WO2016123861A1 - 一种催化裂化汽油的提质方法 - Google Patents

一种催化裂化汽油的提质方法 Download PDF

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WO2016123861A1
WO2016123861A1 PCT/CN2015/075888 CN2015075888W WO2016123861A1 WO 2016123861 A1 WO2016123861 A1 WO 2016123861A1 CN 2015075888 W CN2015075888 W CN 2015075888W WO 2016123861 A1 WO2016123861 A1 WO 2016123861A1
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Prior art keywords
gasoline
gasoline fraction
fraction
desulfurization
heavy
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PCT/CN2015/075888
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English (en)
French (fr)
Inventor
高金森
赵亮
徐春明
郝天臻
韩晓娜
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中国石油大学(北京)
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Priority claimed from CN201510059630.0A external-priority patent/CN104667861B/zh
Priority claimed from CN201510058274.0A external-priority patent/CN104673377B/zh
Priority claimed from CN201510058454.9A external-priority patent/CN104673363B/zh
Application filed by 中国石油大学(北京) filed Critical 中国石油大学(北京)
Priority to US14/940,027 priority Critical patent/US10266778B2/en
Publication of WO2016123861A1 publication Critical patent/WO2016123861A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
    • 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

Definitions

  • the invention belongs to the technical field of petrochemical industry, and particularly relates to a method for upgrading catalytic cracked gasoline.
  • Oil resources are becoming increasingly heavy and inferior, and environmental protection requirements are increasing.
  • the world's new environmental regulations are stricter on the quality of gasoline.
  • the national V motor gasoline standard to be implemented on January 1, 2018 will require an olefin content of less than 24%, a sulfur content of less than 10 ppm, and an octane number of 93 or more.
  • the improvement in gasoline quality standards is mainly reflected in the increase in octane number while the olefin content and sulfur content are further reduced.
  • Catalytic cracking gasoline is the main component of China's motor gasoline, accounting for about 75% in the gasoline pool. About 90% of the olefin content and sulfur content in the finished gasoline come from catalytic cracking gasoline, which makes China's gasoline products far from meeting the new requirements of sulfur content ⁇ 10ppm and olefin content ⁇ 24%. On the other hand, at present, China's 93rd gasoline is the mainstay. However, with the continuous improvement of the domestic automobile manufacturing industry and the continuous increase in the number of imported automobile vehicles, the demand for gasoline of 95 or higher octane number is increasing. Catalytic cracking gasoline is limited by the process itself, and its octane number is mainly maintained by a large amount of olefins.
  • the RON is generally around 90, so the octane number directly affects the octane level of the finished gasoline.
  • the current mainstream process for catalytic cracking gasoline desulfurization and olefin reduction is catalytic hydrogenation, which inevitably brings about a large amount of olefin saturation, resulting in a large octane loss, which seriously affects the economic benefits of the enterprise.
  • the existing catalytic cracking gasoline sulfur reduction technology is mainly represented by Sinopec S-zorb, Stone Institute RSDS and French Prime-G+.
  • S-zorb is developed by Conocophillips Company of the United States.
  • Sinopec Group bought out and perfected it to desulfurize the whole fraction catalytic gasoline. After desulfurization, the sulfur content can be controlled to less than 10ppm.
  • the octane loss of the whole fraction gasoline is 1.0 to 2.0 units.
  • RSDS is developed by the petrochemical science research institute. The technology first cuts the catalytic gasoline into light and heavy fractions, the light fractions are extracted and desulfurized, and the heavy fractions are selectively hydrodesulfurized.
  • Prime-G+ uses a full-fraction pre-hydrogenation, light-weight gasoline partitioning and heavy-distillation selective hydrodesulfurization process, which is characterized by the use of light sulfides and diolefins in the full hydrogenation pre-hydrogenation process.
  • the action forms a high-boiling sulfide, the olefin is not saturated, and then cut by light heavy gasoline to obtain a light fraction and a high sulfur heavy fraction having a sulfur content of less than 10 ppm, and the heavy fraction is hydrodesulfurized; this technique is the same as RSDS, although partially low sulfur
  • the components may be unhydrogenated, but due to the low yield of light components of less than 10 ppm, most require hydrotreating, resulting in an octane loss of between 3.0 and 4.0 for the whole fraction gasoline.
  • CN1611572A discloses a catalytic conversion process for increasing the octane number of gasoline.
  • the method comprises contacting a heavy gasoline fraction having an initial boiling point greater than 100 ° C with a catalyst having a temperature lower than 700 ° C, at a temperature of 300-660 ° C, 130-450 KPa, a weight hourly space velocity of 1 to 120 h -1 , and a weight of the catalyst and the gasoline fraction.
  • the reaction is carried out under the condition that the ratio of water vapor to gasoline fraction is 0-0.1, the reaction product and the living agent are separated, and the raw agent is recycled after being stripped and regenerated.
  • the octane number of the catalytically cracked gasoline can be increased by 3 to 10 units by the method provided by the present invention.
  • the method follows the catalytic cracking mechanism of petroleum hydrocarbons, and the gasoline undergoes hydrogen transfer reaction and cracking reaction. Although the gasoline octane number can be improved, it is necessary to first perform fractional cutting and only collect heavy gasoline fraction having an initial boiling point greater than 100 ° C for reaction. The amount of loss is large.
  • CN 1160746 A discloses a catalytic conversion process for increasing the octane number of low quality gasoline.
  • the method is that the low octane gasoline is injected into the riser reactor from the upstream of the inlet of the conventional catalytic cracking raw material, and is contacted with the high temperature catalyst from the regenerator at a reaction temperature of 600-730 ° C and a ratio of the agent to the oil of 6 to 180.
  • the reaction was carried out under conditions of a weight hourly space velocity of from 1 to 180 hrs .
  • This method can increase the octane number of gasoline, but all of the low octane gasoline in this method needs to participate in the reaction, and the loss of gasoline is large.
  • CN 103805269 A proposes a method for catalytic deep hydrodesulfurization of gasoline, wherein the light gasoline and the middle gasoline fraction are subjected to alkali-free deodorization, and then the light and medium gasoline are separated by the hydrogenation pre-fractionation tower, and the hydrogenation pre-fractionation tower is simultaneously introduced into the hot diesel; The separated medium gasoline is mixed with heavy gasoline for selective hydrogenation, and the obtained distillate oil is mixed with the alkali-free deodorized light gasoline to obtain a clean gasoline product.
  • the method can effectively desulfurize, the degree of reduction of octane number is also moderated, but the octane number cannot be effectively improved, and the process flow is greatly different from the present invention.
  • the existing technology for reducing the sulfur content of catalytic cracking gasoline generally has a problem of large hydrotreating ratio and high octane loss when dealing with deep desulfurization requirements.
  • the octane recovery process of some supporting hydrodesulfurization processes is also ineffective.
  • Degree desulfurization technology There is an urgent need in the market to develop catalytically cracked gasoline deep with less octane loss or significantly increased octane number.
  • the invention solves the above technical problem and provides a method for upgrading catalytic cracked gasoline, which can not only deeply remove the sulfide content of the catalytic cracking gasoline below 10 ppm, but also can significantly increase the octane number of the catalytic cracking gasoline by 3-5. unit.
  • a method for upgrading catalytic cracked gasoline comprising the following steps:
  • the cutting temperature of the light and medium gasoline fraction is 35-60 ° C
  • the cutting temperature of the medium and heavy gasoline fraction is 70-160 ° C.
  • the cutting according to the present invention divides the catalytically cracked gasoline into three fractions of light, medium and heavy according to the boiling range from low to high, and controls the boiling range of the middle gasoline fraction to be 35-50 ° C to 130-160 ° C. .
  • the catalytic cracking gasoline (FCC gasoline) is subjected to fraction cutting, and the collected light gasoline fraction is controlled to be classified into a catalytically cracked gasoline rich in olefin and high in octane number by controlling the cutting temperature, and the gasoline fraction is distilled.
  • Catalytic cracking gasoline with medium olefin and aromatic content and lowest octane number heavy gasoline fraction is classified as catalytic cracking gasoline with lower olefin content, higher aromatic content and higher octane number.
  • the present invention performs aromatization/hydroisomerization reaction on the middle gasoline fraction having the lowest octane number, and further mixes the reaction product with other gasoline fractions to obtain a FCC gasoline having a significantly improved octane number.
  • the range and range of the raw gasoline can be comprehensively considered to determine the distillation range of the middle gasoline fraction.
  • the inventors' research found that the middle gasoline fraction with a distillation range of 40-160 ° C accounts for about 40 m% of the catalytic cracked gasoline, which is basically the lowest octane fraction, and its RON is below 80, and even a small part is below 70. Therefore, the middle gasoline fraction can be controlled to a gasoline fraction having a distillation range of 40 to 160 ° C, preferably a gasoline fraction having a distillation range of 40 to 150 ° C.
  • the cutting temperature of the medium and heavy gasoline fractions can be further set to 70-130. °C.
  • the catalyst used for the aromatization/hydroisomerization reaction of the middle fraction may be The art of aromatization/hydrogen isomerization catalysts commonly used in FCC gasoline processing.
  • the catalyst used in the aromatization/hydroisomerization reaction is obtained by using a molecular sieve and a metal oxide as a composite support for supporting an active metal component, wherein the active metal is zinc and/or gallium.
  • the molecular sieve may be one or more of an MFI type molecular sieve, an MCM type molecular sieve, and an LTL type molecular sieve, wherein the metal oxide is alumina; wherein the MFI type molecular sieve may be ZSM-5, HZSM- The molecular sieve of 5, MCM type molecular sieve may be molecular sieve such as MCM-41, and the molecular sieve of LTL type may be molecular sieve of L type.
  • the mass ratio of the molecular sieve to the metal oxide in the catalyst used for the aromatization/hydroisomerization reaction is 1: (0.2-0.5), and the loading of the active metal on the composite carrier is 0.5-3. %.
  • the catalyst may be obtained by impregnating the composite carrier with a soluble salt solution of the active metal, drying the impregnated material, and calcining; the impregnation may be an equal volume impregnation.
  • the aromatization/hydroisomerization reaction has a reaction temperature of 260-400 ° C, a reaction pressure of 0.8-2.0 MPa, a hydrogen oil volume ratio of 200-800:1, and a weight hourly space velocity of 1.0- 6.0h -1 .
  • the aromatization/hydroisomerization reaction of the present invention can be carried out using a fixed bed reactor to facilitate control of the reaction process and increase in catalyst efficiency and life.
  • the middle gasoline fraction may be desulfurized to obtain a first desulfurized gasoline fraction. And then aromatizing/hydroisomerizing the gasoline fraction in the first desulfurization in the presence of a catalyst to obtain a second desulfurized gasoline fraction, and then the light gasoline fraction, the second desulfurization gasoline fraction, and The heavy gasoline fraction is mixed to obtain modified gasoline.
  • the desulfurization of the medium gasoline fraction is solvent extraction desulfurization
  • the solvent extraction desulfurization can be carried out using techniques well known in the art without very severe limitations.
  • it may be treated by the gasoline fraction solvent extraction desulfurization method disclosed in the publication No. CN103555359A, and specifically includes the following steps: the middle gasoline fraction is entered from the lower part of the extraction tower, and the solvent enters from the top of the extraction tower, and is pumped.
  • the bottom reflux device of the tower is injected with C5 alkane to control the temperature of the top of the extraction tower to be 55-100 ° C, the temperature of the bottom of the extraction column is 40-80 ° C, the top pressure of the extraction column is 0.2-0.7 MPa, the solvent and the middle gasoline fraction
  • the feed ratio is controlled at 1.0-5.0, and the ratio of C5 alkane to medium gasoline fraction feed is controlled at 0.1-0.5.
  • the addition of C5 alkane in the process of solvent extraction and desulfurization is to increase the separation efficiency.
  • the C5 alkane may be selected from one or both of n-pentane and isopentane.
  • the medium gasoline fraction is first subjected to solvent extraction and desulfurization, thereby separating the desulfurized middle fraction and the raffinate oil, and then performing the aromatization/hydrogen isomerization reaction on the desulfurized middle fraction, which not only reduces the need Aromatization
  • the treatment amount of the fraction of the hydrogenation/hydroisomerization reaction is also advantageous for improving the reaction efficiency of aromatization/hydroisomerization.
  • the raffinate oil may be mixed with the aromatization/hydrogen isomerization reaction product, the light gasoline fraction and the heavy gasoline fraction to obtain a modified gasoline.
  • desulfurization by solvent extraction, solvent selection and separation operations and steps can be determined based on basic knowledge and skill possessed by those skilled in the art.
  • the extraction can be carried out in an extraction column, and the solvent can be selected from diethylene glycol, triethylene glycol, tetraethylene glycol, dimethyl sulfoxide, sulfolane, N-formylmorpholine, N-methylpyrrolidone, polyethylene.
  • the solvent can be selected from diethylene glycol, triethylene glycol, tetraethylene glycol, dimethyl sulfoxide, sulfolane, N-formylmorpholine, N-methylpyrrolidone, polyethylene.
  • an alcohol and propylene carbonate preferably tetraethylene glycol and/or sulfolane.
  • the desulfurization of the middle gasoline fraction is adsorption desulfurization, and the adsorption desulfurization is performed by using a desulfurization adsorbent, wherein the desulfurization adsorbent is supported by an alkali-treated molecular sieve and activated carbon as a composite carrier.
  • the composition is obtained wherein the active metal is selected from one or more of the group consisting of Groups IA, VIII, IB, IIB and VIB of the Periodic Table.
  • the mass ratio of the molecular sieve to the activated carbon is (20-80): (80-20), preferably (20-60): (80-40); the desulfurization adsorbent
  • the type of molecular sieve in the composite carrier is of the X type, Y type or ZSM-5 type.
  • the present invention is not strictly limited to the X-type and ZSM-5 type molecular sieves used; the Y-type molecular sieve has a skeleton silicon-aluminum atomic ratio of not less than 3.0 (determined by XRD method).
  • the present invention is not limited to the activated carbon used, and its specific surface may be usually about 1000 m 2 /g.
