate of 10,800 tons/year into the extracted oil hydrogenation unit 3-2, and the catalyst in the extracted oil hydrogenation unit 3-2 is the entire hydrogenation catalyst.
  • the obtained blended gasoline has a distillation range of 30-205, a sulfur content of 12.3 ppm, a mercaptan content of less than 1.0 ppm (a trace amount, not detected), an olefin content of 24.3% (v), and a diene content of less than 0.01. % (v) (trace, not detected); aromatics content 24.3% (v), octane number (RON) 98.4, density 727.5 kg/ m3 , and production volume 32,900 tons/year.
  • the distillation range of the obtained high-quality ethylene or reforming material is 110-205 ° C, the sulfur content is trace amount, the detection is not found, the mercaptan content is less than 1.0 ppm (; trace amount, not detected;), the olefin is trace
  • the amount, which could not be detected, has a bromine index (bromine number) of 39 (0.039), an aromatic content of 3.7% (v), an octane number (RON) of 75.0, a density of 728.7 kg/m3, and a yield of 27,100. Tons per year.
  • FIG. 4 it is a schematic flowchart of Embodiment 5 of the present invention.
  • 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).
  • Stable gasoline catalytic gasoline
  • RON octane number
  • sulfur content is 200ppm, mercaptan content is lppm, olefin content is less than 0.1% (v) (trace, not detected), diolefin content is less than 0.01% (v) (; trace, not detected; ), the aromatics content is 8% (v), the octane number (RON) is 82, and the naphtha having a density of 732 kg/ m3 is extracted and separated in the extraction tower 1 at a flow rate of 20,000 tons/year;
  • the distillation range is 30-205 ° C
  • the sulfur content is 150 ppm
  • the mercaptan content is 1 ppm
  • the olefin content is 6% (v)
  • the diene content is less than 0.01% (v) (trace, not detected)
  • the aromatic hydrocarbon content is 10% (v)
  • the octane number (RON) is 79, and the hydrogenated coking gasoline having a density of 721 kg/ m3 is extracted and separated
  • the solvent used in the extraction tower 1 is N Monomethylpyrrolidone
  • extraction temperature is 130 ° C
  • solvent ratio (solvent / feed) is 2.5 (mass)
  • raffinate water wash ratio is 0.25 (mass)
  • solvent recovery temperature is 177 ° C
  • solvent recovery pressure is 0.15 MPa (absolute)
  • the raffinate oil separated by the extraction column 1 enters the distillation column 2 at a flow rate of 85,000 tons/year for cutting and fractionation, and the temperature of the top of the distillation column 2 is 87 ° C, and the temperature at the bottom of the column is At 187 ° C, the top pressure is 0.20 MPa (absolute), and the bottom pressure is 0.25 MPa (absolute), and light gasoline and heavy gasoline are obtained respectively.
  • the light gasoline (distillation range of 30-110 ° C) was distilled off through the upper portion of the distillation column 2, and was produced as a blended gasoline at a flow rate of 37,400 tons/year.
  • the heavy gasoline (distillation range 110-170 ° C) enters the heavy gasoline hydrogenation unit 3-1 at a flow rate of 47,600 tons/year for hydrotreating, and the catalyst in the heavy gasoline hydrogenation unit 3-1 is all Hydrogenation catalyst GHT-22; the volumetric airspeed ratio of the heavy gasoline hydrogenation unit 3-1 is 1.0 Oh; the hydrogen/oil volume ratio is 250; the operating temperature is 250V, and the operating pressure is 1.0 MPa (absolute);
  • the heavy gasoline after hydrogenation is produced as a high-quality vinyl or reforming material at a flow rate of 47,600 tons/year.
  • the extracted oil separated in the extraction column 1 is subjected to hydrotreating at a flow rate of 15,000 tons/year into the extracted oil hydrogenation unit 3-2, and the catalyst in the extracted oil hydrogenation unit 3-2 is all added.
  • the extracted oil after hydrogen was mixed with the light gasoline at a flow rate of 15,000 tons/year and then produced as a blended gasoline.
  • the resulting blended gasoline has a distillation range of 30-205, a sulfur content of 24.0 ppm, a mercaptan content of less than 1.0 ppm (a trace amount, not detected), an olefin content of 16.4% (v), and a diene content of less than 0.01. % (v) (trace, not detected); aromatics content 22.3% (v), octane number (RON) 93.1, density 722.9 kg/ m3 , and production volume 52.40 million/year.
  • the obtained high-quality ethylene material or reforming material has a distillation range of 110-205 ° C, a sulfur content of 4.7 ppm, a mercaptan content of less than 1.0 ppm (a trace amount, which is not detected;), and a trace amount of olefin content, which is not detected.
  • the index (bromine number) is 32 (0.032), the aromatic content is 1.8% (v), the octane number (RON) is 73.9, the density is 732.4 kg/ m3 , and the recovery is 47,600 tons/year.
  • Embodiment 6 of the present invention is the same as Embodiment 5.
  • the distillation range is 30-205 ° C, the sulfur content is 800 ppm, the mercaptan content is 9 ppm, the olefin content is 36% (v), the diolefin content is 0.9% (v), and the aromatic content is 17% (v).
  • Stable gasoline catalytic gasoline with an octane number (RON) of 91 and a density of 731 kg/ m3 is extracted and extracted in the extraction column 1 at a flow rate of 60,000 tons/year; at the same time, the distillation range is 30-205.
  • sulfur content is 200ppm
  • mercaptan content is lppm
  • olefin content is less than 0.1% (v) (trace, not detected)
  • diolefin content is less than 0.01% (v) (; trace, not detected; )
  • the aromatics content is 8% (v)
  • the octane number (RON) is 82
  • the naphtha having a density of 732 kg/ m3 is extracted and separated in the extraction tower 1 at a flow rate of 20,000 tons/year
  • the distillation range is 30-205 ° C
  • the sulfur content is 150 ppm
  • the mercaptan content is 1 ppm
  • the olefin contains The amount is 6% (v), the diene content is less than 0.01% (v) (trace, not detected), the aromatic content is 10% (v), the octane number (RON) is 79, and the density is 721 kg/m.
  • the hydrogenated coking gasoline of 3 is subjected to extraction and separation in the extracting tower 1 at a flow rate of 20,000 tons/year to separate the extracted oil and the raffinate oil;
  • the solvent used in the extracting column 1 is N-formylmorpholine,
  • the extraction temperature is 150 ° C
  • the solvent ratio (solvent / feed) is 6.0 (mass)
  • the raffinate water wash ratio is 0.3 (mass)
  • the solvent recovery temperature is 185 ° C
  • the solvent recovery pressure is 0.2 MPa (absolute)
  • the raffinate oil of the extraction column 1 enters the distillation column 2 at a flow rate of 84,000 tons/year for cutting and fractionation, the temperature of the top of the distillation column 2 is 95 ° C, the temperature at the bottom of the column is 194 ° C, and the pressure at the top of the column It is 0.25 MPa (absolute), and the bottom pressure is 0.30 MPa (absolute), and light gasoline and heavy gasoline are respectively obtained.
