WO2018072343A1 - 一种汽油的处理方法 - Google Patents

一种汽油的处理方法 Download PDF

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Publication number
WO2018072343A1
WO2018072343A1 PCT/CN2017/000634 CN2017000634W WO2018072343A1 WO 2018072343 A1 WO2018072343 A1 WO 2018072343A1 CN 2017000634 W CN2017000634 W CN 2017000634W WO 2018072343 A1 WO2018072343 A1 WO 2018072343A1
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Prior art keywords
gasoline
catalyst
aromatization
desulfurization
fraction
Prior art date
Legal status (The legal status is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the status listed.)
Ceased
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PCT/CN2017/000634
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English (en)
French (fr)
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WO2018072343A8 (zh
Inventor
许友好
罗一斌
王新
欧阳颖
达志坚
舒兴田
汪燮卿
Current Assignee (The listed assignees may be inaccurate. Google has not performed a legal analysis and makes no representation or warranty as to the accuracy of the list.)
Sinopec Research Institute of Petroleum Processing
China Petroleum and Chemical Corp
Original Assignee
Sinopec Research Institute of Petroleum Processing
China Petroleum and Chemical Corp
Priority date (The priority date is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the date listed.)
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Publication date
Priority claimed from CN201610921999.2A external-priority patent/CN107974296B/zh
Priority claimed from CN201610921955.XA external-priority patent/CN107974290B/zh
Priority claimed from CN201610921806.3A external-priority patent/CN107974293B/zh
Priority claimed from CN201610922006.3A external-priority patent/CN107974292B/zh
Priority claimed from CN201610921175.5A external-priority patent/CN107974297B/zh
Priority to KR1020197011089A priority Critical patent/KR102485001B1/ko
Priority to CN201780064852.9A priority patent/CN110139917B/zh
Priority to SG11201903075WA priority patent/SG11201903075WA/en
Application filed by Sinopec Research Institute of Petroleum Processing , China Petroleum and Chemical Corp filed Critical Sinopec Research Institute of Petroleum Processing
Priority to US16/339,670 priority patent/US10907110B2/en
Priority to RU2019115348A priority patent/RU2742646C2/ru
Publication of WO2018072343A1 publication Critical patent/WO2018072343A1/zh
Publication of WO2018072343A8 publication Critical patent/WO2018072343A8/zh
Anticipated expiration legal-status Critical
Ceased legal-status Critical Current

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    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/30—Aromatics

Definitions

  • the invention relates to a method for treating gasoline, and more particularly to a method for desulfurizing gasoline.
  • Catalytic cracking gasoline is a major component of motor gasoline, accounting for about 75% in gasoline pools, and is characterized by high olefin and sulfur content. It is not difficult to achieve deep desulfurization of gasoline and reduce the content of olefins in catalytic cracking gasoline by hydrogenation technology. However, since olefin is a high octane component, a large decrease in its content will lead to a serious loss of gasoline octane number, thus affecting The performance of gasoline and the economic benefits of refineries, so the deep desulfurization of gasoline while maintaining the gasoline octane number has become a hot spot in the production of clean gasoline.
  • the deep desulfurization of gasoline is mainly by hydrodesulfurization or adsorption desulfurization.
  • Selective hydrodesulfurization is one of the main ways to remove thiophene sulfides, but the olefin saturation and other reactions also occur in large quantities, resulting in a large loss of octane number.
  • a deep hydrogenation process to restore octane number is also recognized as a second reactor to promote low-octane hydrocarbon cracking, isomerization and alkylation while performing deep desulfurization and olefin saturation. The reaction is achieved to achieve the purpose of restoring the octane number.
  • Chinese patent application CN101845322A discloses a method for reducing the content of sulfur and olefin in gasoline.
  • the raw material catalytic cracking gasoline is first removed from the diene by a prehydrogenation reactor, and then cut into fractional distillation towers to be fractionated into light and heavy gasoline, and light gasoline is hydrogen.
  • heavy gasoline enters the selective hydrogenation reactor for hydrodesulfurization, and the reaction effluent enters the hydro-upgrading reactor for hydro-upgrading, reducing the olefin content, and the modified heavy gasoline and light gasoline adsorbed.
  • the desulfurization products are blended to obtain clean gasoline that meets the standard requirements.
  • the adsorption desulfurization catalyst has a good removal effect on the sulfide in the gasoline, but the adsorption desulfurization is carried out in the presence of hydrogen, and the olefin in the saturated catalytic cracking gasoline, especially the light gasoline, is adsorbed and desulfurized, and the olefin group in the light gasoline. A higher octane number will result in a large loss of the octane number of the gasoline.
  • the adsorption method removes the sulfur-containing compound in the fuel, and uses the adsorbent to react and adsorb the light oil in the presence of hydrogen to form a metal sulfide or to remove sulfur by using a sulfide, the hydrogen consumption is low, and the desulfurization efficiency is high. It is possible to produce gasoline having a sulfur content of 10 ⁇ g/g or less. Although the adsorption process achieves deep desulfurization of gasoline under low hydrogen consumption conditions, the gasoline product octane number is still slightly lost. Especially when dealing with gasoline raw materials with high olefin content and high sulfur content, the loss of gasoline octane number is still large.
  • the present invention provides a method of treating gasoline, comprising:
  • the mixed catalyst comprises an adsorbed desulfurization catalyst and an aromatization catalyst, and said at least about 50 wt%, preferably at least about 80 wt%, more preferably at least about 90 wt%, particularly preferably at least about 95 wt%, most preferably about 100 wt%
  • the aromatization catalyst is passivated and/or aged.
  • the present invention provides a method of treating gasoline comprising:
  • the mixed catalyst comprises an adsorption desulfurization catalyst and an aromatization catalyst
  • the aromatization catalyst is selected from the group consisting of a fresh aromatization catalyst, a passivated aromatization catalyst, an aged aromatization catalyst, Or any combination of them.
  • the present invention provides a gasoline treatment system comprising a desulfurization-aromatization reactor, a high pressure separator, a gasoline full fraction fractionation column, and an etherification unit, wherein the desulfurization-aromatization reactor Provided with a gasoline feedstock inlet, a hydrogen inlet, and a desulfurization-aromatization product outlet, the gasoline full fraction fractionator is provided with a gasoline full fraction inlet, a light gasoline fraction outlet, and a heavy gasoline fraction outlet; the desulfurization-aromatization reaction The desulfurization-aromatization product outlet of the apparatus is in fluid communication with the gasoline full fraction inlet of the gasoline full fraction fractionation column via the high pressure separator; the light gasoline fraction outlet of the gasoline full fractionation column and the inlet fluid of the etherification unit Connected,
  • the gasoline feedstock enters the desulfurization-aromatization reactor through the inlet of the gasoline feedstock, wherein desulfurization and aromatization reaction are carried out in the presence of hydrogen to obtain a desulfurization-aromatization product;
  • the desulfurization-aromatization product Dividing a hydrogen-containing tail gas and a gasoline full fraction in the high-pressure separator, the whole gasoline fraction enters the gasoline full-fraction fractionation column through a gasoline full-fraction inlet, and fractionally distilling to obtain a light gasoline fraction and a heavy gasoline fraction;
  • the gasoline fraction enters the etherification unit where it undergoes an etherification reaction to give an etherified product.
  • the method of the present invention desulfurizes and aromatizes a high sulfur and high olefin gasoline feedstock with an adsorbed desulfurization catalyst and an aromatization catalyst, and can cleave and aromatize olefins in gasoline while reducing gasoline sulfur content.
  • the production of hydrocarbons below C 5 and preferably cutting in the light gasoline fraction for etherification reaction reduces both the olefin content of the gasoline and the octane number of the gasoline and maintains a high yield of the gasoline feedstock.
  • the desulfurization and aromatization reactions of the present invention can be carried out using existing adsorption desulfurization reactors without the need to modify existing reactors.
  • the desulfurization and aromatization reaction of the present invention are carried out in one reactor by using two kinds of catalysts, and the stepwise treatment of gasoline (aromatization after desulfurization of gasoline or desulfurization after aromatization of gasoline) needs to be separately increased.