  • the active metal selected from Group IA of the periodic table is, for example, potassium (K), sodium (Na) or the like; and the active metal selected from Group VIII of the periodic table is, for example, iron (Fe), cobalt (Co), nickel ( Ni) or the like; an active metal selected from Group IB of the periodic table is, for example, copper (Cu), silver (Ag) or the like; an active metal selected from Group IIB of the periodic table is, for example, zinc (Zn) or the like; and is selected from Group VIB of the periodic table.
  • the active metal is, for example, molybdenum (Mo) or the like.
  • the active metal in the desulfurization adsorbent is at least two selected from the group consisting of Ni, Fe, Ag, Co, Mo, Zn, and K.
  • the loading of Ni on the composite carrier may be 10-30%; the loading of Fe on the composite carrier may be 5-15%; the loading of Ag on the composite carrier may be 5-10%; Co is compounded
  • the loading on the support may be 5-10%; the loading of Mo on the composite support may be 5-10%; the loading of Zn on the composite support may be 5-15%; the loading of K on the composite support Can be 5-15%.
  • the loading is the amount of each active metal supported on the composite support.
  • the active metal in the desulfurization adsorbent is supported on the composite carrier in an amount of 2 to 30%, preferably 5 to 25%, more preferably 5 to 20%.
  • the loading is the total loading of the active metal.
  • the active metal is K and Ni; further, the loading of K on the composite support is 5-15%, the loading of Ni on the composite carrier is 10-25%; further, the mass ratio of K to Ni supported on the composite carrier is (0.2-0.5):1.
  • the active metal is Zn and Fe; further, the loading of Zn on the composite support is 5-15%, and the loading of Fe on the composite support is 8-15%; further
  • the mass ratio of Zn to Fe supported on the composite carrier is (0.5-1):1.
  • the preparation method of the above desulfurization adsorbent may include the following steps:
  • the alkali-treated molecular sieve and the activated carbon are proportionally made into a composite carrier;
  • the composite carrier is impregnated with a soluble salt solution of the active metal, and the impregnated material is dried and calcined to obtain the desulfurization adsorbent.
  • the alkali treatment comprises separately mixing the molecular sieve and the activated carbon according to a mass ratio of molecular sieve or activated carbon: alkali:water to (0.1-2):(0.05-2):(4-15), and maintaining 0.
  • the mixture is stirred at a temperature of -120 ° C for 0.1-24 hours and then dried, and the alkali treatment process includes at least one time.
  • the base used in the alkali treatment of the present invention is not critical, and for example, a 0.1-1.0 mol/L NaOH solution can be used.
  • the temperature of the stirring treatment may be 30-100 ° C, and the time may be 1-10 h; further, the temperature of the stirring treatment may be 70-80 ° C, and the time may be 2-4 h.
  • the drying temperature after the stirring treatment may be, for example, 100 to 120 ° C, and the time may be, for example, 5 to 8 hours.
  • the alkali treatment process can be one or two times.
  • the soluble salt solution of the active metal may be, for example, a sulfate solution, a nitrate solution or the like, preferably a sulfate solution.
  • the impregnation may be an equal volume impregnation, which is a conventional impregnation method in the art, for example, the soluble salt solution of the active metal may be added dropwise to the composite carrier under normal temperature and stirring conditions until The composite carrier is polymerized into a spherical shape and then allowed to stand for a while (for example, 1-3 h).
  • the soluble salt solution of the first active metal is first impregnated on the composite carrier, and after washing, drying and calcining, the solubility of the second active metal is further impregnated.
  • the salt solution is washed, dried and calcined to prepare a composite carrier loaded with two active metal components.
  • the active metal soluble salt When immersing, the active metal soluble salt can be converted according to the loading requirement of each of the above various active metals on the composite carrier and the total loading requirement of the active metal on the composite carrier (loading two or more active metal components). Dosage.
  • the drying of the impregnated material is carried out at 90-120 ° C for 12-24 h, preferably at 110-120 ° C for 18-24 h.
  • the impregnated material is dried and then calcined to be calcined at 450-640 ° C for 4-6 h.
  • the method comprises cooling the dried material to the chamber.
  • the temperature is first raised to 400 ° C at a rate of 6 ° C / min, and then raised to 450-640 ° C at a rate of 3 ° C / min.
  • the adsorptive desulfurization is carried out using a fixed bed atmospheric pressure, and the temperature for controlling the adsorption desulfurization is 20-100 ° C, for example, 30-80 ° C, and the flow rate of the middle gasoline fraction is 0.3-1 mL/min, for example, 0.5 mL. /min.
  • the method for upgrading the catalytic cracked gasoline of the present invention may further comprise:
  • the selective hydrodesulfurization is carried out after mixing the sulfur-rich component with the heavy gasoline fraction.
  • the upgrading method of the catalytic cracking gasoline further comprises:
  • the desulfurization adsorbent after adsorption desulfurization is washed with water vapor, dried with nitrogen at 200-400 ° C, and the desulfurized adsorbent after drying is cooled with nitrogen to realize regeneration of the desulfurization adsorbent.
  • the method for regenerating the desulfurization adsorbent includes sequentially performing steam washing of the desulfurization adsorbent to be regenerated, nitrogen drying at 200 to 400 ° C, and nitrogen cooling.
  • the desulfurization adsorbent after adsorbing desulfurization can be washed with water at 130-180 ° C for 1-3 h, then dried by nitrogen-purging at 200-400 ° C for 10-60 min, and finally purged with nitrogen at room temperature. Cooling was carried out for 60 min.
  • the method for upgrading catalytic cracked gasoline of the present invention may further perform desulfurization treatment on the catalytically cracked gasoline before cutting the catalytic cracked gasoline into light, medium and heavy gasoline fractions; or, in the light gasoline Before the fraction, the desulfurized gasoline fraction and the heavy gasoline fraction are mixed, the light gasoline fraction is subjected to mercaptan treatment to obtain a mercaptan light gasoline fraction, and then the desulfurized light gasoline fraction and the desulfurized gasoline fraction and the heavy gasoline fraction are obtained. Mix and get modified gasoline.
  • the mercaptan treatment may be carried out by a conventional method such as an alkali extraction method or a mercaptan conversion method or the like.
  • the alkali extraction method uses an alkali solution to extract the mercaptan into the alkali solution and removes it.
  • the alkali content in the alkali solution may be 5-50%, and the oil-base volume ratio may be (1-15): 1, the operating temperature. It can be 10-60 ° C;
  • the thiol conversion method is to convert small molecule thiol into other sulfides and can be removed by conventional alkali-free deodorization process, pre-hydrogenation in Prime-G+ process, etc.
  • the alkali deodorization process conditions may be: reactor operating pressure 0.2-1.0 MPa, reaction temperature 20-60 ° C, feed space velocity 0.5-2.0 h-1, volume ratio of air flow to feed amount is 0.2-1.0, catalyst used Both the cocatalyst and the cocatalyst can be used as catalysts in the art.
  • the heavy gasoline fraction may be selectively hydrodesulfurized prior to mixing the light gasoline fraction, the desulfurized gasoline fraction and the heavy gasoline fraction. Desulfurizing the heavy gasoline fraction, and then mixing the desulfurized heavy gasoline fraction with the light gasoline fraction and the desulfurized gasoline fraction to obtain Modified gasoline.
  • the heavy gasoline fraction and the hydrogen gas may be subjected to selective hydrodesulfurization under the action of a selective hydrodesulfurization catalyst to obtain a desulfurized heavy gasoline fraction, wherein the temperature of the selective hydrodesulfurization is 200-300. °C, the pressure is 1.5-2.5MPa, the volumetric space velocity is 1-5h -1 , and the hydrogen oil volume ratio is 400-600.
  • the selective hydrodesulfurization catalyst of the present invention may be a conventional catalyst for selective hydrodesulfurization of gasoline in the prior art, such as RSDS-I, RSDS-21, RSDS-22 catalyst in the RSDS process, Prime-G+ The HR806 and HR841 catalysts in the process, the FGH-20/FGH-11 combination catalyst in the OCT-M process, the HDOS series deep hydrodesulfurization catalyst in the CDOS process, and the like.
  • the hydrodesulfurization catalyst is obtained by supporting a third active metal component, wherein the carrier is a molecular sieve (for example, X type, Y type or ZSM-5 type) or a metal oxide (for example, three Alaluminum oxide), the third active metal comprising Co and Mo.
  • the carrier is a molecular sieve (for example, X type, Y type or ZSM-5 type) or a metal oxide (for example, three Alaluminum oxide), the third active metal comprising Co and Mo.
  • the total loading of Co and Mo on the support is 5-20%.
  • the mass ratio of Co to Mo supported on the carrier is (0.2-0.6):1.
  • the method for upgrading catalytic cracked gasoline of the present invention cuts gasoline raw materials into light, medium and heavy gasoline fractions, and separately treats the characteristics of each gasoline fraction, which is not only flexible in operation, but also beneficial in reducing hydrodesulfurization components.
  • the method can realize deep desulfurization of gasoline raw materials, and the octane number of the whole fraction gasoline is increased by 3 to 5 units, which has great practical value.
  • the specific desulfurization adsorbent can be used in the upgrading method of the catalytic cracking gasoline of the present invention, which has high sulfur capacity, good selectivity to sulfur, and high desulfurization depth, and can remove sulfur to 1 ppmw (by mass). One millionth); in addition to long life and friendly to the environment.
  • the method for upgrading catalytic cracked gasoline of the present invention can wash the desulfurization adsorbent after adsorption desulfurization, and the sulfur-rich component formed by the washing can be mixed with the heavy gasoline fraction to perform selective hydrodesulfurization, thereby avoiding the raw material. Waste, improve the utilization rate of raw materials; at the same time, the desulfurization adsorbent can be regenerated by drying and cooling after washing.
  • the operation is simple, and the regenerated desulfurization adsorbent does not need hydrogen reduction before use, and is environmentally friendly and economical; The agent can be regenerated many times and can still maintain high sulfur capacity and good desulfurization effect after regeneration.
  • the upgrading method of the catalytic cracking gasoline of the present invention, the aromatization/hydroisomerization reaction of the gasoline fraction in the first desulfurization can be carried out on a fixed bed, since the gas residence time in the fixed bed reactor can be strictly controlled, The temperature distribution can be adjusted to improve the conversion and selectivity of the chemical reaction; and the catalyst in the fixed bed reactor is not easy to wear and can be used continuously for a long time; the fixed bed reactor is simple in structure, stable in operation, easy to control, and easy to use. Achieve large-scale and continuous production.
  • Embodiment 1 is a process flow diagram of a method for upgrading a catalytically cracked gasoline according to Embodiment 1 of the present invention
  • FIG. 2 is a process flow diagram of a method for upgrading a catalytically cracked gasoline according to Embodiment 3 of the present invention
  • FIG. 5 is a process flow diagram of a method for upgrading a catalytically cracked gasoline according to Embodiment 5 of the present invention
  • FIG. 6 is a process flow diagram of a method for upgrading a catalytically cracked gasoline according to Embodiment 6 of the present invention.
  • FIG. 7 is a process flow diagram of a method for upgrading a catalytically cracked gasoline according to Embodiment 7 of the present invention.
  • Figure 8 is a process flow diagram of a method for upgrading catalytic cracked gasoline according to Example 8 of the present invention.
  • the HZSM-5 molecular sieve and the alumina were uniformly mixed at a mass ratio of 70:30 to prepare a composite carrier, wherein the mass ratio of the molecular sieve to the alumina was 1:0.4.
  • the composite carrier prepared above was subjected to an equal volume impregnation using an aqueous solution of Ga 2 (SO 4 ) 3 ⁇ 16H 2 O, and the impregnated material was washed with deionized water, dried at 120 ° C for 20 hours, and the dried material was cooled to After room temperature, the temperature is raised to 400 ° C at a rate of 6 ° C / min, and then heated to 550 ° C at a rate of 3 ° C / min, and calcined at 550 ° C for 4 hours to prepare a catalyst, wherein the loading of Ga on the composite carrier It is about 1.8%.
  • the catalytic gasoline produced by catalytic cracking of Daqing atmospheric heavy oil is used as raw material (the composition and properties are shown in Table 1).
  • the process of increasing the octane number of the catalytic cracking gasoline is shown in Figure 1. Specifically:
  • Step 11 The catalytically cracked gasoline is divided into light, medium and heavy gasoline fractions according to the distillation range from low to high, wherein the distillation range of the controlled gasoline fraction is 40-160 °C.
  • Step 12 After placing the catalyst prepared above in a fixed bed reactor, the middle gasoline fraction is introduced into a fixed bed reactor at a reaction temperature of 380 ° C, a reaction pressure of 1.5 MPa, a weight hourly space velocity of 5.0 h -1 , and hydrogen.
  • the aromatization/hydroisomerization reaction was continuously carried out for 200 hours in a fixed bed reactor under an oil to volume ratio of 500:1.
  • Step 13 The reaction product of the above step is taken out, and then mixed with the light gasoline fraction and the heavy gasoline fraction to obtain the modified gasoline.
  • the composition and properties thereof are shown in Table 1. It can be seen from the results in Table 1 that the modified gasoline octane number is greatly improved.
  • the MCM-41 molecular sieve and alumina were mixed at a mass ratio of 80:20 to prepare a composite carrier in which the mass ratio of the molecular sieve to the alumina was 1:0.25.
  • the composite carrier prepared above was impregnated with an equal volume using a ZnSO 4 solution, and the impregnated material was washed with deionized water, dried at 110 ° C for 24 hours, and the dried material was cooled to room temperature, first at 6 ° C / min. The temperature was raised to 400 ° C, and the temperature was raised to 450 ° C at a rate of 3 ° C / min, and calcined at 450 ° C for 6 hours to prepare a catalyst, wherein the loading of Zn on the composite carrier was about 0.5%.
  • the method for increasing the octane number of the catalytic cracking gasoline by using the catalytic gasoline of Example 1 as a raw material is:
  • the catalytically cracked gasoline is divided into light, medium and heavy gasoline fractions according to the distillation range from low to high, wherein the distillation range of the controlled gasoline fraction is 40-160 °C.