  • the light gasoline (distillation range 30-110 ° C) was distilled off through the upper portion of the distillation column 2, and was produced as a blended gasoline at a flow rate of 38,600 tons/year.
  • the heavy gasoline (distillation range 110-205 ° C) enters the heavy gasoline hydrogenation unit 3-1 at a flow rate of 45,400 tons/year for hydrotreating, and the catalyst in the heavy gasoline hydrogenation unit 3-1 is all Hydrogenation catalyst GHT-22; the volumetric space velocity ratio of the heavy gasoline hydrogenation unit 3-1 is SJh- 1 ; hydrogen/oil volume ratio is 300; operating temperature is 285 ° C, operating pressure is 2.5 MPa (absolute)
  • the hydrogenated heavy gasoline is produced as a high-quality vinyl or reformate at a flow rate of 45,400 tons/year.
  • the obtained blended gasoline has a distillation range of 30-205, a sulfur content of 38.1 ppm, a mercaptan content of less than 1.0 ppm (a trace amount, not detected), an olefin content of 20.3% (v), and a diene content of less than 0.01. % (v) (trace, not detected); aromatics content 20.9% (v), octane number (RON) 93.7, density 721.3 kg/ m3 , and production volume 54.46 million tons/year.
  • the obtained high-quality ethylene or reforming material has a distillation range of 110-205 ° C, a sulfur content of 4.0 ppm, a mercaptan content of less than 1.0 ppm (trace, not detected), an olefin content of 2.0% (v), an index ( The bromine number is 25 (0.025), and the aromatic content is trace. It is not detected, the octane number (RON) is 74.8, the density is 738.9 kg/ m3 , and the recovery is 45,400 tons/year.
  • the system and method of the invention can not only de-olefin, but also desulfurized alcohol, diene; the hydrogenation device of the invention has small scale and low cost; finally, the raw materials processed by the invention are diversified, not only stable gasoline but also stable in treatment.
  • the invention combines extraction, distillation cutting and hydrogenation treatment to optimize the quality of the ethylene material, increase the extracted oil in the gasoline, and increase the octane number.

Description

催化烃軍组后加氢制备高质暈汽油的系统和方法 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, 特别涉及一种组分炼油 烃重组后加氢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, 可以分馏出干气、 液化气、 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 ΠΙ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降烯烃转变为降硫的问 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。
用碱洗脱硫不能脱烯烃, 而且还会造成环境污染; 通过对轻汽油加氢处理, 不 仅能耗高, 费用也高。
因此, 提供一种低成本、 低能耗、 无污染制备低硫含量、 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亟需解决的技术难 题。
发明内容
本发明的目的之一是提供一种低成本、 低能耗、 无污染制备低硫含量、 低烯烃 含量并且提高辛烷值 (RON) 的汽油的系统。
为实现上述目的, 本发明采取以下技术方案:
方案之一:
一种组分炼油烃重组后加氢制备高质量汽油的系统, 其特征在于: 包括抽提系 统, 蒸馏系统和加氢装置, 所述抽提系统上部通过管线与所述蒸馏系统相连接; 所 述抽提系统下部通过管线与所述加氢装置相连接,所述加氢装置通过管线与所述蒸 馏系统上部的管线相连接; 所述蒸馏系统上部通过管线直接采出产品, 所述蒸馏系 统中部通过管线与另一加氢装置相连接; 所述蒸馏系统下部通过管线直接采出产
P
方案之二:
一种组分炼油烃重组后加氢制备高质量汽油的系统, 其特征在于: 包括抽提系 统, 蒸馏系统和加氢装置, 所述抽提系统上部通过管线与所述蒸馏系统相连接; 所 述抽提系统下部通过管线与抽出油加氢装置相连接;所述蒸馏系统上部通过管线采 出轻汽油; 所述蒸馏系统下部通过管线与重汽油加氢装置相连接, 该重汽油加氢装 置下部通过管线采出重整料或乙烯料。
本发明的另一目的是提供上述制备高质量汽油的方法。
方案之一:
一种组分炼油烃重组后加氢制备高质量汽油的方法, 其步骤如下: 原料进入抽 提系统进行萃取分离, 分离出抽余油和抽出油; 所述抽提系统的抽余油进入蒸馏系 统进行切割分馏, 所述蒸馏系统上部蒸出轻汽油, 并作为调和汽油采出; 所述蒸馏 系统中部侧线抽出化工轻油; 所述化工轻油进入化工轻油加氢装置进行加氢处理; 所述加氢后的化工轻油作为优质乙烯料或重整料采出;所述抽提系统中的抽出油进 入抽出油加氢装置进行加氢处理; 加氢后的抽出油与所述轻汽油混合后作为调和汽 油采出; 所述蒸馏系统切割后的柴油直接采出。
一种优选方案, 其特征在于: 所述蒸馏系统为蒸馏塔, 塔顶温度为 77〜95°C, 塔底温度为 173〜194°C ; 所述蒸馏塔的塔顶压力为 0. 15〜0. 25MPa (绝), 塔底压力 为 0. 20〜0. 30MPa (绝); 所述轻汽油的馏程控制在 30°C〜110°C; 所述化工轻油的 馏程为 110〜160°C ; 所述柴油的馏程为 160〜205°C。
一种优选方案, 其特征在于: 所述蒸馏系统为蒸馏塔, 塔顶温度为 87°C, 塔底 温度为 184°C ; 所述蒸馏塔的塔顶压力为 0. 2MPa (绝), 塔底压力为 0. 25MPa (绝)。
一种优选方案, 其特征在于: 所述抽提系统中所用溶剂为环丁砜, 萃取温度为 120°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 165°C, 溶剂回收压力为 O. lMPa (绝)。
一种优选方案, 其特征在于: 所述抽提系统中所用溶剂为 N—甲基吡咯烷酮, 萃取温度为 130°C,溶剂比(溶剂 /进料)为 2.5 (;质量),抽余油水洗比为 0.25 (;质量;), 溶剂回收温度为 177°C, 溶剂回收压力为 0.15MPa (绝)。
一种优选方案, 其特征在于: 所述抽提系统中所用溶剂为 N—甲酰基吗啉, 萃 取温度为 150°C, 溶剂比 (溶剂 /进料) 为 6(质量), 抽余油水洗比为 0.3(质量), 溶 剂回收温度为 185°C, 溶剂回收压力为 0.2MPa (绝)。
一种优选方案, 其特征在于: 所述化工轻油加氢装置中的催化剂为全部加氢催 化剂 GHT-22 ; 所述化工轻油加氢装置的体积空速比为 l^h ; 氢 /油体积比为 250-500; 操作温度为 250~320°C, 操作压力为 l~4MPa (绝)。
一种优选方案, 其特征在于: 所述化工轻油加氢装置的体积空速比为 2.5h— 氢 /油体积比为 300; 操作温度为 285°C, 操作压力为 2.5MPa (绝)。
一种优选方案, 其特征在于: 所述抽出油加氢装置中的催化剂为全部加氢催化 剂 GHT-22, 所述抽出油加氢装置的体积空速比为 l〜4h— 氢 /油体积比为 250〜500; 操作温度为 250 290 °C, 操作压力为 l~4MPa (绝)。
—种优选方案, 其特征在于: 所述抽出油加氢装置的体积空速比为 2.51^ ; 氢 / 油体积比为 300; 操作温度为 270°C, 操作压力为 2.5MPa (绝)。
一种优选方案, 其特征在于: 所述化工轻油和抽出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 imgf000005_0001
堆密度 g/ml 0.73
比表面积 m2/g 180
孔容 ml/g 0.5-0.6
wo3 m% 15
NiO m% 1.7
CoO m% 0.15
Na20 m% <0.09
Fe203 m% <0.06
Si02 m% <0.60
载体 m% 82.4
方案之二:
一种组分炼油烃重组后加氢制备高质量汽油的方法, 其步骤如下: 原料进入抽 提系统进行萃取分离, 分离出抽余油和抽出油; 所述抽提系统分离出的抽余油进入 蒸馏系统进行切割分馏, 所述蒸馏系统上部蒸出轻汽油, 并作为调和汽油采出; 所 述蒸馏系统下部蒸出重汽油, 并进入重汽油加氢装置进行加氢处理; 所述重汽油加 氢处理后的重汽油作为乙烯料或重整料采出;所述抽提系统分离出的抽出油进入抽 出油加氢装置进行加氢处理;加氢后的抽出油与所述轻汽油混合后作为调和汽油采 出。
一种优选方案, 其特征在于: 所述蒸馏系统为蒸馏塔, 塔顶温度为 77〜95°C, 塔底温度为 173~194°C ; 塔顶压力为 0. 15〜0. 25MPa (绝), 塔底压力为 0. 20〜 0. 30MPa (绝); 所述轻汽油的馏程控制在 30°C〜110°C ; 所述重汽油的馏程为 110〜 205 °C。
一种优选方案, 其特征在于: 所述蒸馏塔的塔顶温度为 87°C, 塔底温度为 187 °C ; 塔顶压力为 0. 20MPa (绝), 塔底压力为 0. 25MPa (绝)。
一种优选方案, 其特征在于: 所述抽提系统中所用溶剂为环丁砜, 萃取温度为
120°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 165°C, 溶剂回收压力为 O. lMPa (绝)。
一种优选方案, 其特征在于: 所述抽提系统中所用溶剂为 N—甲基吡咯烷酮, 萃取温度为 130°C, 溶剂比(溶剂 /进料)为 2.5(质量), 抽余油水洗比为 0.25(质量), 溶剂回收温度为 177°C, 溶剂回收压力为 0.15MPa (绝)。
一种优选方案, 其特征在于: 所述抽提系统中所用溶剂为 N—甲酰基吗啉, 萃 取温度为 150°C, 溶剂比 (溶剂 /进料) 为 6.0(质量), 抽余油水洗比为 0.3(质量), 溶剂回收温度为 185°C, 溶剂回收压力为 0.2MPa (绝)。 一种优选方案, 其特征在于: 所述重汽油加氢装置和所述抽出油加氢装置中的 催化剂为全部加氢催化剂 GHT-22;体积空速比为 1.0〜4. Oh"1 ;氢 /油体积比为 250〜 500; 操作温度为 250〜290°C, 操作压力为 1.0〜4· OMPa (绝)。
一种优选方案, 其特征在于: 所述重汽油加氢装置的体积空速比为 2. 5^ ; 氢 /油体积比为 300; 操作温度为 270°C, 操作压力为 2. 50MPa (绝)。
一种优选方案, 其特征在于: 所述抽出油加氢装置的体积空速比为 2. 5^ ; 氢 /油体积比为 300; 操作温度为 285°C, 操作压力为 2. 50MPa (绝)。