  • the problem of aromatization reactors and ancillary systems can also avoid the need to change the existing catalyst preparation process for gasoline desulfurization and aromatization coupling catalysts.
  • the problem of low catalyst wear strength improves the reaction efficiency and reduces the investment cost.
  • the aromatization catalyst of the present invention is preferably subjected to passivation and/or aging treatment to impart a moderate activity to the aromatization catalyst, thereby facilitating the progress of aromatization.
  • the present invention preferably etherifies the light gasoline fraction to reduce both olefins and high-octane etherified oils, thereby increasing the octane number of the gasoline product.
  • the etherification treatment preferably employed in the present invention is also capable of reducing light components in gasoline products and reducing gasoline vapor pressure.
  • Figure 1 is a schematic illustration of a preferred embodiment of the method and system of the present invention.
  • any specific numerical values (including the endpoints of the numerical ranges) disclosed herein are not limited to the precise value of the numerical value, but should be understood to cover the value close to the precise value. And, for the range of values disclosed, one or more new ranges of values can be obtained by arbitrarily combining between the endpoint values of the range, the endpoint values and the particular point values in the range, and the specific point values. These new numerical ranges are also considered to be specifically disclosed herein.
  • the RIPP test method in the present invention can be found in the "Petrochemical Analysis Method", edited by Yang Cuiding, Science Press, September 1990, pages 263-268 and 303-304, ISBN: 7-03-001894-X.
  • the present invention provides a method of treating gasoline comprising:
  • the mixed catalyst comprises an adsorbed desulfurization catalyst and an aromatization catalyst, and said at least about 50 wt%, preferably at least about 80 wt%, more preferably at least about 90 wt%, particularly preferably at least about 95 wt%, most preferably about 100 wt%
  • the aromatization catalyst is passivated and/or aged.
  • the process further comprising: at least a portion of the resulting desulfurized - aromatization product is distilled to give a liquid C 3 -C 4 hydrocarbon components and stable gasoline fractions.
  • the liquid hydrocarbon component is sent to a gas separation unit for further processing.
  • the stabilized gasoline fraction can be used directly as a gasoline product or as a feedstock for other processes.
  • the processing method further includes the following steps:
  • processing method further includes the following steps:
  • At least a portion of the etherified oil is combined with at least a portion of the heavy gasoline fraction to provide a gasoline product.
  • the present invention provides a method of treating gasoline comprising:
  • the mixed catalyst comprises an adsorption desulfurization catalyst and an aromatization catalyst
  • the aromatization catalyst is selected from the group consisting of a fresh aromatization catalyst, a passivated aromatization catalyst, an aged aromatization catalyst, Or any combination of them.
  • the "desulfurization and aromatization reaction” refers to a process in which a gasoline feedstock is subjected to desulfurization and conversion of an olefin to an aromatic hydrocarbon by the action of an adsorption desulfurization catalyst and an aromatization catalyst in the presence of hydrogen, which may be accompanied by Cracking reaction.
  • the desulfurization and aromatization reactions are carried out under conditions such as a reaction temperature of from about 350 to 500 ° C, preferably from about 380 to 420 ° C, and a weight hourly space velocity of from about 2 to about 50 hours - 1 , preferably about 5-20 hours -1 , a reaction pressure of about 0.5-3.0 MPa, preferably about 1.5-2.5 MPa, a volume ratio of hydrogen to gasoline raw materials (in standard conditions (STP) 0 ° C (273 K), 1.01 ⁇ 10 5 Pa) is about 1-500, preferably about 50-200.
  • a reaction temperature of from about 350 to 500 ° C, preferably from about 380 to 420 ° C
  • a weight hourly space velocity of from about 2 to about 50 hours - 1 , preferably about 5-20 hours -1
  • a reaction pressure of about 0.5-3.0 MPa, preferably about 1.5-2.5 MPa
  • a volume ratio of hydrogen to gasoline raw materials in standard conditions (STP) 0 ° C (273 K
  • the adsorptive desulfurization catalyst used in the present invention is not particularly limited, and may be any catalyst suitable for adsorption desulfurization of gasoline which is well known to those skilled in the art.
  • the adsorptive desulfurization catalyst comprises silica, alumina, zinc oxide, and a desulfurization active metal, and the desulfurization active metal is selected from the group consisting of cobalt, nickel, copper, iron, manganese, molybdenum, and tungsten. At least one of silver, tin and vanadium.
  • the zinc oxide comprises about 10 to 90% by weight, based on the dry basis weight of the adsorptive desulfurization catalyst, and based on the weight of the oxide.
  • About 5 to 85% by weight, alumina accounts for about 5 to 30% by weight; based on the dry basis weight of the adsorbed desulfurization catalyst and based on the weight of the element, the content of the desulfurization active metal in the adsorptive desulfurization catalyst is about 5-30% by weight.
  • the adsorptive desulfurization catalyst may further contain from 1 to 10% by weight of a coke-like substance.
  • Industrial practice shows that the carbon content of the adsorption desulfurization catalyst has an effect on the desulfurization efficiency of the adsorption desulfurization catalyst and the loss of gasoline octane. With the increase of carbon content of the adsorption desulfurization catalyst, the desulfurization efficiency of the adsorption desulfurization catalyst is gradually decreased, and the gasoline octane number is gradually decreased. The loss is reduced.
  • the adsorptive desulfurization catalyst maintains a certain sulfur content.
  • the sulfur-carrying capacity of the adsorbed desulfurization catalyst is about 9-10% by weight, and the sulfur content of the regenerated adsorption desulfurization catalyst is about 5-6 wt%, and the adsorption desulfurization catalyst and regeneration absorption
  • the sulfur difference of the desulfurization catalyst is preferably about 4% by weight or so.
  • the "aromatization catalyst” means a catalyst capable of converting a hydrocarbon such as an olefin or the like in a gasoline feedstock to an aromatic hydrocarbon, and generally includes a molecular sieve, preferably including a molecular sieve, a support, and a metal.
  • the aromatization catalyst comprises from about 10% to about 30% by weight molecular sieves, from about 0% to about 20% by weight, based on the total weight of the aromatization catalyst. Aromatization of the active metal oxide and from about 50 to 90% by weight of the support.
  • the molecular sieve comprises a Y molecular sieve and/or an MFI structural molecular sieve, preferably a five-membered high silica molecular sieve, which may be of the hydrogen type or may be a rare earth and/or Phosphorus modified, preferably having a silicon to aluminum ratio greater than 100, more preferably greater than 150.
  • the aromatization active metal may function as a partial desulfurization or hydrocarbon conversion, for example, may be a metal element selected from Group IVB, a metal element of Group VB, Group VIB At least one of a metal element, a metal element of Group VIII, a metal element of Group IB, a metal element of Group IIB, and a metal element of Group IIIA.
  • the metal element of Group IVB is Zr or/and Ti
  • the metal element of Group VB is V
  • the metal element of Group VIB is Mo or/and W
  • Group VIII The metal element is one or more selected from the group consisting of Fe, Co, and Ni
  • the metal element of Group IB is Cu
  • the metal element of Group IIB is Zn
  • the metal element of Group IIIA is Ga.
  • the aromatization active metal is at least one selected from the group consisting of Fe, Zn and Ga, and the content is preferably from 0.5 to 5% by weight.
  • the support preferably comprises silica and/or alumina.
  • the aromatization catalyst typically has a particle size of from 20 to 120 microns, which is comparable to the particle size of the adsorptive desulfurization catalyst.
  • the present invention mixes the adsorbed desulfurization catalyst and the aromatization catalyst separately after molding (for example, spray drying).
  • the ratio of the adsorbed desulfurization catalyst to the aromatization catalyst in the mixed catalyst of the present invention may vary depending on the olefin and sulfur content in the gasoline feedstock. In certain preferred embodiments, the ratio of the aromatization catalyst to the mixed catalyst is from about 1% to about 30% by weight, preferably from about 3% to about 15% by weight.