  • the middle gasoline fraction is introduced into a fixed bed reactor containing the catalyst prepared above, and the reaction temperature is 260 ° C, the reaction pressure is 0.8 MPa, the weight hourly space velocity is 1 h -1 , and the hydrogen oil volume ratio is 200:1.
  • the aromatization/hydroisomerization reaction was carried out continuously for 200 hours in a fixed bed reactor.
  • the reaction product of the above step is taken out, and then mixed with the light gasoline fraction and the heavy gasoline fraction to obtain the modified gasoline.
  • the composition and properties thereof are shown in Table 1. It can be seen from the results in Table 1 that the modified gasoline octane number is greatly improved.
  • Step 21 The catalytically cracked gasoline is divided into three gasoline fractions of light, medium and heavy according to the distillation range from low to high, wherein the distillation range of the controlled middle fraction is 40-150 °C.
  • Step 22 The middle gasoline fraction is entered from the lower part of the extraction tower, the tetraethylene glycol is introduced from the top of the extraction tower, and the n-pentane is injected into the reflux device at the bottom of the extraction tower to control the temperature of the top of the extraction tower to be 80 ° C, and the extraction is performed.
  • the bottom temperature is 60 ° C
  • the top pressure of the extraction tower absolute pressure
  • the weight ratio of tetraethylene glycol to the middle fraction is controlled at 3.0
  • the weight ratio of n-pentane to middle fraction is controlled at 0.3;
  • the middle fraction and the tetraethylene glycol are contacted in multiple stages in the upper section of the extraction tower, while the n-pentane and tetraethylene glycol are in full contact in the lower part of the extraction tower, wherein the desulfurized middle fraction is carried by the tetraethylene glycol.
  • the bottom is distilled off, and tetraethylene glycol is eluted by water to obtain a desulfurized middle fraction;
  • n-pentane is in full contact with the middle distillate which continues down with tetraethylene glycol in the lower part of the stripping column, and is discharged from the bottom of the column with n-pentane; the n-pentane therein is returned to the stripping column reflux device, The water therein is returned as water washing water to the desulfurized middle distillate water to elute the solvent, and the tetraethylene glycol is returned to the top of the extraction tower to collect the spent sulfur-rich oil.
  • Step 23 The desulfurized middle fraction was introduced into the fixed bed reactor containing the catalyst prepared in Example 1, at a reaction temperature of 300 ° C, a reaction pressure of 1 MPa, a weight hourly space velocity of 2.5 h -1 , and a hydrogen oil volume ratio of Under the conditions of 350:1, aromatization/hydroisomerization reaction was continuously carried out for 200 hours in a fixed bed reactor.
  • Step 24 The reaction product of the above step is taken out, and then with the light gasoline fraction, the spent sulfur-rich oil and the heavy steam The oil fraction is mixed to obtain the modified gasoline, and its composition and properties are shown in Table 2. From the results of Table 2, it can be seen that the modified gasoline octane number is greatly improved.
  • the ZSM-5 molecular sieve and alumina were mixed at a mass ratio of 83:17 to prepare a composite carrier in which the mass ratio of the molecular sieve to the alumina was 1:0.2.
  • the composite carrier prepared above was subjected to an equal volume impregnation using an aqueous solution of Ga 2 (SO 4 ) 3 ⁇ 16H 2 O, and the impregnated material was washed with deionized water, dried at 120 ° C for 18 hours, and the dried material was cooled to After room temperature, the temperature is raised to 400 ° C at a rate of 6 ° C / min, and then heated to 640 ° C at a rate of 3 ° C / min, and calcined at 640 ° C for 5 hours to obtain a catalyst, wherein the loading of Ga on the composite carrier About 3%.
  • the method for increasing the octane number of the catalytic cracking gasoline by using the catalytic gasoline of Example 3 as a raw material is:
  • the catalytically cracked gasoline is divided into three gasoline fractions of light, medium and heavy according to the distillation range from low to high, wherein the distillation range of the controlled gasoline fraction is 50-130 °C.
  • the middle gasoline fraction is taken from the lower part of the extraction tower, the sulfolane enters from the top of the extraction tower, and the isopentane is injected into the reflux device at the bottom of the extraction tower to control the temperature of the top of the extraction tower to be 60 ° C, and the temperature of the bottom of the extraction tower is 40. °C, the top pressure of the extraction tower (absolute pressure) is 0.2MPa, and the weight ratio of sulfolane to middle fraction is controlled at 1.0, the weight ratio of isopentane to middle fraction is controlled at 0.1, and the desulfurized medium gasoline fraction and the surplus are collected. Sulfur oil.
  • the desulfurized medium gasoline fraction is introduced into a fixed bed reactor containing the catalyst prepared above at a reaction temperature of 400 ° C, a reaction pressure of 2 MPa, a weight hourly space velocity of 6 h -1 , and a hydrogen oil volume ratio of 800:1.
  • the aromatization/hydroisomerization reaction was carried out continuously for 200 hours in a fixed bed reactor.
  • the reaction product of the above step is taken out, and then mixed with the light gasoline fraction, the spent sulfur-rich oil fraction and the heavy gasoline fraction to obtain the modified gasoline, and the composition and properties thereof are shown in Table 2. From the results of Table 2, it can be seen that the modified gasoline octane number is greatly improved.
  • the ZSM-5 molecular sieve before alkali treatment exhibits a type I isotherm characteristic of microporous properties, and the desorption isotherm almost coincides with the adsorption isotherm; while the alkali treated ZSM-5 molecular sieve exhibits obvious characteristics.
  • Type IV isotherm which exhibits a continuous adsorption state throughout the measured pressure range up to saturation pressure, with desorption followed by pressure The decrease of force is slowly desorbed. When the pressure reaches a certain value, the amount of desorption suddenly increases, forming a steep curve, and then coincides with the adsorption isotherm as the pressure continues to decrease, thereby indicating the ZSM after alkali treatment. A large number of mesopores (mesopores) are produced in the -5 molecular sieve.
  • the ZSM-5 molecular sieve before the alkali treatment is mainly microporous, has a broad distribution before 2 nm, has a small peak at 3.5 nm, and basically has no pores after 4 nm, using the t-plot method.
  • the calculated average pore diameter is about 2.3 nm; the alkali treated ZSM-5 molecular sieve still has a partial pore distribution before 2 nm, and there is a strong peak around 3.8 nm, and the peak height is almost the ZSM-5 molecular sieve before the alkali treatment. About 11 times, there is a wider pore distribution after 4 nm.
  • the above-mentioned alkali-treated ZSM-5 type molecular sieve and alkali-treated activated carbon were mixed at a mass ratio of 40:60, and then ground into a powder in a mortar, and then dried in an oven at 120 ° C for 6 hours to obtain the first A composite carrier.
  • the first composite carrier prepared above is firstly impregnated with a K 2 SO 4 solution, and after washing, drying and calcining, the first composite carrier impregnated with the K 2 SO 4 solution is subjected to an equal volume impregnation using NiSO 4 . After washing, drying and calcining, a desulfurization adsorbent is prepared;
  • the washing, drying and calcining are as follows: after the immersed material is washed with deionized water, dried at 120 ° C for 20 hours, and after the dried material is cooled to room temperature, the temperature is raised to 400 ° C at a rate of 6 ° C / min. The temperature was further raised to 550 ° C at a rate of 3 ° C/min, and calcined at 550 ° C for 4 hours.
  • the loading of K on the first composite support is about 5%
  • the loading of Ni on the first composite support is about 10%
  • the loading of K and Ni on the first composite support The mass ratio is 0.5:1.
  • the desulfurization adsorbent has a sulfur capacity of 0.514 and a service life of 8-9 hours.
  • the sulfur capacity is 1 g of the desulfurization adsorbent, and the total sulfur content (in grams) removed when the total sulfur content in the gasoline raw material is reduced to less than 10 ppmw, for example, when the sulfur capacity is 0.514, it represents 1 g of desulfurization adsorption.
  • the total sulfur content removed by the agent when the total sulfur content in the gasoline raw material was reduced to 10 ppmw or less was 0.514 g.
  • ZSM-5 molecular sieves were firstly impregnated with CoSO 4 solution, washed, dried and calcined, and then ZSM- impregnated with CoSO 4 solution was treated with an aqueous solution of (NH 4 ) 6 Mo 7 O 24 .4H 2 O.
  • the type 5 molecular sieve is subjected to equal volume impregnation, and after washing, drying and calcination, a selective hydrodesulfurization catalyst is prepared; wherein the specific operations of washing, drying and calcination are as shown in step 1.
  • the selective hydrodesulfurization catalyst prepared above has a total specific surface area of about 356 m 2 /g, a total pore volume of about 0.315 cm 3 ⁇ g -1 , a loading of Co on the carrier of about 4%, and Mo on the carrier.
  • the loading was approximately 10% and the mass ratio of Co to Mo supported on the support was 0.4:1.
  • the HZSM-5 molecular sieve and the alumina were uniformly mixed at a mass ratio of 70:30 to prepare a second composite carrier, wherein the mass ratio of the molecular sieve to the alumina was 1:0.4.
  • the second composite support prepared above is subjected to an equal volume impregnation using an aqueous solution of Ga 2 (SO 4 ) 3 ⁇ 16H 2 O, and after washing, drying and calcination, an aromatization/hydroisomerization catalyst is prepared; For specific operations of washing, drying and calcination, see step 1.
  • the loading of Ga on the second composite support is about 1.8%.
  • the catalytic gasoline produced by catalytic cracking of Daqing atmospheric heavy oil is used as raw material (the composition is shown in Table 4), and the process flow of producing modified gasoline from the gasoline raw material is shown in Fig. 5.
  • the gasoline feedstock was cut into light, medium and heavy gasoline fractions, wherein the cut temperature of the light and medium gasoline fractions was 60 ° C, and the cut temperature of the medium and heavy gasoline fractions was 100 °C.
  • the light gasoline fraction is contacted with the alkali solution for desulfurization, wherein the alkali used is a 20% by mass NaOH solution, the volume ratio of the light gasoline fraction to the NaOH solution is 5:1, and the operating temperature is 30.
  • the mercaptan light gasoline fraction is withdrawn and the oil is withdrawn, and the extracted oil is combined into the heavy gasoline fraction for the next step.
  • the desulfurization adsorbent prepared above is filled in a fixed bed reactor, and the gasoline fraction is adsorbed and desulfurized at a flow rate of 30 ° C and a normal pressure at a flow rate of 0.5 mL/min to obtain a gasoline fraction of the first desulfurization; Further, after the adsorption desulfurization, the desulfurization adsorbent after adsorption and desulfurization is purged with water at 150 ° C for 3 hours for washing, the sulfur-rich component is collected, and the sulfur-rich component is incorporated into the heavy gasoline fraction for the next step.
  • the washed desulfurization adsorbent was dried by nitrogen gas at 300 ° C for 30 min, and the dried desulfurization adsorbent was purged with nitrogen gas (30 ° C) for 30 min at room temperature to cool, the desulfurization adsorbent was regenerated, and the desulfurization adsorption was regenerated three times.
  • the sulfur capacity of the agent is 0.473, and the service life is about 7h.
  • the first desulfurized gasoline fraction is introduced into a fixed bed reactor at a reaction temperature of 380 ° C, a reaction pressure of 1.5 MPa, and a heavy space-time.
  • a speed of 5.0 h -1 and a hydrogen to oil volume ratio of 500:1 aromatization/hydroisomerization reaction was continuously carried out for 200 hours in a fixed bed reactor to obtain a second desulfurized gasoline fraction.
  • the selective hydrodesulfurization catalyst prepared above is packed in a fixed bed reactor, and has a reaction temperature of 260 ° C, a reaction pressure of 1.8 MPa, a volumetric space velocity of 3.0 h -1 , and a hydrogen oil volume ratio of 500.
  • the heavy gasoline fraction from which the oil and the sulfur-rich component are withdrawn is subjected to selective hydrodesulfurization to obtain a desulfurized heavy gasoline fraction.
  • the desulfurized heavy gasoline fraction is mixed with the desulfurized light gasoline fraction and the second desulfurized gasoline fraction to obtain a modified gasoline, and the composition thereof is shown in Table 4.
  • a selective hydrodesulfurization catalyst was prepared according to the method of Example 5, except that the supported Co loading on the carrier was about 6%, the loading of Mo on the carrier was about 10%, and the supported Co and Mo were supported on the carrier.
  • the mass ratio is 0.6:1.
  • the MCM-41 molecular sieve and alumina were mixed at a mass ratio of 80:20 to prepare a second composite carrier in which the molecular sieve to alumina mass ratio was 1:0.25.
  • the second composite carrier prepared above is impregnated with an equal volume using a ZnSO 4 solution, and the impregnated material is washed with deionized water, dried at 110 ° C for 24 hours, and the dried material is cooled to room temperature, first at 6 ° C / The temperature of min was raised to 400 ° C, and then heated to 450 ° C at a rate of 3 ° C / min, and calcined at 450 ° C for 6 hours to prepare a catalyst, wherein the loading of Zn on the second composite support was about 0.5%.
  • the gasoline feedstock is cut into light, medium and heavy gasoline fractions, wherein the light and medium gasoline fractions have a cutting temperature of 50 ° C and the medium and heavy gasoline fractions have a cutting temperature of 90 °C.
  • the light gasoline fraction is contacted with an alkali solution for desulfurization, wherein the alkali used is a 10% by mass NaOH solution, the volume ratio of the light gasoline fraction to the NaOH solution is 5:1, and the operating temperature is 45.
  • the mercaptan light gasoline fraction is withdrawn and the oil is withdrawn, and the extracted oil is combined into the heavy gasoline fraction for the next step.
  • the middle gasoline fraction is taken from the lower part of the extraction tower, the tetraethylene glycol is introduced from the top of the extraction tower, and the n-pentane is injected into the reflux device at the bottom of the extraction tower to control the temperature of the top of the extraction tower to be 80 ° C, and the temperature of the bottom of the extraction tower is extracted.