一种优选方案, 其特征在于: 所述重汽油加氢装置和所述抽出油加氢装置中的 全部加氢催化剂 GHT-22的理化性质如下表所示。
Figure imgf000007_0001
本发明的溶剂也可以是其它溶剂,或者这些溶剂中两种或两种以上溶剂的任意 比例的混合。
本发明的石脑油、 稳定汽油和加氢焦化汽油可以是任意比例。
本发明的轻汽油、 化工轻油以及柴油的切割点 (馏程) 可以调整。 如轻汽油的 馏程控制在 30°C〜70°C ; 所述化工轻油的馏程为 70〜160°C ; 所述柴油的馏程为 160〜205°C ; 轻汽油的馏程控制在 30°C~90°C ; 所述化工轻油的馏程为 90〜160°C; 所述柴油的馏程为 160〜205 °C。
本发明所用蒸馏系统(蒸馏塔)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。 所用的抽提系统 (抽提塔) 为专利号为 200310103541.9和 200310103540.4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的流程示意图。
图 4为本发明实施例 5和 6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 1的流程示意图。 将馏程为 35-205°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% (v), 芳烃 含量为 15% ( v), 辛烷值 (RON) 为 89, 密度为 728千克 /米 3的稳定汽油 (催化 汽油) 以 6万吨 /年的流量在抽提塔 1 中进行萃取分离, 分离出抽出油和抽余油; 所述抽提塔 1 中所用溶剂为环丁砜, 萃取温度为 120°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 165°C, 溶剂回收压力为 O. lMPa (绝); 所述抽提塔 1上部出来的抽余油以 4.8万吨 /年的流量进入蒸馏塔 2 中进行切割分馏, 蒸馏塔 2 的塔顶温度为 77°C, 塔底温度为 173 °C, 塔顶压力为 0.15MPa (绝), 塔底压力为 0.20MPa (绝), 分别得到轻汽油, 化工轻油和柴油。 所述轻汽油 (馏程 30-110°C ) 通过蒸馏塔 2上部蒸出, 以 1.95万吨 /年的流量作为 调和汽油采出。 所述化工轻油 (馏程 110-160°C ) 通过蒸馏塔 2侧线抽出, 其总的 蒸出量为 2.25万吨 /年;然后以 2.25万吨 /年的流量进入加氢装置 3-1进行加氢处理, 所述加氢装置 3-1中的催化剂为全部加氢催化剂 GHT-22;所述加氢装置 3-1的体积 空速比为 41^ ; 氢 /油体积比为 500; 操作温度为 320°C, 操作压力为 4MPa (绝); 所述加氢后的化工轻油以 2.25万吨 /年的流量作为优质乙烯料或重整料采出。 所述 抽提塔 1下部出来的抽出油以 1.20万吨 /年的流量进入加氢装置 3-2进行加氢处理, 所述加氢装置 3-2中的催化剂为全部加氢催化剂 GHT-22;所述加氢装置 3-2的体积 空速比为 41^ ; 氢 /油体积比为 500; 操作温度为 290°C, 操作压力为 4MPa (绝); 加氢后的抽出油以 1.20万吨 /年的流量与所述轻汽油混合后作为调和汽油采出。 通 过蒸馏塔 2切割后的柴油(馏程 160-205°C )以 0.60万吨 /年作为柴油产品直接采出。
所得调和汽油的馏程为 30-205°C, 含硫量为 3.1ppm, 硫醇含量小于 l.Oppm (;痕 量, 检测不出), 烯烃含量为 24.1% (v), 二烯烃含量 0.05% (v); 芳烃含量为 26.7% (v), 辛烷值 (RON) 为 95.3, 密度为 727.6千克 /米 3, 采出量为 3.15万吨 /年。 所得优质乙烯料或重整料的馏程为 110-160°C, 含硫量为痕量, 检测不出, 硫 醇含量小于 l.Oppm (;痕量, 检测不出;), 烯烃含量为痕量, 检测不出, 溴指数(溴价) 为 39 (0.039), 芳烃含量为 1.3% (v), 辛烷值 (RON) 为 83.0, 密度为 729.0千克 /米 3, 采出量为 2.25万吨 /年。
所得柴油的馏程为 160-205°C, 含硫量为 26.3ppm, 硫醇含量 1.55ppm, 烯烃含 量为 27.8% (v), 二烯烃含量 0.04% (v), 芳烃含量为 5.6% (v), 十六烷值为 45.8, 密度为 751.7千克 /米 3, 采出量为 0.6万吨 /年。
所述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 imgf000009_0001
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(二烯烃)的测定: 滴定法;
9、 加氢催化剂分析方法:
Figure imgf000010_0001
10、 溴指数检测: GB/T 11136— 1989。
实施例 2
如图 2所示, 为本发明实施例 2的流程示意图。 将馏程为 35-205°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% (v), 芳烃 含量为 15% (v), 辛烷值 (RON) 为 89, 密度为 728千克 /米 3的稳定汽油 (催化 汽油) 以 6万吨 /年的流量在抽提塔 1中进行萃取分离; 同时, 将馏程为 30-205 °C, 含硫量为 200ppm, 硫醇含量为 lppm, 烯烃含量小于 0.1% (v) (痕量, 检测不出), 二烯烃含量小于 0.01% (v) (;痕量,检测不出;), 芳烃含量为 8% (v), 辛烷值(RON) 为 82, 密度为 732千克 /米 3的石脑油以 2万吨 /年的流量在抽提塔 1中进行萃取分 离; 与此同时, 将馏程为 30-205°C, 含硫量为 150ppm, 硫醇含量为 lppm, 烯烃含 量为 6% (v), 二烯烃含量小于 0.01% (v) (痕量,检测不出), 芳烃含量为 10% (v), 辛烷值 (RON) 为 79, 密度为 721千克 /米 3的加氢焦化汽油以 2万吨 /年的流量在 抽提塔 1中进行萃取分离; 三个原料在抽提塔 1中混合后进行萃取分离, 分离出抽 出油和抽余油;所述抽提塔 1中所用溶剂为 N—甲基吡咯烷酮,萃取温度为 130°C, 溶剂比 (溶剂 /进料) 为 2.5(质量;), 抽余油水洗比为 0.25(质量;), 溶剂回收温度为 11TC , 溶剂回收压力为 0.15MPa (绝); 所述抽提塔 1的抽余油以 8.4万吨 /年的流 量进入蒸馏塔 2中进行切割分馏,蒸馏塔 2的塔顶温度为 87°C,塔底温度为 184 °C, 塔顶压力为 0.2MPa (绝), 塔底压力为 0.25MPa (绝), 分别得到轻汽油, 化工汽 油和柴油。 