  • the aromatization catalyst of the present invention is prepared by mixing and mixing the preparation raw material of the aromatization catalyst and water, and then spray-drying and calcining the slurry; Weight and dry with the preparation of the raw materials Based on the basis weight, the preparation material comprises about 15 to 60% by weight of natural minerals, about 10 to 30% by weight of a precursor of an inorganic oxide binder, and about 20 to 80% by weight of phosphorus and metal-loaded MFI structural molecular sieve.
  • the preparation material comprises, by weight and on a dry basis weight of the preparation material, about 20 to 55% by weight of natural minerals, and about 12 to 28% by weight of inorganic oxidation.
  • the "natural mineral” refers to a natural element or compound formed under the combined action of various substances in the earth's crust (called geological action), and has a characteristic chemical composition and a relatively fixed chemical composition.
  • geological action a natural element or compound formed under the combined action of various substances in the earth's crust
  • at least one selected from the group consisting of kaolin, halloysite, montmorillonite, diatomaceous earth, attapulgite, sepiolite, halloysite, hydrotalcite, bentonite, and rector soil may be included.
  • the "inorganic oxide binder” means an inorganic oxide which functions as a binder in the catalyst, and may, for example, be selected from the group consisting of silica, alumina, zirconia, titania and At least one of the shaped silicon aluminum.
  • the "precursor of an inorganic oxide binder” means a catalytic cracking catalyst preparation material generally used for producing an inorganic oxide binder in a catalytic cracking catalyst, and may, for example, be selected from silica sol, aluminum. At least one of a sol, a peptized pseudoboehmite, a silica alumina sol, and a phosphorus-containing aluminum sol.
  • the "supporting metal” means a metal supported on a molecular sieve by a load, excluding aluminum and an alkali metal such as sodium or potassium.
  • the supporting metal used in the present invention is not particularly limited and may be zinc, gallium and/or iron, and may also include other metals.
  • the supported metal is zinc and/or gallium
  • the natural mineral comprises a component selected from the group consisting of kaolin, halloysite, montmorillonite, diatomaceous earth, convex louver, sepiolite.
  • the inorganic oxide binder comprising at least one selected from the group consisting of silica, alumina, zirconia, titania, and amorphous silica-aluminakind.
  • the MFI structural molecular sieve is at least one selected from the group consisting of ZSM-5, ZSM-8, and ZSM-11.
  • the MFI structural molecular sieve of the aromatization catalyst has n(SiO 2 )/n(Al 2 O 3 ) greater than about 100; in terms of P 2 O 5 and dried by molecular sieves Based on the basis weight, the molecular sieve has a phosphorus content of about 0.1 to 5% by weight; the molecular weight of the molecular sieve is from about 0.5 to 5% by weight based on the oxide of the supported metal and based on the dry weight of the molecular sieve.
  • the crystal grain edge of the molecular sieve crystal grain is inwardly located within an H distance of more than 100 square nanometers.
  • the content, M(C) represents the loading metal content of the geometric center of the crystal plane of the molecular sieve grain determined by the TEM-EDS method and the area outside the H distance is more than 100 square nanometers; the mesopore volume of the molecular sieve accounts for the total pore volume.
  • the volume ratio is about 40-80% by volume, and the mesopore volume having a pore diameter of about 2 nm to about 20 nm preferably accounts for more than about 90% by volume of the total mesopore volume; the strong acid amount of the molecular sieve accounts for the total acid amount.
  • the ratio is about 60-80%, and the ratio of the amount of Bronsted acid (B acid) acid to the amount of Lewis acid (L acid) is about 15-80.
  • the MFI structural molecular sieve has n(SiO 2 )/n(Al 2 O 3 ) greater than about 120; based on P 2 O 5 and based on the dry basis weight of the molecular sieve, The molecular sieve has a phosphorus content of about 0.2 to 4% by weight; based on the metal-supporting oxide and based on the dry weight of the molecular sieve, the molecular weight of the molecular sieve is from about 0.5 to 3% by weight; the molecular sieve of Al
  • the distribution parameter D(Al) satisfies: 0.65 ⁇ D(Al) ⁇ 0.8; the supported metal distribution parameter D(M) of the molecular sieve satisfies: 3 ⁇ D(M) ⁇ 6;
  • the mesopore volume of the molecular sieve accounts for the total pore
  • the ratio of the volume is about 50-70% by volume, the ratio of the mesopore volume having a pore diameter of about 2 nm to about 20
  • the TEM-EDS method for determining the aluminum content and the supported metal content of the molecular sieve is well known to those skilled in the art, wherein the determination of the geometric center is also well known to those skilled in the art and can be calculated according to the formula. The invention will not be described again.
  • the geometric center of a general symmetrical figure is the intersection of the lines connecting the opposite vertices, for example, a regular hexagonal shape
  • the geometric center of the hexagonal crystal face of the ZSM-5 molecular sieve is at the intersection of three lines of opposite vertices, which are one face of regular grains, both of which are inward and outward directions The inward and outward directions on the facet.
  • the ratio of the mesopore volume of the molecular sieve to the total pore volume is measured by a nitrogen adsorption BET specific surface area method, wherein the mesopore volume is a pore volume having a pore diameter greater than about 2 nm and less than about 100 nm;
  • the ratio of the amount of strong acid to the total amount of acid is measured by the NH 3 -TPD method, and the acid center of the strong acid is the acid center corresponding to the NH 3 desorption temperature greater than 300 ° C; the amount of B acid and the amount of L acid
  • the ratio is measured by the pyridine adsorption infrared acid method.
  • the phosphorus-containing and metal-loaded MFI molecular sieves can be prepared by the following steps:
  • step b the water-washed molecular sieve obtained in step a is desiliconized in an alkali solution, and after filtration and washing, a desiliconized molecular sieve is obtained;
  • the desiliconized molecular sieve obtained in step b is subjected to ammonium exchange treatment to obtain an ammonium exchange molecular sieve; wherein the ammonium exchange molecular sieve has a sodium content of less than 0.2 by weight based on the total dry basis weight of the ammonium exchange molecular sieve. %;
  • step c the ammonium exchange molecular sieve obtained in step c is subjected to dealumination treatment in a composite acid dealuminant solution consisting of fluorosilicic acid, organic acid and inorganic acid, and after filtration and washing, a de-aluminated molecular sieve is obtained;
  • the modified molecular sieve obtained in the step e is hydrothermally calcined to obtain the phosphorus-containing and metal-loaded MFI molecular sieve.
  • the "crystallized MFI structure molecular sieve slurry" used in the step a of the present invention can be obtained according to a method well known to those skilled in the art, and will not be described herein.
  • the MFI structural molecular sieve is also well known to those skilled in the art, and may be obtained by amine-free crystallization, or may be a molecular sieve prepared by a templating method, wherein the molecular sieve obtained without amine crystallization does not need to be calcined, and the templating agent is used.
  • the molecular sieve prepared by the method needs to be dried and then calcined in air, and the silicon-aluminum ratio of the ZSM-5 molecular sieve is generally less than 100.
  • the desiliconization treatment using the alkali solution in the step b of the present invention is well known to those skilled in the art.
  • the alkali solution in the step b may be selected from the group consisting of sodium hydroxide solution and/or potassium hydroxide solution, preferably hydrogen.
  • the sodium oxide solution, the conditions of the desiliconization treatment may include: the weight ratio of the molecular sieve to the alkali in the alkali solution on a dry basis weight is about 1: (0.1-1), preferably about 1: (0.15-0.4)
  • the desiliconization treatment is carried out at a temperature of from room temperature to about 100 ° C, preferably from about 50 to 85 ° C, for a period of from about 15 minutes to about 8 hours, preferably from about 30 minutes to about 4 hours.
  • the alkali-treated desiliconized molecular sieve can be about 1: (0.1-1) according to molecular sieve: ammonium salt: H 2 O.
  • the weight ratio of (5-10) is exchanged at room temperature to about 100 ° C for about 0.5-2 hours and then filtered so that the Na 2 O content on the zeolite is less than about 0.2% by weight.
  • the ammonium salt may be a commonly used inorganic ammonium salt, for example, at least one selected from the group consisting of ammonium chloride, ammonium sulfate, and ammonium nitrate.