  • the top pressure of the extraction column absolute pressure
  • the weight ratio of tetraethylene glycol to the gasoline fraction is controlled at 3.0, n-pentane and medium gasoline fraction The weight ratio is controlled at 0.3;
  • the middle gasoline fraction and the tetraethylene glycol are contacted in multiple stages in the upper section of the extraction tower, while the n-pentane and tetraethylene glycol are in full contact in the lower part of the extraction tower, wherein the desulfurized medium gasoline fraction is tetraethylene glycol.
  • n-pentane and the middle gasoline fraction which continues downward with the tetraethylene glycol are sufficiently contacted in the lower portion of the extraction column, and are discharged from the bottom of the column with n-pentane; the n-pentane therein is returned to the extraction column reflux device, Returning the water therein as water washing water to the desulfurized medium gasoline fraction water to elute the solvent, returning the tetraethylene glycol to the top of the extraction tower, and collecting the sulfur-rich oil component and incorporating the heavy gasoline fraction Go to the next step.
  • the first desulfurized gasoline fraction is introduced into a fixed bed reactor containing the aromatization/hydrogen isomerization catalyst prepared above, at a reaction temperature of 260 ° C, a reaction pressure of 0.8 MPa, a weight hourly space velocity of 1 h -1 , and hydrogen. Under the condition of an oil volume ratio of 200:1, aromatization/hydroisomerization reaction was continuously carried out for 200 hours in a fixed bed reactor to obtain a second desulfurized gasoline fraction.
  • the selective hydrodesulfurization catalyst prepared above is packed in a fixed bed reactor, and has a reaction temperature of 300 ° C, a reaction pressure of 1.5 MPa, a volume space velocity of 4.0 h -1 , and a hydrogen oil volume ratio of 600.
  • the heavy gasoline fraction from which the oil and the sulfur-rich component are withdrawn is subjected to selective hydrodesulfurization to obtain a desulfurized heavy gasoline fraction.
  • the desulfurized heavy gasoline fraction is mixed with the desulfurized light gasoline fraction and the second desulfurized gasoline fraction to obtain a modified gasoline, and the composition thereof is shown in Table 4.
  • the method for upgrading gasoline according to Embodiment 5 and Embodiment 6 of the present invention can not only reduce the sulfur content in the gasoline raw material. It is less than 10 ppm, and it is also possible to control the olefin content to 24% or less, and the octane number is remarkably improved.
  • the alkali-treated Y-type molecular sieve and the alkali-treated activated carbon are mixed at a mass ratio of 20:80, and then ground into a powder in a mortar, and then dried in an oven at 110 ° C for 6 hours to obtain a first composite. Carrier.
  • ZnSO 4 using the first impregnation solution prepared as described above for the first support composite and the like is performed, after washing, drying and calcining, and then using the first complex carrier Fe 2 (SO 4) 3 have been impregnated with ZnSO 4 solution incipient wetness impregnation After washing, drying and calcining, preparing a desulfurization adsorbent;
  • the above washing, drying and roasting are specifically: after the immersed material is washed with deionized water, dried at 110 ° C for 24 hours, and after the dried material is cooled to room temperature, the temperature is raised to 400 ° C at a rate of 6 ° C / min. The temperature was further raised to 450 ° C at a rate of 3 ° C / min, and calcined at 450 ° C for 6 hours.
  • the loading of Zn on the first composite support is about 10%
  • the loading of Fe on the first composite support is about 10%
  • the loading of Zn and Fe on the first composite support The mass ratio is 1:1. Inspection
  • the desulfurization adsorbent has a sulfur capacity of 0.481 and a life of 7-8 hours.
  • a selective hydrodesulfurization catalyst was prepared according to the method of Example 5, except that the loading of Co on the support was about 2%, the loading of Mo on the support was about 8%, and the supported Co and Mo on the support.
  • the mass ratio is 0.25:1.
  • the mercaptan feedstock is subjected to mercaptan treatment using a mercaptan conversion method (alkali-free deodorization process), wherein the reactor operating pressure can be controlled to about 0.5 MPa, the reaction temperature is about 40 ° C, and the feed space velocity is 1.0 h -1 .
  • the volume ratio of the air flow to the feed amount is about 0.5, and the mercaptan gasoline is charged.
  • the mercaptan gasoline was cut into light, medium and heavy gasoline fractions, wherein the light and medium gasoline fractions had a cutting temperature of 60 ° C and the medium and heavy gasoline fractions had a cutting temperature of 100 ° C.
  • the desulfurization adsorbent prepared above is filled in a fixed bed reactor, and the gasoline fraction is adsorbed and desulfurized at a flow rate of 30 ° C and a normal pressure at a flow rate of 0.3 mL/min to obtain a gasoline fraction of the first desulfurization; Further, after the adsorptive desulfurization, the desulfurization adsorbent after adsorption and desulfurization is purged with water at 180 ° C for 1 hour for washing, the sulfur-rich component is collected, and the sulfur-rich component is incorporated into the heavy gasoline fraction for the next step.
  • the washed desulfurization adsorbent was dried by nitrogen gas at 400 ° C for 10 min, and the dried desulfurization adsorbent was purged with nitrogen gas (10 ° C) at room temperature for 10 min to cool, the desulfurization adsorbent was regenerated, and the desulfurization adsorption was regenerated three times.
  • the sulfur capacity of the agent is 0.481, and the service life is about 7h.
  • the first desulfurized gasoline fraction was introduced into a fixed bed reactor equipped with the aromatization/hydrogen isomerization catalyst prepared in Example 5, at a reaction temperature of 300 ° C, a reaction pressure of 1 MPa, a weight hourly space velocity of 2.5 h -1 .
  • the aromatization/hydroisomerization reaction was continuously carried out for 200 hours in a fixed bed reactor under a condition of a hydrogen to oil volume ratio of 350:1 to obtain a second desulfurized gasoline fraction.
  • the selective hydrodesulfurization catalyst prepared above is packed in a fixed bed reactor, and has a reaction temperature of 300 ° C, a reaction pressure of 1.5 MPa, a volume space velocity of 4.0 h -1 , and a hydrogen oil volume ratio of 600.
  • the heavy gasoline fraction of the sulfur-rich component is subjected to selective hydrodesulfurization to obtain a desulfurized heavy gasoline fraction.
  • the desulfurized heavy gasoline fraction was mixed with the light gasoline fraction and the second desulfurized gasoline fraction to prepare modified gasoline, and the composition thereof is shown in Table 6.
  • the ZSM-5 molecular sieve and alumina were mixed at a mass ratio of 83:17 to prepare a composite carrier in which the mass ratio of the molecular sieve to the alumina was 1:0.2.
  • the composite carrier prepared above was subjected to an equal volume impregnation using an aqueous solution of Ga 2 (SO 4 ) 3 ⁇ 16H 2 O, and the impregnated material was washed with deionized water, dried at 120 ° C for 18 hours, and the dried material was cooled to After room temperature, the temperature is raised to 400 ° C at a rate of 6 ° C / min, and then heated to 640 ° C at a rate of 3 ° C / min, and calcined at 640 ° C for 5 hours to obtain a catalyst, wherein the loading of Ga on the composite carrier About 3%.
  • the mercaptan feedstock is subjected to mercaptan treatment using a mercaptan conversion process (alkali-free deodorization process), wherein the reactor operating pressure can be controlled to about 0.3 MPa, the reaction temperature is about 60 ° C, and the feed space velocity is 1.5 h -1 .
  • the volume ratio of the air flow to the feed amount is about 1.0, and the mercaptan gasoline is charged.
  • the mercaptan gasoline was cut into light, medium and heavy gasoline fractions, wherein the light and medium gasoline fractions were cut at a temperature of 50 ° C, and the medium and heavy gasoline fractions were cut at a temperature of 90 ° C.
  • the middle gasoline fraction is taken from the lower part of the extraction tower, the sulfolane enters from the top of the extraction tower, and the isopentane is injected into the reflux device at the bottom of the extraction tower to control the temperature of the top of the extraction tower to be 60 ° C, and the temperature of the bottom of the extraction tower is 40.
  • the top pressure of the extraction tower absolute pressure
  • the weight ratio of sulfolane to medium gasoline fraction is controlled at 1.0
  • the weight ratio of isopentane to medium gasoline fraction is controlled at 0.1, respectively, and the gasoline fraction of the first desulfurization is charged and Sulfur-rich oil component.
  • the first desulfurized gasoline fraction is introduced into a fixed bed reactor containing the aromatization/hydrogen isomerization catalyst prepared above, at a reaction temperature of 400 ° C, a reaction pressure of 2 MPa, a weight hourly space velocity of 6 h -1 , and hydrogen oil.
  • the aromatization/hydroisomerization reaction was continuously carried out for 200 hours in a fixed bed reactor under a condition of a volume ratio of 800:1 to obtain a second desulfurized gasoline fraction;
  • the selective hydrodesulfurization catalyst prepared in Example 5 was packed in a fixed bed reactor under the conditions of a reaction temperature of 300 ° C, a reaction pressure of 2.5 MPa, a volume space velocity of 2.0 h -1 , and a hydrogen oil volume ratio of 400.
  • the heavy gasoline fraction combined with the extracted oil and the sulfur-rich component is subjected to selective hydrodesulfurization to obtain a desulfurized heavy gasoline fraction.
  • the desulfurized heavy gasoline fraction was mixed with the light gasoline fraction and the second desulfurized gasoline fraction to prepare modified gasoline, and the composition thereof is shown in Table 6.
  • the gasoline upgrading method of the seventh embodiment and the eighth embodiment of the present invention can not only reduce the sulfur content in the gasoline raw material to less than 10 ppm, but also control the olefin content to 24% or less, and the octane number is remarkably improved.
  • the alkali-treated ZSM-5 type molecular sieve was prepared according to the method of Example 5, the alkali treated ZSM-5 type molecular sieve was uniformly impregnated with the K 2 SO 4 solution and the NiSO 4 solution according to the method of Example 5, and Washing, drying, and roasting to obtain a desulfurization adsorbent. After testing, the desulfurization adsorbent has a sulfur capacity of 0.286 and a lifetime of only 3-4 hours.
  • the alkali-treated activated carbon was prepared according to the method of Example 5, the alkali-treated activated carbon was uniformly impregnated with K 2 SO 4 solution and NiSO 4 solution according to the method of Example 5, and washed, dried, and calcined. Desulfurization adsorbent.
  • the desulfurization adsorbent has a sulfur capacity of 0.236 and a lifetime of only 3-4 hours.
  • the ZSM-5 type molecular sieve of Example 5 (not treated with alkali) and activated carbon (not treated with alkali) were directly mixed at a mass ratio of 40:60, placed in a mortar for grinding, and then placed in an oven at 120 ° C. After drying for 6 h, a composite carrier was obtained.
  • the composite support was subjected to an equal volume impregnation using the K 2 SO 4 solution and the NiSO 4 solution according to the method of Example 5, and washed, dried, and calcined to obtain a desulfurization adsorbent. After testing, the desulfurization adsorbent has a sulfur capacity of 0.155 and a lifetime of only 2-3 hours.