所述轻汽油 (馏程 30-110°C )通过蒸馏塔 2上部蒸出, 以 3.3万吨 /年的 流量作为调和汽油采出。 所述化工轻油 (馏程 110-160°C )通过蒸馏塔 2侧线抽出, 其总的蒸出量为 4.1万吨 /年; 然后以 4.1万吨 /年的流量进入加氢装置 3-1进行加氢 处理, 所述加氢装置 3-1中的催化剂为全部加氢催化剂 GHT-22; 所述加氢装置 3-1 的体积空速比为 l.Oh ;氢 /油体积比为 250;操作温度为 250V,操作压力为 l.OMPa (绝); 所述化工轻油以 4.1万吨 /年的流量作为优质乙烯料或重整料采出。 所述抽 提塔 1中的抽出油以 1.6万吨 /年的流量进入加氢装置 3-2进行加氢处理, 所述加氢 装置 3-2中的催化剂为全部加氢催化剂 GHT-22;所述加氢装置 3-2的体积空速比为 l.Oh"1; 氢 /油体积比为 250; 操作温度为 250°C, 操作压力为 l.OMPa (绝); 加氢后 的抽出油以 1.6万吨 /年的流量与所述轻汽油混合后作为调和汽油采出。通过蒸馏塔 2切割后的柴油 (馏程 160-205°C) 以 1.0万吨 /年作为柴油产品直接采出。
所得调和汽油的馏程为 30-205°C, 含硫量为 3.4ppm, 硫醇含量小于 l.Oppm (;痕 量, 检测不出), 烯烃含量为 16.6% (v), 二烯烃含量 0.05% (v); 芳烃含量为 23.7% (v), 辛烷值 (RON) 为 95.0, 密度为 719.2千克 /米 3, 采出量为 4.9万吨 /年。
所得优质乙烯料或重整料的馏程为 110-160°C, 含硫量为痕量, 检测不出, 硫 醇含量小于 l.Oppm (;痕量, 检测不出;), 烯烃含量为痕量, 检测不出, 溴指数(溴价) 为 32 (0.032), 芳烃含量为 1.3% (v), 辛烷值(RON) 为 76.2, 密度为 731.0千克 /米 3, 采出量为 4.1万吨 /年。
所得柴油的馏程为 160-205°C, 含硫量为 46.5ppm, 硫醇含量为 0.17ppm (;痕量, 检测不出), 烯烃含量为 15.0% (v), 二烯烃含量为 0.04% (v), 芳烃含量为 3.6% (v), 十六烷值为 45.5, 密度为 753.9千克 /米 3, 采出量为 1.0万吨 /年。
所述加氢装置中的全部加氢催化剂 GHT-22的理化性质以及所用测试方法均与 实施例 1相同。
实施例 3
流程同实施例 2。 将馏程为 30-205°C, 含硫量为 800ppm, 硫醇含量为 9ppm, 烯烃含量为 36%(v),二烯烃含量为 0.9%(v),芳烃含量为 17%(v),辛烷值(RON) 为 91, 密度为 731千克 /米 3的稳定汽油 (催化汽油) 以 6万吨 /年的流量在抽提塔 1中进行萃取分离;同时,将馏程为 30-205 ,含硫量为 200ppm,硫醇含量为 lppm, 烯烃含量小于 0.1% (v) (痕量, 检测不出), 二烯烃含量小于 0.01% (v) (痕量, 检 测不出), 芳烃含量为 8% (v), 辛烷值 (RON) 为 82, 密度为 732千克 /米 3的石 脑油以 2万吨 /年的流量在抽提塔 1中进行萃取分离;与此同时,将馏程为 30-205 , 含硫量为 150ppm,硫醇含量为 lppm, 烯烃含量为 6% (v), 二烯烃含量小于 0.01% (V) (痕量, 检测不出), 芳烃含量为 10% (v), 辛烷值(RON) 为 79, 密度为 721 千克 /米 3的加氢焦化汽油以 2万吨 /年的流量在抽提塔 1 中进行萃取分离; 三种原 料在抽提塔 1中混合后萃取分离, 分离出抽出油和抽余油; 所述抽提塔 1中所用溶 剂为 N—甲酰基吗啉, 萃取温度为 150°C, 溶剂比 (溶剂 /进料) 为 6.0(质量), 抽余 油水洗比为 0.3(质量), 溶剂回收温度为 185°C, 溶剂回收压力为 0.2MPa (绝); 所 述抽提塔 1上部出来的抽余油以 8.2万吨 /年的流量进入蒸馏塔 2中进行切割分馏, 蒸馏塔 2的塔顶温度为 95°C, 塔底温度为 194°C, 塔顶压力为 0.25MPa (绝), 塔 底压力为 0.30MPa (绝), 分别得到轻汽油, 化工汽油和柴油。 所述轻汽油 (馏程 30-110°C ) 通过蒸馏塔 2上部蒸出, 以 3.3万吨 /年的流量作为调和汽油采出。 所述 化工轻油(馏程 110-160°C )通过蒸馏塔 2侧线抽出, 其总的蒸出量为 4.1万吨 /年; 然后以 4.1万吨 /年的流量进入加氢装置 3-1进行加氢处理, 所述加氢装置 3-1中的 催化剂为全部加氢催化剂 GHT-22; 所述加氢装置 3-1 的体积空速比为 SJh ; 氢 / 油体积比为 300; 操作温度为 285°C, 操作压力为 2.5MPa (绝); 所述加氢后的化 工轻油以 4.1万吨 /年的流量作为优质乙烯料或重整料采出。所述抽提塔 1中的抽出 油以 1.8万吨 /年的流量进入加氢装置 3-2进行加氢处理, 所述加氢装置 3-2中的催 化剂为全部加氢催化剂 GHT-22;所述加氢装置 3-2的体积空速比为 2.51^ ;氢 /油体 积比为 300; 操作温度为 270°C, 操作压力为 2.5MPa (绝); 加氢后的抽出油以 1.8 万吨 /年的流量与所述轻汽油混合后作为调和汽油采出。 通过蒸馏塔 2 切割后的柴 油 (馏程 160-205°C ) 以 0.8万吨 /年作为柴油产品直接采出。
所得调和汽油的馏程为 30-205°C, 含硫量为 3.2ppm, 硫醇含量小于 l.Oppm (;痕 量, 检测不出), 烯烃含量为 19.5% (v), 二烯烃含量 0.05% (v); 芳烃含量为 25.0% (v), 辛烷值 (RON) 为 94.7, 密度为 723.3千克 /米 3, 采出量为 5.1万吨 /年。
所得优质乙烯料或重整料的馏程为 110-160°C, 含硫量为痕量, 检测不出, 硫 醇含量小于 l.Oppm (;痕量, 检测不出;), 烯烃含量为痕量, 检测不出, 溴指数(溴价) 为 25 (0.025 ), 芳烃含量为 1.3% (v), 辛烷值(RON) 为 77.6, 密度为 731.0千克 /米 3, 采出量为 4.1万吨 /年。
所得柴油的馏程为 160-205°C, 含硫量为 478.1ppm, 硫醇含量 2.61ppm, 烯烃 含量为 17.8% (v), 二烯烃含量为 0.04% (v), 芳烃含量为 5.6% (v), 十六烷值为 45.1, 密度为 759.1千克 /米 3, 采出量为 0.8万吨 /年。