  • the " dealumination treatment" referred to in the step d of the present invention is known per se to those skilled in the art, but it has not been reported that the inorganic acid, the organic acid and the fluorosilicic acid are used together in the dealumination treatment as in the step d of the present invention.
  • the dealumination treatment may be carried out once or in multiple times, and the organic acid may be first mixed with the ammonium exchange molecular sieve, and then the fluorosilicic acid and the inorganic acid are mixed with the ammonium exchange molecular sieve, that is, the organic acid may be first added.
  • the fluorosilicic acid and the inorganic acid are then slowly co-currently added, or the fluorosilicic acid is added first and then the inorganic acid is added, preferably the fluorosilicic acid and the inorganic acid are added in a slow cocurrent manner.
  • the dealumination treatment can be carried out under the following conditions: the weight ratio of molecular sieve, fluorosilicic acid, organic acid and inorganic acid on a dry basis weight is about 1: (0.02-0.5): (0.05-0.5): 0.05-0.5), preferably about 1: (0.05-0.3): (0.1-0.3): (0.1-0.3); the treatment temperature is about 25-100 ° C, and the treatment time is about 0.5-6 hours.
  • the experiment proves that the invention adopts the composite acid system for dealumination treatment, and under the synergistic action of the three acids, the silicon-aluminum ratio of the molecular sieve can be effectively improved and the aluminum distribution can be adjusted under the premise of ensuring the crystal structure of the molecular sieve and the structural integrity of the mesoporous pore. And improve the acid distribution.
  • the organic acid and inorganic acid used in the step d of the present invention may be those conventionally used in the art, for example, the organic acid may be at least selected from the group consisting of ethylenediaminetetraacetic acid, oxalic acid, citric acid, and sulfosalicylic acid.
  • the organic acid may be at least selected from the group consisting of ethylenediaminetetraacetic acid, oxalic acid, citric acid, and sulfosalicylic acid.
  • One is preferably oxalic acid
  • the inorganic acid may be at least one selected from the group consisting of hydrochloric acid, sulfuric acid and nitric acid, preferably hydrochloric acid.
  • step d of the present invention is well known to those skilled in the art, for example, by rinsing about 5-10 times water at about 30-60 ° C against filtered molecular sieves.
  • the phosphorus modification treatment and the loading treatment of the supported metal carried out in the step e of the present invention are well known to those skilled in the art.
  • the phosphorus modification treatment in the step e may include: selecting from phosphoric acid, ammonium hydrogen phosphate, phosphoric acid.
  • the at least one phosphorus-containing compound of ammonium dihydrogen phosphate and ammonium phosphate impregnates and/or ion-exchanges the molecular sieve;
  • the loading treatment of the load metal in step e may include: containing at least one supported metal selected from the group consisting of zinc and gallium
  • the soluble salt is dissolved in deionized water, the pH is adjusted with ammonia water to precipitate the supported metal in the form of hydroxide, and the resulting precipitate is uniformly mixed with the molecular sieve.
  • the hydrothermal calcination treatment of the molecular sieve carried out in the step f of the present invention is well known to those skilled in the art.
  • the hydrothermal calcination treatment in the step f can be carried out under the following conditions: the atmosphere of the calcination treatment is a water vapor atmosphere; roasting The temperature is about 400-800 ° C and the calcination time is about 0.5-8 hours.
  • the inventors of the present invention have surprisingly discovered that the acrobatic activity of the aromatization catalyst has a better effect at about 20-55, which uses RIPP 92-90.
  • the micro-reaction activity test method of the catalytic cracking industrial equilibrium catalyst is determined, and the specific measurement process can be referred to the examples.
  • the acceptable fresh aromatization catalysts commercially available from the manufacturer or made by the manufacturer are generally above 60, with higher activity and stronger cracking performance. Therefore, prior to use in the desulfurization and aromatization reactions, such fresh aromatization catalysts are preferably treated to reduce the amount of acid and increase the acid strength, thereby facilitating the reduction of the hydrogen transfer reaction.
  • the aromatization catalyst is passivated prior to use in the desulfurization and aromatization reactions.
  • the passivation treatment comprises contacting an aromatization catalyst, such as a fresh aromatization catalyst, with a carbon, sulfur and/or nitrogen containing compound and subjecting the passivation reaction.
  • the passivated treatment of the aromatization catalyst comprises from about 0.1 to 5.0% by weight of a passivation, wherein the passivation comprises at least one selected from the group consisting of carbon, sulfur, and nitrogen. element.
  • the passivation treatment can be carried out in a reactor located outside of the desulfurization-aromatization reactor, or in a fluidization reactor used as a desulfurization-aromatization reactor.
  • the pre-elevation section is carried out;
  • the carbon, sulfur and/or nitrogen-containing compound used for the passivation treatment may be gasoline, hydrogen sulfide, carbon disulfide, ammonia, aniline, pyridine or quinoline, etc.
  • the gasoline feedstock may also be other gasoline from the outside, such as catalytically cracked gasoline, steam cracked gasoline or other olefin-containing gasoline.
  • the aromatization catalyst is subjected to an aging treatment prior to use in the desulfurization and aromatization reactions.
  • the "aging treatment” means that an aromatization catalyst such as a fresh aromatization catalyst is subjected to high temperature treatment in the presence of steam to lower its activity.
  • the aging treatment is carried out under conditions of a temperature of from about 500 to 800 ° C, preferably from about 600 to 800 ° C, more preferably from about 700 to 800 ° C; and a time of from about 1 to 360 hours Preferably, it is about 2 to 48 hours, more preferably about 4 to 24 hours; the aging atmosphere is an atmosphere containing water vapor, preferably about 100% of a water vapor atmosphere.
  • the fresh aromatization catalyst is contacted with an aging medium of steam or water vapor, and aged for about 1 hour in a hydrothermal environment at a temperature of about 500 ° C to 800 ° C. After 360 hours, an aged aromatization catalyst was obtained.
  • the aging medium comprises air, dry gas, regenerated flue gas, air or dry gas burned gas or air and combustion oil burned gas, or other gases such as nitrogen.
  • the weight ratio of water vapor to aged medium in the aqueous vapor-containing aging medium is from about 0.2 to 0.9, more preferably from about 0.40 to 0.60.
  • the regenerated flue gas may be from the catalyst regenerating device of the present invention or may be derived from a regenerative device of other processes.
  • the aging treatment is carried out in an aged reactor, preferably a dense phase fluidized bed.
  • the aging treatment comprises aging the fresh aromatization catalyst at 800 ° C and 100% steam for a period of from 4 h to 17 h.
  • At least about 50 wt%, preferably at least about 80 wt%, more preferably at least about 90 wt%, particularly preferably at least about 95 wt%, and most preferably about 100 wt% of the aromatization catalyst is passivated and / or aging treatment.
  • the desulfurization and aromatization reactions of the present invention are preferably carried out in a fluidized reactor to facilitate rapid regeneration of the aromatization catalyst.
  • the fluidization reactor is selected from the group consisting of a fluidized bed, a riser, a down-conveying line reactor, a composite reactor composed of a riser and a fluidized bed, and a riser and a composite reactor composed of a down-conveying line, a composite reactor composed of two or more risers, a composite reactor composed of two or more fluidized beds, and two or more
  • the composite reactor constituted by the descending conveying line is preferably a riser reactor, a fluidized bed reactor or a combination thereof.
  • each of the above reactors may be divided into two or more reaction zones.
  • the fluidized bed reactor is one or more selected from the group consisting of a fixed fluidized bed, a bulk fluidized bed, a bubbling bed, a turbulent bed, a fast bed, a transport bed, and a dense phase fluidized bed.
  • the riser reactor is one or more selected from the group consisting of an equal diameter riser, a constant line riser, and various variable diameter risers.
  • the fluidization reactor is a dense phase fluidization reactor.
  • the gasoline feedstock used in the present invention may be any conventional gasoline feedstock commonly used in the art.
  • the gasoline feedstock may be at least one selected from the group consisting of catalytically cracked gasoline, catalytically cracked gasoline, cokered gasoline, thermally cracked gasoline, and straight run gasoline, or a fraction thereof.