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Abstract

本发明提供一种催化裂化汽油的提质方法。该提质方法包括如下步骤：将催化裂化汽油切割为轻、中、重汽油馏分；在催化剂存在下对所述中汽油馏分进行芳构化/临氢异构化反应，得到脱硫中汽油馏分；将所述轻汽油馏分、脱硫中汽油馏分和重汽油馏分混合，获得改质汽油；其中，轻、中汽油馏分的切割温度为35-60℃，中、重汽油馏分的切割温度为70-160℃。本发明的方法不仅能够实现催化裂化汽油的深度脱硫，而且显著提高了辛烷值。

Description

一种催化裂化汽油的提质方法 技术领域
本发明属于石油化工技术领域，特别涉及一种催化裂化汽油的提质方法。
背景技术
石油资源日趋重质化、劣质化，环保要求日益提高，世界新的环保法规对汽油质量的要求愈加严格。如2018年1月1日即将实施的国V车用汽油标准将要求烯烃含量在24％以下，硫含量在10ppm以下，而辛烷值为93以上。汽油质量标准的提高主要表现为：在烯烃含量及硫含量进一步降低的同时，提高辛烷值。
目前，发达国家主要是从改善“配方”着手来达到相应的质量标准。他们利用多种工艺生产汽油，然后将多种汽油进行调配。一般含烯烃的催化裂化汽油约占1/3以下，含芳烃但不含烯烃的重整汽油约占1/3以上，其它既不含烯烃又不含芳烃的烷基化、异构化、醚化等清洁汽油组分约占1/3。硫含量和烯烃含量低，辛烷值高。
催化裂化汽油是我国车用汽油的主要组成部分，在汽油池中占75％左右。成品汽油中烯烃含量及硫含量的大约90％来自催化裂化汽油，因而造成我国汽油产品远不能满足硫含量≤10ppm、烯烃含量≤24％的新指标要求。另一方面，目前我国以93号汽油为主，但随着国内汽车制造业水平的不断提高以及国内进口汽车保有量的不断增加，对95号或者更高辛烷值的汽油的需求日益增加。催化裂化汽油由于受工艺自身的限制，其辛烷值主要由大量的烯烃来维持，RON一般在90左右，因而其辛烷值的高低直接影响着成品汽油的辛烷值水平。而且，目前催化裂化汽油脱硫降烯烃的主流工艺是催化加氢，不可避免地带来大量烯烃的饱和，造成了较大的辛烷值损失，严重影响了企业的经济效益。
随着原油日趋重质化、重油催化裂化能力不断扩大以及环保法规的日益严格，这一问题显得更为突出，这在客观上迫使石油化工行业不得不研究开发新的催化裂化汽油高效提质工艺，尤其是催化裂化汽油深度脱硫和辛烷值提高同时实现的高效改质工艺。
现有催化裂化汽油降硫技术，主要以中国石化S-zorb、石科院RSDS和法国Prime-G+为代表。S-zorb是美国Conocophillips公司开发，中国石化集团买断并加以完善，用于全馏分催化汽油脱硫，脱后硫含量可以控制到10ppm以下，全馏分汽油的辛烷值损失在 1.0～2.0个单位。RSDS是石油化工科学研究院开发，该技术先将催化汽油切割成轻重馏分，轻馏分经过抽提脱硫醇，重馏分去选择性加氢脱硫；由该技术生产硫含量小于10ppm的产品时，轻馏分产量约20％，大部分需要加氢，全馏分汽油的辛烷值损失在3.0～4.0之间。Prime-G+由法国Axens公司开发，采用全馏分预加氢、轻重汽油分割和重馏分选择性加氢脱硫的工艺流程，其特点是在全馏分预加氢过程中，将轻硫化物与二烯烃作用形成高沸点的硫化物，烯烃没有被饱和，然后通过轻重汽油切割得到硫含量小于10ppm的轻馏分和高硫重馏分，重馏分去加氢脱硫；该技术与RSDS一样，虽然部分低硫轻组分可以不经加氢处理，但由于小于10ppm的轻组分产量很少，大部分都需加氢处理，导致全馏分汽油的辛烷值损失也在3.0～4.0之间。
CN1611572A公开了一种提高汽油辛烷值的催化转化方法。该方法是使初馏点大于100℃的重汽油馏分与温度低于700℃的催化剂接触，在300～660℃、130～450KPa、重时空速为1～120h-1、催化剂与汽油馏分的重量比为2～20、水蒸汽与汽油馏分的重量比为0～0.1的条件下发生反应，分离反应产物和待生剂，待生剂经汽提、再生后循环使用。采用本发明提供的方法可以使催化裂化汽油的辛烷值提高3～10个单位。该方法遵循石油烃类催化裂化机理，使汽油进行氢转移反应和裂化反应，汽油辛烷值虽能提高，但需要先进行馏分切割而只收取初馏点大于100℃的重汽油馏分进行反应，损耗量较大。
CN 1160746A公开了一种提高低品质汽油辛烷值的催化转化方法。该方法是将低辛烷值汽油由常规催化裂化原料入口的上游注入提升管反应器中，与来自再生器的高温催化剂接触，在反应温度为600～730℃、剂油比为6～180、重时空速为1～180时-1的条件下进行反应。该方法可使汽油的辛烷值提高，但该方法中全部的低辛烷值汽油都需要参加反应，汽油的损耗量很大。
CN 103805269 A提出了一种催化汽油深度加氢脱硫方法，轻汽油和中汽油馏分进行无碱脱臭，然后通过加氢预分馏塔分出轻、中汽油，加氢预分馏塔同时引入热柴油；分出的中汽油与重汽油混合后进行选择性加氢，所得馏分油与无碱脱臭的轻汽油混合，得到清洁汽油产品。该方法虽然能够有效脱硫，辛烷值的降低程度也得到了一定程度的缓和，但不能有效提高辛烷值，且工艺流程与本发明有很大差别。
综上所述，现有降低催化裂化汽油硫含量的技术在应对深度脱硫要求时，普遍都存在加氢处理比例大、辛烷值损失多的问题。一些配套加氢脱硫过程的辛烷值恢复工艺也效果不明显。市场上迫切要求开发辛烷值损失少或者辛烷值显著提高的催化裂化汽油深 度脱硫技术。
发明内容
本发明为解决上述的技术问题，提供一种催化裂化汽油的提质方法，不仅能够深度脱除催化裂化汽油所含硫化物在10ppm以下，而且能够显著提高催化裂化汽油辛烷值3-5个单位。
本发明的目的是通过以下技术方案实现的：
一种催化裂化汽油的提质方法，包括以下步骤：
将催化裂化汽油切割为轻、中、重汽油馏分；
在催化剂存在下对所述中汽油馏分进行芳构化/临氢异构化反应，得到脱硫中汽油馏分；
将所述轻汽油馏分、脱硫中汽油馏分和重汽油馏分混合，获得改质汽油；
其中，轻、中汽油馏分的切割温度为35-60℃，中、重汽油馏分的切割温度为70-160℃。
本发明所述的切割是将催化裂化汽油按照沸程从低到高分割为轻、中、重三个汽油馏分，且控制所述中汽油馏分的沸程是35-50℃至130-160℃。
根据本发明的提质方法，首先对催化裂化汽油(FCC汽油)实施馏分切割，通过控制切割温度，使收集的轻汽油馏分为富含烯烃、辛烷值高的催化裂化汽油，中汽油馏分为烯烃和芳烃含量居中而辛烷值最低的催化裂化汽油，重汽油馏分为烯烃含量较低而芳烃含量较高、辛烷值较高的催化裂化汽油。进一步地，本发明对辛烷值最低的中汽油馏分进行芳构化/临氢异构化反应，将反应产物再与其它汽油馏分混合，即可调配得到辛烷值显著提高的FCC汽油。
本发明的实施方案中，根据FCC汽油的情况，可以综合考虑原料汽油的处理量和效果来确定中汽油馏分的馏程范围。发明人的研究发现，馏程为40-160℃的中汽油馏分大约占催化裂化汽油的40m％，基本是辛烷值最低的部分，其RON在80以下，甚至有小部分在70以下，因此中汽油馏分可以控制为馏程为40-160℃的汽油馏分，优选馏程在40-150℃之间的汽油馏分。显然，馏程越长，收取的馏分越多，芳构化/临氢异构化反应需要处理的油量就越大，因此，进一步可以将中、重汽油馏分的切割温度设为70-130℃。
在本发明具体方案中，对中馏分进行芳构化/临氢异构化反应所使用的催化剂可以为 本领域对于FCC汽油处理常用的芳构化/临氢异构化反应催化剂。在一实施方式中，进行芳构化/临氢异构化反应所采用的催化剂由分子筛和金属氧化物作为复合载体负载活性金属成分而得到，其中，活性金属为锌和/或镓。
更具体地，所述分子筛可以是MFI型分子筛、MCM型分子筛和LTL型分子筛中的一种或多种，所述金属氧化物为氧化铝；其中，MFI型分子筛可以是ZSM-5、HZSM-5等分子筛，MCM型分子筛可以是MCM-41等分子筛，LTL型分子筛可以是L型分子筛等。
进一步地，进行芳构化/临氢异构化反应所采用的催化剂中的分子筛与金属氧化物的质量比为1：(0.2-0.5)，活性金属在复合载体上的负载量为0.5-3％。可以通过将所述复合载体用所述活性金属的可溶性盐溶液进行浸渍，将浸渍后的物料干燥后焙烧，得到所述催化剂；所述浸渍可以为等体积浸渍。
更进一步地，所述芳构化/临氢异构化反应的反应温度为260-400℃，反应压力为0.8-2.0MPa，氢油体积比为200-800:1，重时空速为1.0-6.0h-1。并且，本发明所述芳构化/临氢异构化反应可以利用固定床反应器进行，从而利于反应过程的控制和催化剂效率和寿命提升。
本发明的催化裂化汽油的提质方法，还可以在对所述中汽油馏分进行芳构化/临氢异构化反应之前，先对所述中汽油馏分进行脱硫，得到第一脱硫中汽油馏分，然后在催化剂存在下对所述第一脱硫中汽油馏分进行芳构化/临氢异构化反应，得到第二脱硫中汽油馏分，再将所述轻汽油馏分、第二脱硫中汽油馏分和重汽油馏分混合，获得改质汽油。
在一实施方式中，对所述中汽油馏分进行的脱硫为溶剂抽提脱硫，溶剂抽提脱硫可以利用本领域中熟知的技术进行，没有非常严格的限制。例如可以是采用公开号为CN103555359A专利中的汽油馏分溶剂抽提脱硫方法进行处理，具体包括以下步骤：使所述中汽油馏分从抽提塔中下部进入，溶剂从抽提塔顶部进入，从抽提塔底部回流装置注入C5烷烃，控制抽提塔顶温度为55-100℃，抽提塔底温度为40-80℃，抽提塔顶压为0.2-0.7MPa，所述溶剂与中汽油馏分进料比控制在1.0-5.0，C5烷烃与中汽油馏分进料比控制在0.1-0.5。其中，在溶剂抽提脱硫的过程添加C5烷烃是为了增加分离效率。在本发明的一个实施方案中，C5烷烃可选自正戊烷和异戊烷中的一种或两种。
该方式先对所述中汽油馏分进行溶剂抽提脱硫，从而分离出脱硫的中馏分和抽余油，随后对脱硫的中馏分实施芳构化/临氢异构化反应，其不仅减少了需要进行芳构 化/临氢异构化反应的馏分的处理量，同时也利于提高芳构化/临氢异构化的反应效率。进一步地，所述抽余油可与芳构化/临氢异构化反应产物、轻汽油馏分和重汽油馏分混合，获得改质汽油。
此外，利用溶剂抽提进行脱硫，对于溶剂的选择及分离操作和步骤，均可基于本领域技术人员所具备的基础常识和技能进行确定。例如，抽提可以在抽提塔中完成，溶剂可以选择二甘醇、三甘醇、四甘醇、二甲亚砜、环丁砜、N-甲酰吗啉、N-甲基吡咯烷酮、聚乙二醇和碳酸丙烯酯中一种或多种；优选为四甘醇和/或环丁砜。
在另一实施方式中，对所述中汽油馏分进行的脱硫为吸附脱硫，利用脱硫吸附剂进行所述吸附脱硫，所述脱硫吸附剂由分别经碱处理的分子筛和活性炭作为复合载体负载活性金属成分而得到，其中，活性金属选自周期表IA、VIII、IB、IIB和VIB族中的一种或多种元素。
本发明所述脱硫吸附剂的复合载体中，分子筛与活性炭的质量比为(20-80):(80-20)，优选为(20-60):(80-40)；所述脱硫吸附剂的复合载体中的分子筛的类型为X型、Y型或ZSM-5型。本发明对所采用X型和ZSM-5型分子筛无严格限制；所述Y型分子筛的骨架硅铝原子比不小于3.0(XRD法测定)。此外，本发明对所采用的活性炭无严格限制，其比表面通常可为1000m2/g左右。
在本发明中，选自周期表IA族的活性金属例如为钾(K)、钠(Na)等；选自周期表VIII族的活性金属例如为铁(Fe)、钴(Co)、镍(Ni)等；选自周期表IB族的活性金属例如为铜(Cu)、银(Ag)等；选自周期表IIB族的活性金属例如为锌(Zn)等；选自周期表VIB族的活性金属例如为钼(Mo)等。
进一步地，所述脱硫吸附剂中的活性金属选自Ni、Fe、Ag、Co、Mo、Zn和K中的至少2种。其中，Ni在复合载体上的负载量可为10-30％；Fe在复合载体上的负载量可为5-15％；Ag在复合载体上的负载量可为5-10％；Co在复合载体上的负载量可为5-10％；Mo在复合载体上的负载量可为5-10％；Zn在复合载体上的负载量可为5-15％；K在复合载体上的负载量可为5-15％。该负载量为每种活性金属各自在复合载体上的负载量。
进一步地，所述脱硫吸附剂中的活性金属在复合载体上的负载量为2-30％，优选为5-25％，进一步优选为5-20％。在复合载体上负载两种以上活性金属时，所述负载量为活性金属的总负载量。
在一实施方式中，所述活性金属为K和Ni；进一步地，K在复合载体上的负载量为 5-15％，Ni在复合载体上的负载量为10-25％；更进一步地，复合载体上负载的K与Ni的质量比为(0.2-0.5)：1。
在另一实施方式中，所述活性金属为Zn和Fe；进一步地，Zn在复合载体上的负载量为5-15％，Fe在复合载体上的负载量为8-15％；更进一步地，复合载体上负载的Zn与Fe的质量比为(0.5-1)：1。
上述的脱硫吸附剂的制备方法，可以包括以下步骤：
将分别经碱处理的分子筛和活性炭按比例制成复合载体；
将所述复合载体用所述活性金属的可溶性盐溶液进行浸渍，将浸渍后的物料干燥后焙烧，得到所述脱硫吸附剂。
在一实施方式中，所述碱处理包括分别对分子筛和活性炭按照分子筛或活性炭：碱：水为(0.1-2)：(0.05-2)：(4-15)的质量比混合，并维持0-120℃的温度条件下搅拌处理0.1-24h后干燥，且所述碱处理过程包括至少一次。