所述加氢装置中的全部加氢催化剂 GHT-22的理化性质以及所用测试方法均与 实施例 1相同。
实施例 4
如图 3所示, 为本发明实施例 4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% (v), 芳烃 含量为 15% (v), 辛烷值 (RON) 为 89, 密度为 728千克 /米 3的稳定汽油 (催化 汽油) 以 6万吨 /年的流量在抽提塔 1 中进行萃取分离, 分离出抽出油和抽余油; 所述抽提塔 1 中所用溶剂为环丁砜, 萃取温度为 120°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 165°C, 溶剂回收压力为 O. lMPa (绝); 所述抽提塔 1的抽余油以 4.92万吨 /年的流量进入蒸馏塔 2中进行切 割分馏,蒸馏塔 2的塔顶温度为 77°C,塔底温度为 173°C,塔顶压力为 0.15MPa (绝), 塔底压力为 0.20MPa (绝),分别得到轻汽油和重汽油。所述轻汽油(馏程 30-110°C ) 通过蒸馏塔 2上部蒸出, 以 2.21万吨 /年的流量作为调和汽油采出。所述重汽油(馏 程 110-170°C ) 以 2.71万吨 /年的流量进入重汽油加氢装置 3-1进行加氢处理, 所述 重汽油加氢装置 3-1 中的催化剂为全部加氢催化剂 GHT-22; 所述重汽油加氢装置 3-1 的体积空速比为 Oh ; 氢 /油体积比为 500; 操作温度为 320°C, 操作压力为 4.0MPa (绝); 所述抽提塔 1分离出的抽出油以 1.08万吨 /年的流量进入抽出油加氢 装置 3-2 进行加氢处理, 所述抽出油加氢装置 3-2 中的催化剂为全部加氢催化剂 GHT-22; 所述抽出油加氢装置 3-2的体积空速比为 4.011—1 ; 氢 /油体积比为 500; 操 作温度为 290°C, 操作压力为 4.0MPa (绝); 加氢后的抽出油以 1.08万吨 /年的流量 与所述轻汽油混合后作为调和汽油采出。
所得调和汽油的馏程为 30-205 ,含硫量为 12.3ppm,硫醇含量小于 l.Oppm (;痕 量, 检测不出), 烯烃含量为 24.3% (v), 二烯烃含量低于 0.01% (v) (痕量, 检测 不出); 芳烃含量为 24.3% (v), 辛烷值 (RON) 为 94.8, 密度为 727.5千克 /米 3, 采出量为 3.29万吨 /年。
所得优质乙烯料或重整料的馏程为 110-205°C, 含硫量为痕量, 检测不出, 硫醇含 量小于 l.Oppm (;痕量, 检测不出;), 烯烃为痕量, 检测不出, 溴指数 (溴价) 为 39 (0.039), 芳烃含量为 3.7% (v), 辛烷值(RON)为 75.0, 密度为 728.7千克 /米 3, 采出量为 2.71万吨 /年。
所述各个加氢装置中的全部加氢催化剂 GHT-22的理化性质及测定方法同实施 例 1。
实施例 5
如图 4所示, 为本发明实施例 5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% (v), 芳烃 含量为 15% (v), 辛烷值 (RON) 为 89, 密度为 728千克 /米 3的稳定汽油 (催化 汽油) 以 6万吨 /年的流量在抽提塔 1中进行萃取分离; 同时, 将馏程为 30-205°C, 含硫量为 200ppm, 硫醇含量为 lppm, 烯烃含量小于 0.1% (v) (痕量, 检测不出), 二烯烃含量小于 0.01% (v) (;痕量,检测不出;), 芳烃含量为 8% (v), 辛烷值(RON) 为 82, 密度为 732千克 /米 3的石脑油以 2万吨 /年的流量在抽提塔 1中进行萃取分 离; 与此同时, 将馏程为 30-205°C, 含硫量为 150ppm, 硫醇含量为 lppm, 烯烃含 量为 6% (v), 二烯烃含量小于 0.01% (v) (痕量,检测不出), 芳烃含量为 10% (v), 辛烷值 (RON) 为 79, 密度为 721千克 /米 3的加氢焦化汽油以 2万吨 /年的流量在 抽提塔 1中进行萃取分离, 分离出抽出油和抽余油; 所述抽提塔 1中所用溶剂为 N 一甲基吡咯烷酮, 萃取温度为 130°C, 溶剂比(溶剂 /进料)为 2.5(质量), 抽余油水 洗比为 0.25(质量), 溶剂回收温度为 177°C, 溶剂回收压力为 0.15MPa (绝); 所述 抽提塔 1分离出的抽余油以 8.50万吨 /年的流量进入蒸馏塔 2中进行切割分馏, 蒸 馏塔 2的塔顶温度为 87°C, 塔底温度为 187°C, 塔顶压力为 0.20MPa (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 30-110°C) 通过蒸馏塔 2上部蒸出, 以 3.74万吨 /年的流量作为调和汽油采出。所述重汽油(馏 程 110-170°C) 以 4.76万吨 /年的流量进入重汽油加氢装置 3-1进行加氢处理, 所述 重汽油加氢装置 3-1 中的催化剂为全部加氢催化剂 GHT-22; 所述重汽油加氢装置 3-1 的体积空速比为 l.Oh ; 氢 /油体积比为 250; 操作温度为 250V, 操作压力为 l.OMPa (绝);加氢后的重汽油以 4.76万吨 /年的流量作为优质乙烯料或重整料采出。 所述抽提塔 1中分离出的抽出油以 1.50万吨 /年的流量进入抽出油加氢装置 3-2进 行加氢处理, 所述抽出油加氢装置 3-2中的催化剂为全部加氢催化剂 GHT-22; 所 述抽出油加氢装置 3-2 的体积空速比为 l.Oh— 氢 /油体积比为 250; 操作温度为 250V, 操作压力为 l.OMPa (绝); 加氢后的抽出油以 1.50万吨 /年的流量与所述轻 汽油混合后作为调和汽油采出。
所得调和汽油的馏程为 30-205 ,含硫量为 24.0ppm,硫醇含量小于 l.Oppm (;痕 量, 检测不出), 烯烃含量为 16.4% (v), 二烯烃含量低于 0.01% (v) (痕量, 检测 不出); 芳烃含量为 22.3% (v), 辛烷值 (RON) 为 93.1, 密度为 722.9千克 /米 3, 采出量为 5.24万吨 /年。
所得优质乙烯料或重整料的馏程为 110-205°C, 硫含量为 4.7ppm, 硫醇含量小 于 l.Oppm (;痕量,检测不出;),烯烃含量痕量,检测不出,指数(溴价)为 32 (0.032), 芳烃含量为 1.8% (v), 辛烷值(RON) 为 73.9, 密度为 732.4千克 /米 3, 采出量为 4.76万吨 /年。
所述各个加氢装置中的全部加氢催化剂 GHT-22的理化性质以及所用测试方法 均与实施例 1相同。
实施例 6