  • the gasoline feedstock is a high olefin and a high sulfur gasoline having an olefin volume fraction generally greater than about 20% by volume, preferably greater than about 30% by volume, more preferably greater than about 40% by volume, and even more preferably greater than About 50% by volume;
  • the sulfur content is generally above about 10 ⁇ g/g, preferably greater than about 50 ⁇ g/g, more preferably greater than about 100 ⁇ g/g, further preferably greater than about 500 ⁇ g/g, and even more preferably greater than about 1000 ⁇ g/g.
  • the organic sulfide in the gasoline raw material is not particularly limited, and may be, for example, a mercaptan, a thioether, a thiophene, an alkylthiophene, a benzothiophene, and/or a methylbenzothiophene.
  • the cutting point of the light gasoline fraction and the heavy gasoline fraction in the present invention may be changed as needed.
  • the light gasoline fraction and the heavy gasoline fraction have a cut point of from 60 to 80 ° C, more preferably from about 65 to 70 ° C.
  • the gasoline full fraction is cut in a fractionation column from low to high in a distillation range, more preferably, the gasoline-cut fractionation column is operated at a column top temperature of about 60 At -80 ° C, the bottom temperature is about 120-160 ° C and the operating pressure is about 0.05-0.3 MPa.
  • the "etherification treatment” means that an etherification reaction of a low-carbon hydrocarbon (e.g., isopentene and cyclopentene) below C5 in a light gasoline fraction with an alcohol is carried out to produce a high octane number.
  • a low-carbon hydrocarbon e.g., isopentene and cyclopentene
  • Etherified oil e.g., isopentene and cyclopentene
  • the etherification treatment comprises the steps of: contacting the light gasoline fraction with an alcohol to cause an olefin in the light gasoline fraction to ether with an alcohol under the action of an etherification catalyst a reaction to obtain the etherified oil; wherein the etherification reaction temperature is about 20-200 ° C, the pressure is about 0.1-5 MPa, and the weight hourly space velocity is about 0.1-20 hours -1 , the alcohol and The light gasoline fraction has a molar ratio of about 1: (0.1-100).
  • the etherification catalyst comprises at least one selected from the group consisting of a resin, a molecular sieve, and a heteropolyacid.
  • the alcohol is at least one selected from the group consisting of methanol, ethanol, and propanol.
  • the strongly acidic cation exchange resin catalyst is charged in a length of etherification and/or two-stage etherification fixed bed reactor, and subjected to pretreatment such as desulfurization and dediolefinization.
  • the light gasoline fraction is passed to the etherified fixed bed reactor at a reaction temperature of about 50-90 ° C, a liquid hourly space velocity of about 1.0-3.0 h -1 , and a molar ratio of active olefin to methanol in the light gasoline fraction of about 1
  • the etherification reaction is carried out under the conditions of -2, and the etherified product is sent to a rectification column for separation, an etherified oil is obtained at the bottom of the column, and unreacted light hydrocarbons and methanol are recycled.
  • the "active olefin" means an olefin having a double bond at a tertiary carbon atom.
  • the etherification process has a reaction temperature of about 55-60 ° C at an inlet temperature, an outlet temperature of less than about 90 ° C, a space velocity of about 1-2 h -1 , and an active olefin in methanol and light gasoline fractions.
  • the molar ratio is about 1.2-1.4.
  • the one-stage etherified olefin content is relatively high, and a mixed-phase bed reactor is suitable, and the second-stage etherified olefin content is low, and an adiabatic fixed-bed reactor is suitable.
  • the isomerization unit can also be applied to the light gasoline etherification process.
  • Light gasoline etherification has many advantages such as lowering olefin content of gasoline, increasing octane number, lowering vapor pressure, increasing added value and enhancing blending efficiency.
  • Etherified oil can be used as gasoline octane blending component or heavy gasoline. The fractions are mixed as a full-distillate gasoline product.
  • the light gasoline fraction is pretreated prior to the etherification reaction to remove impurities such as sulfides and/or diolefins, thereby extending the life of the etherification catalyst.
  • the pretreatment is at least one selected from the group consisting of an alkali extraction treatment, a mercaptan conversion treatment, and a selective hydrotreatment.
  • the lye extraction process removes the thiol of the light gasoline fraction from the lye using an alkali solution; the thiol conversion treatment converts the small molecule thiol It can be removed for other sulfides, which can be carried out by a conventional alkali-free deodorization process and pre-hydrogenation, and the catalyst and cocatalyst used can be any catalysts commonly used in the art.
  • the selective hydrotreating is well known to those skilled in the art for removing diolefins from gasoline and isomerizing 3-methyl-1-butene to 2-methyl-1-butene .
  • the above pretreatment process is not essential and may be omitted in some embodiments.
  • the present invention provides a gasoline treatment system comprising a desulfurization-aromatization reactor, a high pressure separator, a gasoline full fraction fractionation column, and an etherification unit, wherein the desulfurization-aromatization reaction
  • the apparatus is provided with a gasoline raw material inlet, a hydrogen inlet, and a desulfurization-aromatization product outlet
  • the gasoline full fraction fractionation tower is provided with a gasoline full fraction inlet, and is light a gasoline distillate outlet and a heavy gasoline distillate outlet
  • the desulfurization-aromatization product outlet of the desulfurization-aromatization reactor is in fluid communication with the gasoline full fraction inlet of the gasoline full fraction fractionation column via the high pressure separator
  • the light gasoline fraction outlet of the gasoline full fraction fractionation column is in fluid communication with the inlet of the etherification unit
  • the gasoline feedstock enters the desulfurization-aromatization reactor through the inlet of the gasoline feedstock, wherein desulfurization and aromatization reaction are carried out in the presence of hydrogen to obtain a desulfurization-aromatization product;
  • the desulfurization-aromatization product Dividing a hydrogen-containing tail gas and a gasoline full fraction in the high-pressure separator, the whole gasoline fraction enters the gasoline full-fraction fractionation column through a gasoline full-fraction inlet, and fractionally distilling to obtain a light gasoline fraction and a heavy gasoline fraction;
  • the gasoline fraction enters the etherification unit where it undergoes an etherification reaction to give an etherified product.
  • the system further comprises an etherification product fractionation column and a mixer, wherein the etherification product fractionation column is provided with an etherification product inlet, an alcohol-containing tail gas outlet, and an etherified oil outlet; An outlet of the etherification unit is in fluid communication with an etherification product inlet of the etherification product fractionation column, the etherified oil outlet of the etherification product fractionation column is in fluid communication with an inlet of the mixer; and the gasoline full fraction fractionation The heavy gasoline fraction outlet of the column is in fluid communication with the inlet of the mixer,
  • the etherification product is fractionated in the etherification product fractionation column to obtain an alcohol-containing tail gas and an etherified oil, and the etherified oil and the heavy gasoline fraction are mixed in the mixer to obtain a gasoline product.
  • gasoline feedstock inlet and the hydrogen inlet of the desulfurization-aromatization reactor are the same inlet.
  • the gasoline processing system of the present invention comprises: a desulfurization-aromatization reactor 3, a high pressure separator 5, a gasoline full fraction fractionation column 8, a pretreatment unit 11, and an etherification unit. 14.
  • the desulfurization and aromatization product outlet of the vessel 3 is in fluid communication with the gasoline full fraction inlet of the gasoline full fraction fractionation column 8 via the high pressure separator 5; the light gasoline fraction outlet of the gasoline full fractionation column 8 is passed through Pretreatment unit 11 and etherification unit 14 and said etherification product fractionation column
  • the etherification product inlet of 16 is in fluid communication, the etherified oil outlet of the etherification product fractionation column 16 is in
  • the process of the present invention comprises mixing a high olefin high sulfur gasoline feedstock 1 with hydrogen 2 and feeding it to a desulfurization-aromatization reactor 3 where it is combined with an adsorption desulfurization catalyst and aromatization.
  • the catalyst is contacted for adsorption desulfurization and aromatization reaction, and after the reaction, the desulfurization-aromatization product 4 enters the high pressure separator 5 to obtain a low sulfur gasoline full fraction 7 and a hydrogen containing tail gas 6.