本发明对碱处理所采用的碱无严格限制，例如可以采用0.1-1.0mol/L的NaOH溶液。进一步地，搅拌处理的温度可以为30-100℃，时间可以为1-10h；更进一步地，搅拌处理的温度可以为70-80℃，时间可以为2-4h。所述搅拌处理后的干燥的温度例如可以为100-120℃，时间例如可以为5-8h。所述碱处理过程可以为一次或两次。
在本发明中，所述活性金属的可溶性盐溶液例如可以为硫酸盐溶液、硝酸盐溶液等，优选为硫酸盐溶液。所述浸渍可以为等体积浸渍，其为本领域常规的浸渍方式，具体操作例如可以为：在常温和搅拌的条件下，向所述复合载体中滴加所述活性金属的可溶性盐溶液，直至复合载体聚合成球状，然后静置一段时间(例如1-3h)。特别是，在复合载体上负载两种活性金属成分时，先在所述复合载体上浸渍第一种活性金属的可溶性盐溶液，经洗涤、干燥和焙烧后，再浸渍第二种活性金属的可溶性盐溶液，经洗涤、干燥和焙烧，即可制得负载两种活性金属成分的复合载体。
浸渍时，可根据上述各种活性金属各自在复合载体上的负载量要求以及活性金属在复合载体上的总负载量要求(负载两种以上活性金属成分)换算各活性金属可溶性盐在浸渍时的用量。
进一步地，对浸渍后的物料的干燥为在90-120℃下干燥12-24h，优选为在110-120℃下干燥18-24h。对浸渍后的物料干燥后进行焙烧为在450-640℃焙烧4-6h。
进一步地，所述对浸渍后的物料干燥后进行焙烧时，包括将干燥后的物料冷却至室 温，先以6℃/min速度升温至400℃，再以3℃/min速度升温至450-640℃。
在本发明中，所述吸附脱硫是利用固定床常压进行，并且控制吸附脱硫的温度为20-100℃，例如30-80℃，中汽油馏分的流速为0.3-1mL/min，例如0.5mL/min。
本发明的催化裂化汽油的提质方法，还可以包括：
采用水蒸气对吸附脱硫后的脱硫吸附剂进行洗涤，收取富硫组分；
将所述富硫组分与所述重汽油馏分混合后进行所述选择性加氢脱硫。
进一步地，所述催化裂化汽油的提质方法还包括：
采用水蒸气对吸附脱硫后的脱硫吸附剂进行洗涤后采用200-400℃的氮气进行干燥，并采用氮气对干燥后的脱硫吸附剂进行冷却，实现对脱硫吸附剂的再生。
即，所述脱硫吸附剂的再生方法，包括对待再生的所述脱硫吸附剂顺序进行水蒸气洗涤、200-400℃的氮气干燥，和氮气冷却。
具体地，可以采用130-180℃水蒸气吹扫吸附脱硫后的脱硫吸附剂1-3h进行洗涤，然后采用200-400℃氮气吹扫10-60min进行干燥，最后采用室温的氮气吹扫10-60min进行冷却。
进一步地，本发明的催化裂化汽油的提质方法，还可以在将催化裂化汽油切割为轻、中、重汽油馏分之前，先对催化裂化汽油进行脱硫醇处理；或者，在将所述轻汽油馏分、脱硫中汽油馏分和重汽油馏分混合之前，先对所述轻汽油馏分进行脱硫醇处理，得到脱硫醇轻汽油馏分，然后将所述脱硫醇轻汽油馏分与脱硫中汽油馏分和重汽油馏分混合，获得改质汽油。
在本发明中，可以采用常规方法进行所述脱硫醇处理，例如碱抽提法或硫醇转化法等。碱抽提法使用碱液将硫醇抽提到碱液中而脱除，碱液中碱的质量含量可为5-50％，油碱体积比可为(1-15)：1，操作温度可为10-60℃；硫醇转化法是将小分子硫醇转化为其它硫化物而脱除，可以采用常规的无碱脱臭工艺、Prime-G+工艺中的预加氢等方式进行，其中无碱脱臭工艺条件可以为：反应器操作压力0.2-1.0MPa，反应温度20-60℃，进料空速0.5-2.0h-1，空气流量与进料量的体积比为0.2-1.0，所用催化剂及助催化剂均可以为本领域常用的催化剂。
进一步地，本发明的催化裂化汽油的提质方法，还可以在将所述轻汽油馏分、脱硫中汽油馏分和重汽油馏分混合之前，先对所述重汽油馏分进行选择性加氢脱硫，得到脱硫重汽油馏分，然后将所述脱硫重汽油馏分与轻汽油馏分和脱硫中汽油馏分混合，获得 改质汽油。
具体地，可以将所述重汽油馏分、氢气在选择性加氢脱硫催化剂的作用下进行选择性加氢脱硫，得到脱硫重汽油馏分，其中，所述选择性加氢脱硫的温度为200-300℃，压力为1.5-2.5MPa，体积空速为1-5h-1，氢油体积比为400-600。
本发明所述的选择性加氢脱硫催化剂可以为现有技术中对汽油进行选择性加氢脱硫的常规催化剂，例如RSDS工艺中的RSDS-Ⅰ、RSDS-21、RSDS-22催化剂，Prime-G+工艺中的HR806和HR841催化剂，OCT-M工艺中的FGH-20/FGH-11组合催化剂，CDOS工艺中的HDOS系列深度加氢脱硫催化剂等。
在一实施方式中，所述加氢脱硫催化剂由载体负载第三活性金属成分而得到，其中，所述载体为分子筛(例如X型、Y型或ZSM-5型)或金属氧化物(例如三氧化二铝)，所述第三活性金属包括Co和Mo。进一步地，Co和Mo在所述载体上的总负载量为5-20％。更进一步地，载体上负载的Co与Mo的质量比为(0.2-0.6)：1。
本发明的实施，至少具有以下优势：
1、本发明的催化裂化汽油的提质方法将汽油原料切割为轻、中、重汽油馏分，并且针对各汽油馏分的特点分别进行处理，不仅操作灵活，还有利于减少加氢脱硫的组分含量；此外，该方法能够实现对汽油原料的深度脱硫，并且全馏分汽油的辛烷值增加3～5个单位，具有极大的实用价值。
2、本发明的催化裂化汽油的提质方法中可使用特定的脱硫吸附剂，其不仅硫容大、对硫的选择性好、而且脱硫深度高，可将硫脱至1ppmw(按质量计的百万分之一)；此外使用寿命长，对环境较为友好。
3、本发明的催化裂化汽油的提质方法可在吸附脱硫后对脱硫吸附剂进行洗涤，洗涤所形成的富硫组分可与重汽油馏分混合后进行选择性加氢脱硫，从而避免了原料浪费，提高了原料利用率；同时，在洗涤后进行干燥和冷却即可实现脱硫吸附剂的再生，该方式操作简单，并且再生的脱硫吸附剂在使用前无需氢气还原，环保经济；此外脱硫吸附剂可多次再生，并且再生后仍然能够维持较高的硫容和良好的脱硫效果。
4、本发明的催化裂化汽油的提质方法在对第一脱硫中汽油馏分的芳构化/临氢异构化反应可以在固定床上进行，由于固定床反应器中气体停留时间可以严格控制，温度分布可以调节，从而有利于提高化学反应的转化率和选择性；并且固定床反应器中催化剂不易磨损，可以较长时间连续使用；固定床反应器结构简单、操作稳定、便于控制、易 实现大型化和连续化生产。
附图说明
图1为本发明实施例1的催化裂化汽油的提质方法的工艺流程图；
图2为本发明实施例3的催化裂化汽油的提质方法的工艺流程图；
图3为实施例5的ZSM-5型分子筛在碱处理前后的吸附脱附等温线；
图4为实施例5的ZSM-5型分子筛在碱处理前后的孔径分布曲线；
图5为本发明实施例5的催化裂化汽油的提质方法的工艺流程图；
图6为本发明实施例6的催化裂化汽油的提质方法的工艺流程图；
图7为本发明实施例7的催化裂化汽油的提质方法的工艺流程图；
图8为本发明实施例8的催化裂化汽油的提质方法的工艺流程图。
具体实施方式
为使本发明的目的、技术方案和优点更加清楚，下面将结合本发明实施例，对本发明实施例中的技术方案进行清楚、完整地描述，显然，所描述的实施例是本发明一部分实施例，而不是全部的实施例。基于本发明中的实施例，本领域普通技术人员在没有做出创造性劳动前提下所获得的所有其他实施例，都属于本发明保护的范围。
实施例1
1、制备催化剂
将HZSM-5分子筛和氧化铝按照质量比70:30混合均匀，制得复合载体，其中分子筛与氧化铝的质量比为1：0.4。
采用Ga2(SO4)3·16H2O的水溶液对上述制备的复合载体进行等体积浸渍，将浸渍后的物料用去离子水洗涤后，120℃干燥20小时，将干燥后的物料冷却至室温后，先以6℃/min的速度升温至400℃，再以3℃/min的速度升温至550℃，在550℃下焙烧4小时，制得催化剂，其中Ga在复合载体上的负载量约为1.8％。
2、汽油改质
以大庆常压重油经过催化裂化生产出的催化汽油为原料(其组成及性质见表1)，提高该催化裂化汽油辛烷值的方法的工艺流程如图1所示，具体为：
步骤11：将催化裂化汽油按照馏程从低到高分割为轻、中、重三个汽油馏分，其中，控制中汽油馏分的馏程为40-160℃。
步骤12：将上述制备的催化剂置于固定床反应器后，将中汽油馏分引入固定床反应器中，在反应温度为380℃、反应压力为1.5MPa、重时空速为5.0h-1、氢油体积比为500:1的条件下，在固定床反应器中连续实施200小时的芳构化/临氢异构化反应。
步骤13：将上述步骤的反应产物引出，然后与轻汽油馏分及重汽油馏分混合，即可获得改质汽油，其组成及性质见表1。由表1结果可知：改质后的汽油辛烷值大幅提高。
实施例2
1、制备催化剂
将MCM-41分子筛和氧化铝按照质量比80:20混合，制得复合载体，其中分子筛与氧化铝的质量比为1：0.25。
采用ZnSO4溶液对上述制备的复合载体进行等体积浸渍，将浸渍后的物料用去离子水洗涤后，110℃干燥24小时，将干燥后的物料冷却至室温后，先以6℃/min的速度升温至400℃，再以3℃/min的速度升温至450℃，在450℃下焙烧6小时，制得催化剂，其中Zn在复合载体上的负载量约为0.5％。
2、汽油改质
以实施例1的催化汽油为原料，提高该催化裂化汽油辛烷值的方法为：
将催化裂化汽油按照馏程从低到高分割为轻、中、重三个汽油馏分，其中，控制中汽油馏分的馏程为40-160℃。
将中汽油馏分引入装有上述制备的催化剂的固定床反应器中，在反应温度为260℃、反应压力为0.8MPa、重时空速为1h-1、氢油体积比为200:1的条件下，在固定床反应器中连续实施200小时的芳构化/临氢异构化反应。
将上述步骤的反应产物引出，然后与轻汽油馏分及重汽油馏分混合，即可获得改质汽油，其组成及性质见表1。由表1结果可知：改质后的汽油辛烷值大幅提高。
表1汽油的组成和性质
Figure PCTCN2015075888-appb-000001
Figure PCTCN2015075888-appb-000002
实施例3
以济南的催化汽油为原料(其组成及性质见表2)，提高该催化裂化汽油辛烷值的方法的工艺流程如图2所示，具体为：
步骤21：将催化裂化汽油按照馏程从低到高分割为轻、中、重三个汽油馏分，其中，控制中馏分的馏程为40-150℃。
步骤22：使中汽油馏分从抽提塔中下部进入，四甘醇从抽提塔顶部进入，同时向抽提塔底部回流装置注入正戊烷，控制抽提塔顶温度为80℃，抽提塔底温度为60℃，抽提塔顶压(绝对压力)为0.5MPa，并且将四甘醇与中馏分重量比控制在3.0，正戊烷与中馏分重量比控制在0.3；
抽提过程中，中馏分与四甘醇在抽提塔上段经多级逆流接触，同时正戊烷与四甘醇在抽提塔下段充分接触，其中的脱硫的中馏分被四甘醇携带从塔顶馏出，经水洗脱去四甘醇，得到脱硫的中馏分；
正戊烷与随四甘醇继续下行的中馏分油在抽提塔下段充分接触，并随正戊烷从塔底出塔排出；将其中的正戊烷返回所述的抽提塔回流装置，将其中的水作为水洗水返回所述的脱硫的中馏分水洗脱去溶剂的步骤，将其中的四甘醇返回所述的抽提塔顶，收取抽余的富硫油分。
步骤23：将脱硫的中馏分引入装有实施例1制备的催化剂的固定床反应器中，在反应温度为300℃、反应压力为1MPa、重时空速为2.5h-1、氢油体积比为350:1的条件下，在固定床反应器中连续实施200小时的芳构化/临氢异构化反应。
步骤24：将上述步骤的反应产物引出，然后与轻汽油馏分、抽余的富硫油分及重汽 油馏分混合，即可获得改质汽油，其组成及性质见表2。由表2结果可知：改质后的汽油辛烷值大幅提高。
实施例4
1、制备催化剂
将ZSM-5分子筛和氧化铝按照质量比83:17混合，制得复合载体，其中分子筛与氧化铝的质量比为1：0.2。
采用Ga2(SO4)3·16H2O的水溶液对上述制备的复合载体进行等体积浸渍，将浸渍后的物料用去离子水洗涤后，120℃干燥18小时，将干燥后的物料冷却至室温后，先以6℃/min的速度升温至400℃，再以3℃/min的速度升温至640℃，在640℃下焙烧5小时，制得催化剂，其中Ga在复合载体上的负载量约为3％。
2、汽油改质
以实施例3的催化汽油为原料，提高该催化裂化汽油辛烷值的方法为：
将催化裂化汽油按照馏程从低到高分割为轻、中、重三个汽油馏分，其中，控制中汽油馏分的馏程为50-130℃。
使中汽油馏分从抽提塔中下部进入，环丁砜从抽提塔顶部进入，同时向抽提塔底部回流装置注入异戊烷，控制抽提塔顶温度为60℃，抽提塔底温度为40℃，抽提塔顶压(绝对压力)为0.2MPa，且将环丁砜与中馏分重量比控制在1.0，异戊烷与中馏分重量比控制在0.1，收取脱硫的中汽油馏分和抽余的富硫油分。
将脱硫的中汽油馏分引入装有上述制备的催化剂的固定床反应器中，在反应温度为400℃、反应压力为2MPa、重时空速为6h-1、氢油体积比为800:1的条件下，在固定床反应器中连续实施200小时的芳构化/临氢异构化反应。
将上述步骤的反应产物引出，然后与轻汽油馏分、抽余的富硫油分及重汽油馏分混合，即可获得改质汽油，其组成及性质见表2。由表2结果可知：改质后的汽油辛烷值大幅提高。
表2汽油的组成和性质
Figure PCTCN2015075888-appb-000003
Figure PCTCN2015075888-appb-000004
实施例5