本发明实施例 6 的流程示意图同实施例 5。 将馏程为 30-205°C, 含硫量为 800ppm, 硫醇含量为 9ppm, 烯烃含量为 36% (v), 二烯烃含量为 0.9% (v), 芳烃 含量为 17% (v), 辛烷值 (RON) 为 91, 密度为 731千克 /米 3的稳定汽油 (催化 汽油) 以 6万吨 /年的流量在抽提塔 1中进行萃取分离; 同时, 将馏程为 30-205°C, 含硫量为 200ppm, 硫醇含量为 lppm, 烯烃含量小于 0.1% (v) (痕量, 检测不出), 二烯烃含量小于 0.01% (v) (;痕量,检测不出;), 芳烃含量为 8% (v), 辛烷值(RON) 为 82, 密度为 732千克 /米 3的石脑油以 2万吨 /年的流量在抽提塔 1中进行萃取分 离; 与此同时, 将馏程为 30-205°C, 含硫量为 150ppm, 硫醇含量为 lppm, 烯烃含 量为 6% (v), 二烯烃含量小于 0.01% (v) (痕量,检测不出), 芳烃含量为 10% (v), 辛烷值 (RON) 为 79, 密度为 721千克 /米 3的加氢焦化汽油以 2万吨 /年的流量在 抽提塔 1中进行萃取分离, 分离出抽出油和抽余油; 所述抽提塔 1中所用溶剂为 N 一甲酰基吗啉, 萃取温度为 150°C, 溶剂比 (溶剂 /进料)为 6.0(质量), 抽余油水洗 比为 0.3(质量), 溶剂回收温度为 185°C, 溶剂回收压力为 0.2MPa (绝); 所述抽提 塔 1的抽余油以 8.40万吨 /年的流量进入蒸馏塔 2中进行切割分馏, 蒸馏塔 2的塔 顶温度为 95°C,塔底温度为 194 °C,塔顶压力为 0.25MPa(绝),塔底压力为 0.30MPa (绝), 分别得到轻汽油和重汽油。 所述轻汽油 (馏程 30-110°C ) 通过蒸馏塔 2上 部蒸出, 以 3.86万吨 /年的流量作为调和汽油采出。 所述重汽油 (馏程 110-205°C ) 以 4.54万吨 /年的流量进入重汽油加氢装置 3-1进行加氢处理, 所述重汽油加氢装 置 3-1中的催化剂为全部加氢催化剂 GHT-22;所述重汽油加氢装置 3-1的体积空速 比为 SJh—1 ; 氢 /油体积比为 300; 操作温度为 285°C, 操作压力为 2.5MPa (绝); 加 氢后的重汽油以 4.54万吨 /年的流量作为优质乙烯料或重整料采出。 所述抽提塔 1 中的抽出油以 1.60万吨 /年的流量进入抽出油加氢装置 3-2进行加氢处理, 所述抽 出油加氢装置 3-2中的催化剂为全部加氢催化剂 GHT-22; 所述抽出油加氢装置 3-2 的体积空速比为 SJh ;氢 /油体积比为 300;操作温度为 270°C,操作压力为 2.5MPa (绝);加氢后的抽出油以 1.60万吨 /年的流量与所述轻汽油混合后作为调和汽油采 出。
所得调和汽油的馏程为 30-205 ,含硫量为 38.1ppm,硫醇含量小于 l.Oppm (;痕 量, 检测不出), 烯烃含量为 20.3% (v), 二烯烃含量低于 0.01% (v) (痕量, 检测 不出); 芳烃含量为 20.9% (v), 辛烷值 (RON) 为 93.7, 密度为 721.3千克 /米 3, 采出量为 5.46万吨 /年。
所得优质乙烯料或重整料的馏程为 110-205°C, 硫含量为 4.0ppm, 硫醇含量小 于 l.Oppm (痕量, 检测不出), 烯烃含量 2.0% (v), 指数 (溴价) 为 25 (0.025 ), 芳烃含量为痕量, 检测不出, 辛烷值 (RON) 为 74.8, 密度为 738.9千克 /米 3, 采 出量为 4.54万吨 /年。
所述各个加氢装置中的全部加氢催化剂 GHT-22的理化性质以及所用测试方法 均与实施例 1相同
工业应用性
本发明的系统和方法不仅能脱烯烃, 还能脱硫醇, 双烯; 本发明的加氢装置规 模小, 成本低; 最后, 本发明处理的原料多样化, 不仅处理稳定汽油, 还可以处理 稳定汽油和石脑油以及加氢焦化汽油的混合物。 本发明将抽提分离、 蒸馏切割和加 氢处理有机结合,使优质乙烯料得到优化,调和汽油中的抽出油增加,辛烷值提高。

Claims

权 利 要 求 书
1、 一种组分炼油烃重组后加氢制备高质量汽柴油的系统, 其特征在于: 包括 抽提系统, 蒸馏系统和加氢装置, 所述抽提系统上部通过管线与所述蒸馏系统相连 接; 所述抽提系统下部通过管线与所述加氢装置相连接, 所述加氢装置通过管线与 所述蒸馏系统上部的管线相连接; 所述蒸馏系统上部通过管线直接采出产品, 所述 蒸馏系统中部通过管线与另一加氢装置相连接;所述蒸馏系统下部通过管线直接采 出产品。
2、 一种组分炼油烃重组后加氢制备高质量汽柴油的方法, 其步骤如下: 原料 进入抽提系统进行萃取分离, 分离出抽余油和抽出油; 所述抽提系统的抽余油进入 蒸馏系统进行切割分馏, 所述蒸馏系统上部蒸出轻汽油, 并作为调和汽油采出; 所 述蒸馏系统中部侧线抽出化工轻油;所述化工轻油进入化工轻油加氢装置进行加氢 处理; 所述加氢后的化工轻油作为优质乙烯料或重整料采出; 所述抽提系统中的抽 出油进入抽出油加氢装置进行加氢处理;加氢后的抽出油与所述轻汽油混合后作为 调和汽油采出; 所述蒸馏系统切割后的柴油直接采出。
3、 根据权利要求 2所述的组分炼油烃重组后加氢制备高质量汽柴油的方法, 其特征在于: 所述蒸馏系统为蒸馏塔, 塔顶温度为 77〜95°C, 塔底温度为 173〜194 。C ; 塔顶压力为 0. 15〜0. 25MPa (绝), 塔底压力为 0. 20〜0. 30MPa (绝); 所述轻 汽油的馏程控制在 30°C〜110°C ; 所述化工轻油的馏程为 110〜160°C ; 所述柴油的 馏程为 160〜205°C。
4、 根据权利要求 3所述的组分炼油烃重组后加氢制备高质量汽柴油的方法, 其特征在于: 所述蒸馏系统为蒸馏塔, 塔顶温度为 87°C, 塔底温度为 184 °C ; 塔顶 压力为 0. 2MPa (绝), 塔底压力为 0. 25MPa (绝)。
5、 根据权利要求 4所述的组分炼油烃重组后加氢制备高质量汽柴油的方法, 其特征在于: 所述抽提系统中所用溶剂为环丁砜, 萃取温度为 120°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 165°C, 溶剂 回收压力为 O. lMPa (绝)。
6、 根据权利要求 4所述的组分炼油烃重组后加氢制备高质量汽柴油的方法, 其特征在于: 所述抽提系统中所用溶剂为 N—甲基吡咯烷酮, 萃取温度为 130°C, 溶剂比 (溶剂 /进料) 为 2.5 (质量), 抽余油水洗比为 0.25 (质量), 溶剂回收温度为 177°C, 溶剂回收压力为 0.15MPa (绝)。