  • the gasoline full fraction 7 enters the gasoline full fractionation column 8 for fractionation and fractionation to obtain a light gasoline fraction 9 and a low sulfur heavy gasoline fraction 10.
  • the light gasoline fraction 9 is pretreated by desulfurization or the like through the pretreatment unit 11 to obtain the pre-etherification light gasoline 12, and then mixed with the alcohol-containing stream 13 to enter the etherification unit 14 for reaction, and the etherification product 15 is fractionated by the etherification product fractionation column 16.
  • An etherified oil 18 and an alcohol-containing tail gas 17 are obtained.
  • the heavy gasoline fraction 10 is combined with the etherified etherified oil 18 in a mixer 19 to obtain a high octane clean gasoline product 20.
  • the degree of crystallinity is measured by a standard method of ASTM D5758-2001 (2011) e1.
  • n(SiO 2 )/n(Al 2 O 3 ), that is, the ratio of silicon to aluminum is calculated by the content of silicon oxide and aluminum oxide, and the content of silicon oxide and aluminum oxide is determined by the standard method of GB/T 30905-2014. Determination.
  • the phosphorus content is determined by the standard method of GB/T 30905-2014
  • the content of the supported metal is determined by the standard method of GB/T 30905-2014
  • the sodium content is determined by the standard method of GB/T 30905-2014.
  • the total specific surface area (S BET ), the mesopore volume, the total pore volume, and the mesoporous volume of 2-20 nm are determined as follows:
  • Instrument parameters The sample was placed in a sample processing system, vacuumed to 1.33 ⁇ 10 -2 Pa at 300 ° C, and kept for 4 h at room temperature to purify the sample. At the liquid nitrogen temperature of -196 °C, the adsorption amount and desorption amount of nitrogen in the purified sample under different specific pressure P/P 0 conditions were tested, and the N 2 adsorption-desorption isotherm curve was obtained. Then use the two-parameter BET formula to calculate the total specific surface area, the specific surface area of the micropore and the specific surface area of the mesopores.
  • the method for measuring the amount of B acid and the amount of L acid is as follows:
  • the FTS3000 type Fourier infrared spectrometer produced by American BIO-RAD company was used.
  • Test conditions The sample was pressed into a sheet and placed in an in-situ bath of an infrared spectrometer to be sealed, and vacuumed to 10 -3 Pa at 350 ° C for 1 h to desorb the gas molecules on the surface of the sample and cool to room temperature.
  • the pyridine vapor with a pressure of 2.67 Pa was introduced into the in-situ tank, and after equilibration for 30 min, the temperature was raised to 200 ° C, and the vacuum was again evacuated to 10 -3 Pa for 30 min, cooled to room temperature, and scanned in the wave number range of 1400-1700 cm -1 .
  • the infrared spectrum of the adsorption of pyridine at 200 ° C was recorded.
  • the sample in the infrared absorption cell was moved to the heat treatment zone, heated to 350 ° C, evacuated to 10 -3 Pa, held for 30 min, cooled to room temperature, and the infrared spectrum of the pyridine adsorption at 350 ° C was recorded.
  • the instrument automatically integrates the amount of B acid and L acid.
  • the method for measuring the total acid amount and the strong acid amount is as follows:
  • Test conditions Weigh 0.2g sample to be loaded into the sample tube, placed in the heating cell heating furnace, He gas is carrier gas (50mL / min), temperature is raised to 600 ° C at 20 ° C / min, purged for 60 min to drive out Impurities adsorbed on the surface of the catalyst. Then, the temperature was lowered to 100 ° C, the temperature was kept for 30 min, and the mixture was switched to NH 3 -He mixture gas (10.02% NH 3 + 89.98% He) for 30 min, and then purged with He gas for 90 min until the baseline was stable to desorb the physically adsorbed ammonia gas. .
  • the temperature was raised to 600 ° C at a heating rate of 10 ° C / min for desorption, and was maintained for 30 min, and the desorption was completed.
  • the TCD detector is used to detect the change of gas composition, and the instrument automatically integrates to obtain the total acid amount and the strong acid amount.
  • the acid center of the strong acid is the acid center corresponding to the NH 3 desorption temperature greater than 300 °C.
  • the D value is calculated as follows:
  • test method for the dry basis weight is:
  • the molecular sieve or catalyst to be tested was placed in a muffle furnace and calcined in an air atmosphere at 600 ° C for 3 hours, and the obtained calcined product was cooled to room temperature in a closed drying dish, and then weighed.
  • the adsorptive desulfurization catalysts used in the following examples and comparative examples were produced by Catalyst Branch of China Petroleum & Chemical Corporation under the trade name FCAS.
  • the aromatization catalysts used included laboratory-made catalysts, numbered OTAZ-C- 3, and the commercial aromatization catalyst, the commercial brand number is MP051, produced by the Catalyst Branch of China Petroleum & Chemical Corporation.
  • the properties of the respective catalysts used in the examples are shown in Table 1.
  • catalyst FCAS OTAZ-C-3 MP051 Chemical composition, weight% Alumina 13 50.3 50.8 Sodium oxide / / / Nickel oxide twenty one / / Zinc oxide 52 / / Gallium oxide / 1.5 / Silicon oxide 14 48.14 48.7 Apparent density, kg/m 3 1010 800 790 Pore volume, mL/g / 0.27 0.23 Specific surface area, m 2 /g / 218 215 Wear index, weight % -1 / 1.5 1.4 Screening composition, weight% 0 to 40 microns 14.5 13.9 15 40 to 80 microns 51.9 49.5 50.5 >80 microns 33.6 36.6 34.5 Microreaction activity / 80 80 80
  • the properties of the raw materials used are as follows: kaolin (Suzhou China Kaolin Company, solid content: 75% by weight), pseudo-boehmite (Shandong Aluminum Co., Ltd., solid content of 65% by weight, firstly used in the concentration of 31% by weight hydrochloric acid peptization, The molar ratio of hydrochloric acid to pseudoboehmite in terms of alumina was 0.20).
  • the weight of the filter cake was filtered; 100 g of the above molecular sieve (dry basis) was added to 1000 g of a 2.0% NaOH solution, and the temperature was raised to 65 ° C. After reacting for 30 minutes, the mixture was rapidly cooled to room temperature, filtered, and washed until the filtrate was neutral.
  • the filter cake was added to 800 g of water and beaten, 40 g of NH 4 Cl was added, and the temperature was raised to 75 ° C. After exchange treatment for 1 h, the Na 2 O content was less than 0.2% by weight, filtered, and washed to obtain a molecular sieve cake; the above molecular sieve 50 g was taken. (dry basis) adding water to prepare a molecular sieve slurry with a solid content of 10% by weight, adding 11 g of oxalic acid with stirring, and then adding 110 g of hydrochloric acid (10% by mass) and 92 g of fluorosilicic acid (3% by mass) in parallel for 30 min.
  • the pseudo-boehmite is mixed with kaolin and formulated into a slurry having a solid content of 30% by weight with deionized water, stirred uniformly, and the pH of the slurry is adjusted to 2.5 with hydrochloric acid, the pH is maintained, and the mixture is allowed to stand at 50 ° C. After aging for 1 hour, the mixture was stirred for 1 hour to form a colloid, and the prepared molecular sieve A and water were added to form a catalyst slurry (solid content of 35% by weight). The stirring was continued, followed by spray drying to prepare a microsphere catalyst.
  • microsphere catalyst was then calcined at 500 ° C for 1 hour to obtain the OTAZ-C-3 aromatization catalyst used in the present invention, which had a dry basis ratio of 25% by weight of kaolin, 25% by weight of pseudoboehmite and 50% by weight.
  • Molecular sieve A was then calcined at 500 ° C for 1 hour to obtain the OTAZ-C-3 aromatization catalyst used in the present invention, which had a dry basis ratio of 25% by weight of kaolin, 25% by weight of pseudoboehmite and 50% by weight.
  • the passivation treatment of the aromatization catalyst is as follows:
  • the fresh aromatization catalyst numbered OTAZ-C-3 was charged into a dense-phase fluidized bed reactor, and contacted with stabilized gasoline A at a passivation temperature of 410 ° C for 2 hours to obtain a blunt
  • the OTAZ-C-3 catalyst had a passivate content of 0.5% by weight and a microreactivity of 35.