1、制备脱硫吸附剂
1)制备经碱处理的分子筛和活性炭
将两份500mL浓度为0.3mol/L的NaOH溶液水浴升温至70℃左右后，分别向其中加入25g的ZSM-5型分子筛和25g的活性炭，搅拌200分钟左右后，立即将混合物用冰浴降至常温，过滤并用去离子水过滤洗涤滤饼多次，直至滤出液的pH值为7左右后，将所得滤饼置于110℃的烘箱中干燥4h，分别制得经碱处理的ZSM-5型分子筛和经碱处理的活性炭；
此外，采用ASAP2000型自动物理吸附仪测定ZSM-5型分子筛和活性炭的比表面积及孔径分布，结果如表3所示。
表3ZSM-5型分子筛和活性炭的比表面积及孔径
Figure PCTCN2015075888-appb-000005
由图3可知：碱处理前的ZSM-5分子筛表现出微孔性质特有的I型等温线，其脱附等温线几乎与吸附等温线重合；而碱处理后的ZSM-5分子筛表现出特征明显的IV型等温线，其在整个测量压力范围内呈现持续的吸附状态直至饱和压力，而脱附时先随着压 力的减小缓慢脱附，当压力达到某一值时脱附量突然增加，形成较为陡峭的变化曲线，然后随着压力的继续降低而与吸附等温线重合，由此说明碱处理后的ZSM-5分子筛中产生了大量介孔(中孔)。
由图4可知，碱处理前的ZSM-5分子筛主要以微孔为主，在2nm之前有较宽的分布，在3.5nm处有一个小峰，4nm之后基本上没有孔出现，使用t-plot方法计算得到的平均孔径为2.3nm左右；碱处理后的ZSM-5分子筛在2nm之前仍有部分微孔分布，而在3.8nm左右有一处强峰，峰高几乎为碱处理前ZSM-5分子筛的11倍左右，在4nm之后也有较为宽泛的孔分布。
同时，表3结果表明：经碱处理的ZSM-5型分子筛中孔体积和平均孔径显著增大，说明大量微孔转变为中孔，从而形成介孔和微孔复合孔结构；经碱处理的活性炭的总比表面积、总孔体积、中孔体积和平均孔径均有所增加。
2)制备第一复合载体
将上述经碱处理的ZSM-5型分子筛和经碱处理的活性炭按照质量比40:60混合后，置于研钵中研磨成粉状，随后置于120℃的烘箱中干燥6h，制得第一复合载体。
3)制备脱硫吸附剂
先采用K2SO4溶液对上述制备的第一复合载体进行等体积浸渍，经洗涤、干燥和焙烧后，再采用NiSO4对已浸渍K2SO4溶液的第一复合载体进行等体积浸渍，经洗涤、干燥和焙烧后，制得脱硫吸附剂；
上述洗涤、干燥和焙烧具体为：将浸渍后的物料用去离子水洗涤后，120℃干燥20小时，将干燥后的物料冷却至室温后，先以6℃/min的速度升温至400℃，再以3℃/min的速度升温至550℃，在550℃下焙烧4小时。
上述制备的脱硫吸附剂中，K在第一复合载体上的负载量约为5％，Ni在第一复合载体上的负载量约为10％，并且第一复合载体上负载的K与Ni的质量比为0.5：1。经检测，该脱硫吸附剂的硫容为0.514，寿命长达8-9h。
在本发明中，硫容为1g脱硫吸附剂将汽油原料中的总含硫量降至10ppmw以下时所脱除的总硫量(以克计)，例如硫容为0.514时，代表1g脱硫吸附剂将汽油原料中的总含硫量降至10ppmw以下时所脱除的总硫量为0.514g。
2、制备选择性加氢脱硫催化剂
先采用CoSO4溶液对ZSM-5型分子筛进行等体积浸渍，经洗涤、干燥和焙烧后，再 采用(NH4)6Mo7O24.4H2O的水溶液对已浸渍CoSO4溶液的ZSM-5型分子筛进行等体积浸渍，经洗涤、干燥和焙烧后，制得选择性加氢脱硫催化剂；其中，洗涤、干燥和焙烧的具体操作参见步骤1。
上述制备的选择性加氢脱硫催化剂的总比表面为356m2/g左右，总孔体积为0.315cm3·g-1左右，Co在载体上的负载量约为4％，Mo在载体上的负载量约为10％，并且载体上负载的Co与Mo的质量比为0.4：1。
3、制备芳构化/临氢异构化反应催化剂
将HZSM-5分子筛和氧化铝按照质量比70:30混合均匀，制得第二复合载体，其中分子筛与氧化铝的质量比为1：0.4。
采用Ga2(SO4)3·16H2O的水溶液对上述制备的第二复合载体进行等体积浸渍，经洗涤、干燥和焙烧后，制得芳构化/临氢异构化反应催化剂；其中，洗涤、干燥和焙烧的具体操作参见步骤1，Ga在第二复合载体上的负载量约为1.8％。
4、汽油改质
以大庆常压重油经过催化裂化生产出的催化汽油为原料(其组成见表4)，以该汽油原料生产改质汽油的工艺流程如图5所示。
首先，将该汽油原料切割为轻、中和重汽油馏分，其中轻、中汽油馏分的切割温度为60℃，中、重汽油馏分的切割温度为100℃。
在抽提系统中使轻汽油馏分与碱溶液接触进行脱硫醇处理，其中所采用的碱为质量含量20％的NaOH溶液，轻汽油馏分与NaOH溶液的体积比为5：1，操作温度为30℃，收取脱硫醇轻汽油馏分和抽出油，将该抽出油并入重汽油馏分进行下一步骤。
将上述制备的脱硫吸附剂填装于固定床反应器中，在温度为30℃以及常压条件下，以0.5mL/min的流速对中汽油馏分进行吸附脱硫，得到第一脱硫中汽油馏分；并且，在吸附脱硫后，采用150℃的水蒸气吹扫吸附脱硫后的脱硫吸附剂3h进行洗涤，收取富硫组分，将该富硫组分并入重汽油馏分进行下一步骤。此外，采用300℃氮气吹扫经洗涤的脱硫吸附剂30min进行干燥，并采用室温的氮气(30℃)吹扫经干燥的脱硫吸附剂30min进行冷却，使脱硫吸附剂再生，再生三次的脱硫吸附剂的硫容为0.473，寿命达7h左右。
将上述制备的芳构化/临氢异构化反应催化剂置于固定床反应器后，将第一脱硫中汽油馏分引入固定床反应器中，在反应温度380℃、反应压力1.5MPa、重时空速5.0h-1、氢油体积比为500:1的条件下，在固定床反应器中连续实施200小时的芳构化/临氢异构化 反应，获得第二脱硫中汽油馏分。
将上述制备的选择性加氢脱硫催化剂填装于固定床反应器中，在反应温度260℃、反应压力1.8MPa、体积空速3.0h-1、氢油体积比为500的条件下对合并有抽出油和富硫组分的重汽油馏分进行选择性加氢脱硫，得到脱硫重汽油馏分。将脱硫重汽油馏分与脱硫醇轻汽油馏分和第二脱硫中汽油馏分混合，制得改质汽油，其组成见表4。
实施例6
1、制备选择性加氢脱硫催化剂
按照实施例5方法制备选择性加氢脱硫催化剂，不同的是，控制Co在载体上的负载量约为6％，Mo在载体上的负载量约为10％，并且载体上负载的Co与Mo的质量比为0.6：1。
2、制备芳构化/临氢异构化反应催化剂
将MCM-41分子筛和氧化铝按照质量比80:20混合，制得第二复合载体，其中分子筛与氧化铝的质量比为1：0.25。
采用ZnSO4溶液对上述制备的第二复合载体进行等体积浸渍，将浸渍后的物料用去离子水洗涤后，110℃干燥24小时，将干燥后的物料冷却至室温后，先以6℃/min的速度升温至400℃，再以3℃/min的速度升温至450℃，在450℃下焙烧6小时，制得催化剂，其中Zn在第二复合载体上的负载量约为0.5％。
3、汽油改质
以大庆的催化汽油为原料(其组成见表4)，以该汽油原料生产改质汽油的工艺流程如图6所示。
首先，将该汽油原料切割为轻、中和重汽油馏分，其中轻、中汽油馏分的切割温度为50℃，中、重汽油馏分的切割温度为90℃。
在抽提系统中使轻汽油馏分与碱溶液接触进行脱硫醇处理，其中所采用的碱为质量含量10％的NaOH溶液，轻汽油馏分与NaOH溶液的体积比为5：1，操作温度为45℃，收取脱硫醇轻汽油馏分和抽出油，将该抽出油并入重汽油馏分进行下一步骤。
将中汽油馏分从抽提塔中下部进入，四甘醇从抽提塔顶部进入，同时向抽提塔底部回流装置注入正戊烷，控制抽提塔顶温度为80℃，抽提塔底温度为60℃，抽提塔顶压(绝对压力)为0.5MPa，且将四甘醇与中汽油馏分重量比控制在3.0，正戊烷与中汽油馏分 重量比控制在0.3；
抽提过程中，中汽油馏分与四甘醇在抽提塔上段经多级逆流接触，同时正戊烷与四甘醇在抽提塔下段充分接触，其中的脱硫的中汽油馏分被四甘醇携带从塔顶馏出，经水洗脱去四甘醇，得到第一脱硫中汽油馏分；
正戊烷与随四甘醇继续下行的中汽油馏分在抽提塔下段充分接触，并随正戊烷从塔底出塔排出；将其中的正戊烷返回所述的抽提塔回流装置，将其中的水作为水洗水返回所述的脱硫的中汽油馏分水洗脱去溶剂的步骤，将其中的四甘醇返回所述的抽提塔顶，收取富硫油组分并入重汽油馏分进行下一步骤。
将第一脱硫中汽油馏分引入装有上述制备的芳构化/临氢异构化反应催化剂的固定床反应器中，在反应温度260℃、反应压力0.8MPa、重时空速1h-1、氢油体积比为200:1的条件下，在固定床反应器中连续实施200小时的芳构化/临氢异构化反应，得到第二脱硫中汽油馏分。
将上述制备的选择性加氢脱硫催化剂填装于固定床反应器中，在反应温度300℃、反应压力1.5MPa、体积空速4.0h-1、氢油体积比为600的条件下对合并有抽出油、和富硫组分的重汽油馏分进行选择性加氢脱硫，得到脱硫重汽油馏分。将脱硫重汽油馏分与脱硫醇轻汽油馏分和第二脱硫中汽油馏分混合，制得改质汽油，其组成见表4。
表4改质前后汽油的组成
Figure PCTCN2015075888-appb-000006
由表4可知:
本发明实施例5和实施例6的汽油的提质方法，不仅能够将汽油原料中的硫含量降 至10ppm以下，同时还能够将烯烃含量控制在24％以下，并且辛烷值显著提高。
实施例7
1、制备脱硫吸附剂
1)制备经碱处理的分子筛和活性炭
将两份500mL浓度为0.2mol/L的NaOH溶液水浴升温至80℃左右后，分别向其中加入25g的Y型分子筛和25g的活性炭，搅拌120分钟左右后，立即将混合物用冰浴降至常温，过滤并用去离子水过滤洗涤滤饼多次，直至滤出液的pH值为7左右后，将所得滤饼置于120℃的烘箱中干燥3h，分别制得经碱处理的Y型分子筛和经碱处理的活性炭；Y型分子筛和活性炭的比表面积及孔径分布如表5所示。
表5Y型分子筛和活性炭的比表面积及孔径
Figure PCTCN2015075888-appb-000007
2)制备第一复合载体
将上述经碱处理的Y型分子筛和经碱处理的活性炭按照质量比20:80混合后，置于研钵中研磨成粉状，随后置于110℃的烘箱中干燥6h，制得第一复合载体。
3)制备脱硫吸附剂
先采用ZnSO4溶液对上述制备的第一复合载体进行等体积浸渍，经洗涤、干燥和焙烧后，再采用Fe2(SO4)3对已浸渍ZnSO4溶液的第一复合载体进行等体积浸渍，经洗涤、干燥和焙烧后，制得脱硫吸附剂；
上述洗涤、干燥和焙烧具体为：将浸渍后的物料用去离子水洗涤后，110℃干燥24小时，将干燥后的物料冷却至室温后，先以6℃/min的速度升温至400℃，再以3℃/min的速度升温至450℃，在450℃下焙烧6小时。
上述制备的脱硫吸附剂中，Zn在第一复合载体上的负载量约为10％，Fe在第一复合载体上的负载量约为10％，并且第一复合载体上负载的Zn与Fe的质量比为1：1。经检 测，该脱硫吸附剂的硫容为0.481，寿命长达7-8h。
2、制备选择性加氢脱硫催化剂
按照实施例5方法制备选择性加氢脱硫催化剂，不同的是，控制Co在载体上的负载量约为2％，Mo在载体上的负载量约为8％，并且载体上负载的Co与Mo的质量比为0.25：1。
3、汽油改质
以济南的催化汽油为原料(其组成见表6)，对该汽油原料进行脱硫的工艺流程如图7所示。
首先，采用硫醇转化法(无碱脱臭工艺)对汽油原料进行脱硫醇处理，其中可以控制反应器操作压力为0.5MPa左右，反应温度为40℃左右，进料空速为1.0h-1，空气流量与进料量的体积比为0.5左右，收取脱硫醇汽油。
将该脱硫醇汽油切割为轻、中和重汽油馏分，其中轻、中汽油馏分的切割温度为60℃，中、重汽油馏分的切割温度为100℃。
将上述制备的脱硫吸附剂填装于固定床反应器中，在温度为30℃以及常压条件下，以0.3mL/min的流速对中汽油馏分进行吸附脱硫，得到第一脱硫中汽油馏分；并且，在吸附脱硫后，采用180℃的水蒸气吹扫吸附脱硫后的脱硫吸附剂1h进行洗涤，收取富硫组分，将该富硫组分并入重汽油馏分进行下一步骤。此外，采用400℃氮气吹扫经洗涤的脱硫吸附剂10min进行干燥，并采用室温的氮气(10℃)吹扫经干燥的脱硫吸附剂10min进行冷却，使脱硫吸附剂再生，再生三次的脱硫吸附剂的硫容为0.481，寿命达7h左右。
将第一脱硫中汽油馏分引入装有实施例5制备的芳构化/临氢异构化反应催化剂的固定床反应器中，在反应温度300℃、反应压力1MPa、重时空速2.5h-1、氢油体积比为350:1的条件下，在固定床反应器中连续实施200小时的芳构化/临氢异构化反应，得到第二脱硫中汽油馏分。
将上述制备的选择性加氢脱硫催化剂填装于固定床反应器中，在反应温度300℃、反应压力1.5MPa、体积空速4.0h-1、氢油体积比为600的条件下对合并有富硫组分的重汽油馏分进行选择性加氢脱硫，得到脱硫重汽油馏分。
将脱硫重汽油馏分与轻汽油馏分和第二脱硫中汽油馏分混合，制得改质汽油，其组成见表6。
实施例8
1、制备芳构化/临氢异构化反应催化剂
将ZSM-5分子筛和氧化铝按照质量比83:17混合，制得复合载体，其中分子筛与氧化铝的质量比为1：0.2。
采用Ga2(SO4)3·16H2O的水溶液对上述制备的复合载体进行等体积浸渍，将浸渍后的物料用去离子水洗涤后，120℃干燥18小时，将干燥后的物料冷却至室温后，先以6℃/min的速度升温至400℃，再以3℃/min的速度升温至640℃，在640℃下焙烧5小时，制得催化剂，其中Ga在复合载体上的负载量约为3％。