7、 根据权利要求 4所述的组分炼油烃重组后加氢制备高质量汽柴油的方法, 其特征在于: 所述抽提系统中所用溶剂为 N—甲酰基吗啉, 萃取温度为 150°C, 溶 剂比(溶剂 /进料)为 6(质量), 抽余油水洗比为 0.3(质量), 溶剂回收温度为 185°C, 溶剂回收压力为 0.2MPa (绝)。
8、 根据权利要求 5-7 中任一项所述的组分炼油烃重组后加氢制备高质量汽柴 油的方法, 其特征在于: 所述化工轻油加氢装置中的催化剂为全部加氢催化剂 GHT-22; 所述化工轻油加氢装置的体积空速比为 l Ah—1 ; 氢 /油体积比为 250〜500; 操作温度为 250〜320°C, 操作压力为 l〜4MPa (绝); 所述抽出油加氢装置中的催化 剂为全部加氢催化剂 GHT-22, 所述抽出油加氢装置的体积空速比为 l^h ; 氢 /油 体积比为 250 500; 操作温度为 250~320°C, 操作压力为 l~4MPa (绝)。
9、 根据权利要求 8所述的组分炼油烃重组后加氢制备高质量汽柴油的方法, 其特征在于: 所述化工轻油加氢装置的体积空速比为 SJh ; 氢 /油体积比为 300; 操作温度为 285°C, 操作压力为 2.5MPa (绝); 所述抽出油加氢装置的体积空速比 为 ^!!—1 ; 氢 /油体积比为 300; 操作温度为 285°C, 操作压力为 2.5MPa (绝)。
10、 根据权利要求 9所述的组分炼油烃重组后加氢制备高质量汽柴油的方法, 其特征在于: 所述化工轻油和抽出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_0001
11、 一种组分炼油烃重组后加氢制备高质量汽油的系统, 其特征在于: 包括抽 提系统,蒸馏系统和加氢装置,所述抽提系统上部通过管线与所述蒸馏系统相连接; 所述抽提系统下部通过管线与抽出油加氢装置相连接;所述蒸馏系统上部通过管线 采出轻汽油; 所述蒸馏系统下部通过管线与重汽油加氢装置相连接, 该重汽油加氢 装置下部通过管线采出重整料或乙烯料。
12、 一种组分炼油烃重组后加氢制备高质量汽油的方法, 其步骤如下: 原料进 入抽提系统进行萃取分离, 分离出抽余油和抽出油; 所述抽提系统分离出的抽余油 进入蒸馏系统进行切割分馏,所述蒸馏系统上部蒸出轻汽油,并作为调和汽油采出; 所述蒸馏系统下部蒸出重汽油, 并进入重汽油加氢装置进行加氢处理; 所述重汽油 加氢处理后的重汽油作为乙烯料或重整料采出;所述抽提系统分离出的抽出油进入 抽出油加氢装置进行加氢处理;加氢后的抽出油与所述轻汽油混合后作为调和汽油 采出。
13、 根据权利要求 12所述的组分炼油烃重组后加氢制备高质量汽油的方法, 其特征在于: 所述蒸馏系统的蒸馏塔的塔顶温度为 77〜95°C, 塔底温度为 173〜194 。C ; 塔顶压力为 0. 15〜0. 25MPa (绝), 塔底压力为 0. 20〜0. 30MPa (绝); 所述轻 汽油的馏程控制在 30°C〜110°C ; 所述重汽油的馏程为 110〜170°C。
14、 根据权利要求 13所述的组分炼油烃重组后加氢制备高质量汽油的方法, 其特征在于: 所述蒸馏系统的蒸馏塔的塔顶温度为 87°C, 塔底温度为 187°C ; 塔顶 压力为 0. 20MPa (绝), 塔底压力为 0. 25MPa (绝)。
15、 根据权利要求 14所述的组分炼油烃重组后加氢制备高质量汽油的方法, 其特征在于: 所述抽提系统中所用溶剂为环丁砜, 萃取温度为 120°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 165°C, 溶剂 回收压力为 O. lMPa (绝)。
16、 根据权利要求 14所述的组分炼油烃重组后加氢制备高质量汽油的方法, 其特征在于: 所述抽提系统中所用溶剂为 N—甲基吡咯烷酮, 萃取温度为 130°C, 溶剂比 (溶剂 /进料) 为 2.5(质量;), 抽余油水洗比为 0.25(质量;), 溶剂回收温度为 177°C, 溶剂回收压力为 0.15MPa (绝)。
17、 根据权利要求 14所述的组分炼油烃重组后加氢制备高质量汽油的方法, 其特征在于: 所述抽提系统中所用溶剂为 N—甲酰基吗啉, 萃取温度为 150°C, 溶 剂比(溶剂 /进料)为 6.0(质量),抽余油水洗比为 0.3(质量),溶剂回收温度为 185°C, 溶剂回收压力为 0.2MPa (绝)。
18、 根据权利要求 15— 17中任一项所述的组分炼油烃重组后加氢制备高质量 汽油的方法, 其特征在于: 所述重汽油加氢装置和所述抽出油加氢装置中的催化剂 为全部加氢催化剂 GHT-22; 体积空速比为 l.O C 1 ; 氢 /油体积比为 250〜500; 操作温度为 250〜320°C, 操作压力为 1.0〜4. 0MPa (绝)。
19、 根据权利要求 18所述的组分炼油烃重组后加氢制备高质量汽油的方法, 其特征在于: 所述重汽油加氢装置的体积空速比为 2. 51^ ; 氢 /油体积比为 300; 操 作温度为 270°C, 操作压力为 2. 50MPa (绝); 所述抽出油加氢装置的体积空速比 为 2. δΐι·1 ; 氢 /油体积比为 300; 操作温度为 285°C, 操作压力为 2. 50MPa (绝)。
20、 根据权利要求 19所述的组分炼油烃重组后加氢制备高质量汽油的方法, 其特征在于: 所述重汽油加氢装置和所述抽出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 imgf000019_0001
PCT/CN2009/070238 2009-01-21 2009-01-21 催化烃重组后加氢制备高质量汽油的系统和方法 WO2010083642A1 (zh)

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EA201071241A EA019489B1 (ru) 2009-01-21 2009-01-21 Система и способ получения высококачественного бензина гидрированием с рекомбинацией углеводородного компонента перегонки нефти
BRPI0909889A BRPI0909889A2 (pt) 2009-01-21 2009-01-21 sistema e método para preparar gasolina de alta qualidade através do refino de componente d recombinação de hidrocarbonetos de hidrogenação
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