  • the fresh aromatization catalyst of MP051 was charged into a dense-phase fluidized bed reactor, and contacted with stabilized gasoline A at a passivation temperature of 410 ° C for 2 hours to obtain a passivated MP051 catalyst.
  • the passivation content was 0.5% by weight and the microreaction activity was 35.
  • the aging treatment of the aromatization catalyst is as follows:
  • the fresh aromatization catalyst of the grade OTAZ-C-3 was charged into a dense-phase fluidized bed reactor, and continuously aged for 8 hours under the aging temperature of 780 ° C and 100% steam to obtain an aged OTAZ-C. -3 catalyst with a microreactivity of 40.
  • the fresh aromatization catalyst of MP051 was charged into the dense-phase fluidized bed reactor and continuously aged for 8 hours under the aging temperature of 780 ° C and 100% steam to obtain the aged MP051 catalyst.
  • the micro-reaction activity was 40.
  • the contents of Na 2 O, NiO, ZnO, Ga 2 O 3 , Al 2 O 3 and SiO 2 in the catalyst are determined by X-ray fluorescence method, and the specific test method is as follows: GB/T 30905 -2014.
  • the content of the passivation in the passivated aromatization catalyst is determined by the RIPP 107-90 method, and the specific steps are summarized as follows: the sample is burned in a high frequency furnace to generate CO 2 and CO gas. The gas enters the infrared detector through a filter drier to give a percentage of total char. Since the catalyst itself has no induction effect, it is necessary to add a sensing agent, a fluxing agent to promote the combustion of the sample, and because the sample is porous, it has water absorption property, and needs to be pretreated before the test. Specific test and test procedures include:
  • Test preparation Preheat, clean, test leak and check the circuit according to the operating procedures of CS-46 carbon sulfur tester (produced by American LECO company). The catalyst sample was baked at 110 ° C for 2-3 hours and cooled to room temperature in a desiccator.
  • the micro-reaction activity test method of the aromatization catalyst adopts the RIPP 92-90 method, wherein the same equipment and test procedure as ASTM D3907-2013 is used, under the following raw material oil and test reaction conditions Conducted, feedstock oil: Dagang straight-run diesel, relative density d 4 20 is 0.8419, initial boiling point 235 ° C, dry point 337 ° C; reaction conditions: oil input 1.56 g, reaction temperature 460 ° C, the ratio of the mass ratio of the agent oil 3.2, raw materials The oil has a weight hourly space velocity of 16h -1 and a feed time of 70 seconds.
  • Gasoline PONA uses simulated distillation and gasoline monomer hydrocarbon analysis (tested by ASTM D2887 and ASTM D6733-01 (2011) test methods, respectively), and gasoline sulfur content is determined using SH/T0689-2000.
  • the gasoline feedstock used in the following examples and comparative examples was Stabilized Gasoline A, the properties of which are listed in Table 3 below:
  • Examples I-1 to I-10 are examples in which the above-mentioned gasoline raw materials were subjected to desulfurization and aromatization reaction using various types of catalysts, and the results of the reactions are shown in Table 4.
  • Gasoline feedstock in a small continuous fluidized bed reactor and including an adsorption desulfurization catalyst The FCAS is contacted with a mixed catalyst of the passivated OTAZ-C-3 catalyst to carry out adsorption desulfurization and aromatization reactions.
  • the OTAZ-C-3 catalyst accounts for 7% by weight of the total weight of the mixed catalyst.
  • the operating conditions were: a reaction temperature of 400 ° C, a reactor pressure of 2.1 MPa, a weight hourly space velocity of the gasoline feedstock of 6 hours -1 , and a volume ratio of hydrogen to gasoline feedstock of 75.
  • the desulfurization-aromatization product obtained from the top of the reactor was cooled to about 10 ° C, and the gas product and the whole gasoline fraction were separated, and the whole gasoline fraction was not cut and etherified as a gasoline product, and its properties are shown in Table 4.
  • the regeneration temperature of the mixed catalyst was 550 ° C, and the mixed catalyst after the regeneration was returned to the reactor for recycling.
  • Example I-1 The operation was substantially the same as in Example I-1, except that the passivated OTAZ-C-3 catalyst was replaced with an equivalent weight of the aged OTAZ-C-3 catalyst, and the whole fraction of the gasoline was not cut and Etherification is a gasoline product whose properties are listed in Table 4.
  • Example I-1 The operation was substantially the same as in Example I-1, except that the passivated OTAZ-C-3 catalyst was replaced with an equivalent weight of fresh OTAZ-C-3 catalyst, and the whole fraction of gasoline was not cleaved and ether.
  • OTAZ-C-3 catalyst As a gasoline product, its properties are listed in Table 4.
  • Example I-1 The operation was substantially the same as in Example I-1, except that the passivated OTAZ-C-3 catalyst used an equal weight of mixed OTAZ-C-3 catalyst (including 5% by weight of passivation catalyst and 2 The weight percent of fresh catalyst was replaced by the total weight of the mixed adsorptive desulfurization catalyst and the aromatization catalyst.
  • the whole gasoline fraction was cut and etherified as a gasoline product, the properties of which are listed in Table 4.
  • Example I-1 The operation was substantially the same as in Example I-1, except that the passivated OTAZ-C-3 catalyst used an equal weight of mixed OTAZ-C-3 catalyst (including 5% by weight of aged catalyst and 2 weights). The % fresh catalyst was replaced by the total weight of the mixed adsorptive desulfurization catalyst and the aromatization catalyst. The whole gasoline fraction was not cut and etherified as a gasoline product, and its properties are shown in Table 4.
  • Example I-1 The operation was substantially the same as in Example I-1, except that the passivated OTAZ-C-3 catalyst was replaced with an equivalent weight of passivated MP051 catalyst, gasoline The whole fraction was cut and etherified as a gasoline product, the properties of which are listed in Table 4.
  • Example I-1 The operation was substantially the same as in Example I-1, except that the passivated OTAZ-C-3 catalyst was replaced with an equivalent weight of aged MP051 catalyst, and the whole fraction of gasoline was not cut and etherified as gasoline.
  • the properties of the products are listed in Table 4.
  • Example I-1 The operation was substantially the same as in Example I-1, except that the passivated OTAZ-C-3 catalyst was replaced with an equivalent weight of fresh MP051 catalyst, and the whole fraction of gasoline was not cut and etherified as a gasoline product. Its properties are listed in Table 4.
  • Example I-1 The operation was substantially the same as in Example I-1, except that the passivated OTAZ-C-3 catalyst used an equal weight of mixed MP051 catalyst (including 5% by weight of passivating catalyst and 2% by weight of fresh). The catalyst was replaced by the total weight of the mixed adsorptive desulfurization catalyst and the aromatization catalyst. The whole fraction of the gasoline was not cut and etherified as a gasoline product, and its properties are shown in Table 4.
  • Example I-1 The operation was substantially the same as in Example I-1, except that the passivated OTAZ-C-3 catalyst used an equal weight of mixed MP051 catalyst (including 5% by weight of aged catalyst and 2% by weight of fresh catalyst). Replaced by the total weight of the mixed adsorptive desulfurization catalyst and the aromatization catalyst, the whole fraction of gasoline is not cut and etherified as a gasoline product, and its properties are listed in Table 4.
  • Example I-1 The operation was substantially the same as that of Example I-1, except that the passivated OTAZ-C-3 catalyst was replaced with an equivalent weight of the adsorption desulfurization catalyst FCAS, and the whole fraction of the gasoline was not cut and etherified as a gasoline product. Its properties are listed in Table 4.
  • the octane number of the resulting gasoline product of Examples I-6 to I-10 compared to Comparative Example I-1 was significantly improved, particularly when using an aromatization catalyst comprising a passivated or aged MP051 catalyst. More apparent; further, when the heavy gasoline fraction is treated with a mixed catalyst comprising an adsorbed desulfurization catalyst FCAS and an aromatization catalyst OTAZ-C-3, Examples I-1 to I-5 and Examples I-6 The octane number of the resulting gasoline product was further improved compared to I-10, and the improvement effect of the aromatization catalyst comprising the passivated or aged OTAZ-C-3 catalyst was more pronounced.