2、汽油改质
以济南的催化汽油为原料(其组成见表6)，对该汽油原料进行改质的工艺流程如图8所示。
首先，采用硫醇转化法(无碱脱臭工艺)对汽油原料进行脱硫醇处理，其中可以控制反应器操作压力为0.3MPa左右，反应温度为60℃左右，进料空速为1.5h-1，空气流量与进料量的体积比为1.0左右，收取脱硫醇汽油。
将该脱硫醇汽油切割为轻、中和重汽油馏分，其中轻、中汽油馏分的切割温度为50℃，中、重汽油馏分的切割温度为90℃。
将中汽油馏分从抽提塔中下部进入，环丁砜从抽提塔顶部进入，同时向抽提塔底部回流装置注入异戊烷，控制抽提塔顶温度为60℃，抽提塔底温度为40℃，抽提塔顶压(绝对压力)为0.2MPa，且将环丁砜与中汽油馏分重量比控制在1.0，异戊烷与中汽油馏分重量比控制在0.1，分别收取第一脱硫中汽油馏分和富硫油组分。
将第一脱硫中汽油馏分引入装有上述制备的芳构化/临氢异构化反应催化剂的固定床反应器中，在反应温度400℃、反应压力2MPa、重时空速6h-1、氢油体积比为800:1的条件下，在固定床反应器中连续实施200小时的芳构化/临氢异构化反应，得到第二脱硫中汽油馏分；
将实施例5制备的选择性加氢脱硫催化剂填装于固定床反应器中，在反应温度300℃、反应压力2.5MPa、体积空速2.0h-1、氢油体积比为400的条件下对合并有抽出油和富硫组分的重汽油馏分进行选择性加氢脱硫，得到脱硫重汽油馏分。将脱硫重汽油馏分与轻汽油馏分和第二脱硫中汽油馏分混合，制得改质汽油，其组成见表6。
表6改质前后汽油的组成
Figure PCTCN2015075888-appb-000008
由表6可知:
本发明实施例7和实施例8的汽油提质方法，不仅能够将汽油原料中的硫含量降至10ppm以下，同时还能够将烯烃含量控制在24％以下，并且辛烷值显著提高。
对比例1
按照实施例5方法制备经碱处理的ZSM-5型分子筛后，按照实施例5方法先后采用K2SO4溶液和NiSO4溶液对该经碱处理的ZSM-5型分子筛进行等体积浸渍，并洗涤、干燥、焙烧，制得脱硫吸附剂。经检测，该脱硫吸附剂的硫容为0.286，寿命仅为3-4h。
对比例2
按照实施例5方法制备经碱处理的活性炭后，按照实施例5方法先后采用K2SO4溶液和NiSO4溶液对该经碱处理的活性炭进行等体积浸渍，并洗涤、干燥、焙烧，制得脱硫吸附剂。经检测，该脱硫吸附剂的硫容为0.236，寿命仅为3-4h。
对比例3
直接将实施例5的ZSM-5型分子筛(未经碱处理)和活性炭(未经碱处理)按照质量比40:60混合后，置于研钵中进行研磨，随后置于120℃的烘箱中干燥6h，制得复合载体。
按照实施例5方法先后采用K2SO4溶液和NiSO4溶液对该复合载体进行等体积浸渍， 并洗涤、干燥、焙烧，制得脱硫吸附剂。经检测，该脱硫吸附剂的硫容为0.155，寿命仅为2-3h。
最后应说明的是：以上各实施例仅用以说明本发明的技术方案，而非对其限制；尽管参照前述各实施例对本发明进行了详细的说明，本领域的普通技术人员应当理解：其依然可以对前述各实施例所记载的技术方案进行修改，或者对其中部分或者全部技术特征进行等同替换；而这些修改或者替换，并不使相应技术方案的本质脱离本发明各实施例技术方案的范围。

Claims (20)

  1. 一种催化裂化汽油的提质方法，其特征在于：包括以下步骤：
    将催化裂化汽油切割为轻、中、重汽油馏分；
    在催化剂存在下对所述中汽油馏分进行芳构化/临氢异构化反应，得到脱硫中汽油馏分；
    将所述轻汽油馏分、脱硫中汽油馏分和重汽油馏分混合，获得改质汽油；
    其中，轻、中汽油馏分的切割温度为35-60℃，中、重汽油馏分的切割温度为70-160℃。
  2. 根据权利要求1所述的催化裂化汽油的提质方法，其特征在于，进行芳构化/临氢异构化反应所采用的催化剂由分子筛和金属氧化物作为复合载体负载活性金属成分而得到，其中，活性金属为锌和/或镓。
  3. 根据权利要求2所述的催化裂化汽油的提质方法，其特征在于，进行芳构化/临氢异构化反应所采用的催化剂中的分子筛选自MFI型分子筛、MCM型分子筛和LTL型分子筛中的一种或多种，金属氧化物为氧化铝。
  4. 根据权利要求2或3所述的催化裂化汽油的提质方法，其特征在于，进行芳构化/临氢异构化反应所采用的催化剂中，分子筛与金属氧化物的质量比为1：(0.2-0.5)，活性金属在复合载体上的负载量为0.5-3％。
  5. 根据权利要求1至4任一所述的催化裂化汽油的提质方法，其特征在于，所述芳构化/临氢异构化反应的反应温度为260-400℃，反应压力为0.8-2.0MPa，氢油体积比为200-800:1，重时空速为1.0-6.0h-1。
  6. 根据权利要求1至5任一所述的催化裂化汽油的提质方法，其特征在于，在对所述中汽油馏分进行芳构化/临氢异构化反应之前，先对所述中汽油馏分进行脱硫，得到第一脱硫中汽油馏分，然后在催化剂存在下对所述第一脱硫中汽油馏分进行芳构化/临氢异构化反应，得到第二脱硫中汽油馏分，再将所述轻汽油馏分、第二脱硫中汽油馏分和重汽油馏分混合，获得改质汽油。
  7. 根据权利要求6所述的催化裂化汽油的提质方法，其特征在于，对所述中汽油馏分进行的脱硫为溶剂抽提脱硫，所述溶剂抽提脱硫包括以下步骤：使所述中汽油馏分从抽提塔中下部进入，溶剂从抽提塔顶部进入，从抽提塔底部回流装置注入C5烷烃，控制抽提塔顶温度为55-100℃，抽提塔底温度为40-80℃，抽提塔顶压为0.2-0.7MPa，所述溶剂与中汽油馏分进料比控制在1.0-5.0，C5烷烃与中汽油馏分进料比控制在0.1-0.5。
  8. 根据权利要求7所述的催化裂化汽油的提质方法，其特征在于，所述溶剂选自二甘醇、三甘醇、四甘醇、二甲亚砜、环丁砜、N-甲酰吗啉、N-甲基吡咯烷酮、聚乙二醇和碳酸丙烯酯中的一种或多种。
  9. 根据权利要求6所述的催化裂化汽油的提质方法，其特征在于，对所述中汽油馏分进行的脱硫为吸附脱硫，利用脱硫吸附剂进行所述吸附脱硫，所述脱硫吸附剂由分别经碱处理的分子筛和活性炭作为复合载体负载活性金属成分而得到，其中，活性金属选自周期表IA、VIII、IB、IIB和VIB族中的一种或多种元素。
  10. 根据权利要求9所述的催化裂化汽油的提质方法，其特征在于，所述脱硫吸附剂的复合载体中，分子筛与活性炭的质量比为(20-80):(80-20)。
  11. 根据权利要求9或10所述的催化裂化汽油的提质方法，其特征在于，所述脱硫吸附剂的复合载体中的分子筛的类型为X型、Y型或ZSM-5型。
  12. 根据权利要求9至11任一所述的催化裂化汽油的提质方法，其特征在于，所述脱硫吸附剂中的活性金属选自Ni、Fe、Ag、Co、Mo、Zn和K中的至少2种。
  13. 根据权利要求9至12任一所述的催化裂化汽油的提质方法，其特征在于，所述脱硫吸附剂中的活性金属在复合载体上的负载量为2-30％。
  14. 根据权利要求9至13任一所述的催化裂化汽油的提质方法，其特征在于，所述吸附脱硫是利用固定床常压进行，并且控制吸附脱硫的温度为20-100℃，中汽油馏分的流速为0.3-1mL/min。
  15. 根据权利要求1至14任一所述的催化裂化汽油的提质方法，其特征在于，在将催化裂化汽油切割为轻、中、重汽油馏分之前，先对催化裂化汽油进行脱硫醇处理。
  16. 根据权利要求1至14任一所述的催化裂化汽油的提质方法，其特征在于，在将所述轻汽油馏分、脱硫中汽油馏分和重汽油馏分混合之前，先对所述轻汽油馏分进行脱硫醇处理，得到脱硫醇轻汽油馏分，然后将所述脱硫醇轻汽油馏分与脱硫中汽油馏分和重汽油馏分混合，获得改质汽油。
  17. 根据权利要求1至16任一所述的催化裂化汽油的提质方法，其特征在于，在将所述轻汽油馏分、脱硫中汽油馏分和重汽油馏分混合之前，先对所述重汽油馏分进行选择性加氢脱硫，得到脱硫重汽油馏分，然后将所述脱硫重汽油馏分与轻汽油馏分和脱硫中汽油馏分混合，获得改质汽油。
  18. 根据权利要求17所述的催化裂化汽油的提质方法，其特征在于，将所述重汽油 馏分、氢气在选择性加氢脱硫催化剂的作用下进行选择性加氢脱硫，得到脱硫重汽油馏分，其中，所述选择性加氢脱硫的温度为200-300℃，压力为1.5-2.5MPa，体积空速为1-5h-1，氢油体积比为400-600。
  19. 根据权利要求18所述的催化裂化汽油的提质方法，其特征在于，所述选择性加氢脱硫催化剂由载体负载活性金属成分而得到，其中，载体为分子筛或金属氧化物，活性金属包括Co和Mo。
  20. 根据权利要求19所述的催化裂化汽油的提质方法，其特征在于，Co和Mo在载体上的总负载量为5-20％。
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Cited By (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN111135859A (zh) * 2020-01-13 2020-05-12 中化泉州石化有限公司 一种长链正构烷烃临氢异构化催化剂及其制备方法
CN111575045A (zh) * 2020-05-26 2020-08-25 中国海洋石油集团有限公司 一种脱硫汽油降苯增产芳烃的方法

Citations (7)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US5294328A (en) * 1990-05-24 1994-03-15 Uop Production of reformulated gasoline
US6126812A (en) * 1998-07-14 2000-10-03 Phillips Petroleum Company Gasoline upgrade with split feed
CN1464033A (zh) * 2002-06-13 2003-12-31 中国科学院大连化学物理研究所 一种催化裂化汽油的芳构化加氢异构化脱硫改性的工艺
CN1782034A (zh) * 2004-11-30 2006-06-07 中国石油化工股份有限公司 一种同时降低汽油硫、烯烃含量的方法
CN101283074A (zh) * 2005-08-08 2008-10-08 法国石油公司 包括轻馏分吸附脱硫和重馏分加氢脱硫的汽油脱硫的方法
CN101418234A (zh) * 2007-10-26 2009-04-29 中国石油化工股份有限公司 一种催化裂化汽油生产低硫汽油的方法
CN103555359A (zh) * 2013-11-18 2014-02-05 郝天臻 一种催化裂化汽油的深度脱硫方法

Patent Citations (7)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US5294328A (en) * 1990-05-24 1994-03-15 Uop Production of reformulated gasoline
US6126812A (en) * 1998-07-14 2000-10-03 Phillips Petroleum Company Gasoline upgrade with split feed
CN1464033A (zh) * 2002-06-13 2003-12-31 中国科学院大连化学物理研究所 一种催化裂化汽油的芳构化加氢异构化脱硫改性的工艺
CN1782034A (zh) * 2004-11-30 2006-06-07 中国石油化工股份有限公司 一种同时降低汽油硫、烯烃含量的方法
CN101283074A (zh) * 2005-08-08 2008-10-08 法国石油公司 包括轻馏分吸附脱硫和重馏分加氢脱硫的汽油脱硫的方法
CN101418234A (zh) * 2007-10-26 2009-04-29 中国石油化工股份有限公司 一种催化裂化汽油生产低硫汽油的方法
CN103555359A (zh) * 2013-11-18 2014-02-05 郝天臻 一种催化裂化汽油的深度脱硫方法

Cited By (4)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN111135859A (zh) * 2020-01-13 2020-05-12 中化泉州石化有限公司 一种长链正构烷烃临氢异构化催化剂及其制备方法
CN111135859B (zh) * 2020-01-13 2023-01-10 中化泉州石化有限公司 一种长链正构烷烃临氢异构化催化剂及其制备方法
CN111575045A (zh) * 2020-05-26 2020-08-25 中国海洋石油集团有限公司 一种脱硫汽油降苯增产芳烃的方法
CN111575045B (zh) * 2020-05-26 2022-04-05 中国海洋石油集团有限公司 一种脱硫汽油降苯增产芳烃的方法

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