  • Examples II-1 to II-3 are examples in which the gasoline raw materials are subjected to desulfurization and aromatization reaction using various types of catalysts, followed by dividing the whole gasoline fraction and etherifying the light gasoline fraction.
  • the reaction results are shown in Table 6.
  • the gasoline feedstock is contacted with a mixed catalyst comprising an adsorbed desulfurization catalyst FCAS and a passivated OTAZ-C-3 catalyst in a small continuous fluidized bed reactor for adsorption desulfurization and aromatization reactions.
  • a mixed catalyst comprising an adsorbed desulfurization catalyst FCAS and a passivated OTAZ-C-3 catalyst in a small continuous fluidized bed reactor for adsorption desulfurization and aromatization reactions.
  • the OTAZ-C-3 catalyst accounts for 7% by weight of the total weight of the mixed catalyst.
  • the operating conditions were: a reaction temperature of 400 ° C, a reactor pressure of 2.1 MPa, a weight hourly space velocity of the gasoline feedstock of 6 hours -1 , and a volume ratio of hydrogen to gasoline feedstock of 75.
  • the desulfurization-aromatization product obtained from the top of the reactor was cooled to about 10 ° C and separated to obtain a tail gas and a gasoline full fraction.
  • the regeneration temperature of the mixed catalyst was 550 ° C, and the mixed catalyst after the regeneration was returned to the reactor for recycling.
  • the whole gasoline fraction is distilled in a fractionation column, cut into light gasoline fraction and low sulfur heavy gasoline fraction, and the light fraction end point is controlled to be 65-70 ° C (according to ASTM D86 standard).
  • the light gasoline fraction obtained by distillation is referred to as LCN-A-1
  • the heavy gasoline fraction is recorded as HCN-A-1.
  • Table 5 The properties of the light gasoline fraction and the heavy gasoline fraction obtained by distillation are shown in Table 5.
  • the light gasoline fraction is subjected to desulfurization and dediolefin pretreatment in the presence of hydrogen through a refining reactor, so that the sulfur content and the diolefin content in the light gasoline fraction are all reduced to below 10 ppm, and then mixed with methanol (analytically pure) into the etherification reactor.
  • the etherification reaction is carried out under the conditions of a reaction temperature of 55-80 ° C, a space velocity of 1.2 h -1 , a molar ratio of methanol to a reactive olefin in the light gasoline fraction of 1.2, and an etherification catalyst of a strong acid cation exchange resin (type 732 benzene).
  • the etherified oil and gas obtained by the reaction is separated into the etherification fractionation column.
  • the gas phase at the top of the column is methanol-containing tail gas including residual C5 and methanol, and the etherified oil is obtained at the bottom of the column.
  • the temperature at the top of the etherification fractionation column is 60- At 80 ° C, the bottom temperature is 110-140 ° C.
  • the corresponding etherified oil is referred to as LCN-AM-1 and its properties are also listed in Table 5.
  • the obtained adsorbed and desulfurized heavy gasoline fraction HCN-A-1 was mixed with the etherified oil LCN-A-M-1 to obtain a high-octane clean gasoline product, and the properties are listed in Table 6.
  • the gasoline feedstock is contacted with a mixed catalyst comprising an adsorbed desulfurization catalyst FCAS and an aged OTAZ-C-3 catalyst in a small continuous fluidized bed reactor for adsorption desulfurization and aromatization reactions.
  • a mixed catalyst comprising an adsorbed desulfurization catalyst FCAS and an aged OTAZ-C-3 catalyst in a small continuous fluidized bed reactor for adsorption desulfurization and aromatization reactions.
  • the OTAZ-C-3 catalyst accounts for 7% by weight of the total weight of the mixed catalyst.
  • the operating conditions were: a reaction temperature of 400 ° C, a reactor pressure of 2.1 MPa, a weight hourly space velocity of the gasoline feedstock of 6 hours -1 , and a volume ratio of hydrogen to gasoline feedstock of 75.
  • the desulfurization-aromatization product obtained from the top of the reactor was cooled to about 10 ° C and separated to obtain a tail gas and a gasoline full fraction.
  • the regeneration temperature of the mixed catalyst was 550 ° C, and the mixed catalyst after the regeneration was returned to the reactor for recycling.
  • the whole gasoline fraction was treated in the same manner as in Example II-1, and the obtained light gasoline fraction was recorded as LCN-A-2, the heavy gasoline fraction was recorded as HCN-A-2, and the etherified oil was recorded as LCN- AM-2, properties are listed in Table 5; the obtained etherified oil was mixed with the heavy gasoline fraction to obtain a high octane clean gasoline product, the properties of which are listed in Table 6.
  • the gasoline feedstock is contacted with a mixed catalyst comprising an adsorbed desulfurization catalyst FCAS and a fresh OTAZ-C-3 catalyst in a small continuous fluidized bed reactor for adsorption desulfurization and aromatization reactions.
  • a mixed catalyst comprising an adsorbed desulfurization catalyst FCAS and a fresh OTAZ-C-3 catalyst in a small continuous fluidized bed reactor for adsorption desulfurization and aromatization reactions.
  • the OTAZ-C-3 catalyst accounts for 7% by weight of the total weight of the mixed catalyst.
  • the operating conditions were: a reaction temperature of 400 ° C, a reactor pressure of 2.1 MPa, a weight hourly space velocity of the gasoline feedstock of 6 hours -1 , and a volume ratio of hydrogen to gasoline feedstock of 75.
  • the desulfurization-aromatization product obtained from the top of the reactor was cooled to about 10 ° C and separated to obtain a tail gas and a gasoline full fraction.
  • the regeneration temperature of the mixed catalyst was 550 ° C, and the mixed catalyst after the regeneration was returned to the reactor for recycling.
  • the whole gasoline fraction was treated in the same manner as in Example II-1, and the obtained light gasoline fraction was recorded as LCN-A-3, the heavy gasoline fraction was recorded as HCN-A-3, and the etherified oil was recorded as LCN- AM-3, properties are listed in Table 5; the resulting etherified oil was mixed with the heavy gasoline fraction to give a high octane clean gasoline product, the properties of which are listed in Table 6.
  • Example II-1 The procedure was essentially the same as in Example II-1, except that the passivated OTAZ-C-3 catalyst was replaced with an equivalent weight of the adsorbed desulfurization catalyst FCAS.
  • the properties of the resulting gasoline products are listed in Table 6.
  • Comparative Example II-1 of the adsorbed desulfurization catalyst FCAS was used.
  • the obtained gasoline product was significantly lower in octane number (antiknock index) than the raw gasoline;
  • Examples II-1 to II-3 using a mixed catalyst comprising an adsorbed desulfurization catalyst FCAS and an aromatization catalyst OTAZ-C-3 Comparative Example II-1 showed a significant improvement in the octane number of the resulting gasoline product, especially when using an aromatization catalyst comprising a passivated or aged OTAZ-C-3 catalyst.

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Abstract

一种汽油的处理方法及处理系统，该处理方法包括：使汽油原料与混合催化剂接触并在氢气存在下进行脱硫和芳构化反应，得到脱硫-芳构化产物；任选地将所得脱硫-芳构化产物进行分割，得到轻汽油馏分和重汽油馏分；以及任选地将所得轻汽油馏分进行醚化处理，得到醚化油；其中，所述混合催化剂包括吸附脱硫催化剂和芳构化催化剂。该方法能降低汽油中的硫和烯烃含量，并能够同时提高汽油的辛烷值和保持高汽油收率。

Description

一种汽油的处理方法 技术领域
本发明涉及一种汽油的处理方法，更具体地涉及一种对汽油进行脱硫处理的方法。
背景技术
汽车尾气排放导致的大气污染日益严重。随着人们对环境保护的日益重视，世界各国都加快了车用燃料质量升级的步伐，例如中国国家标准GB17930-2016要求汽油中硫含量不大于10μg/g，烯烃体积分数不大于24％。
催化裂化汽油是车用汽油的主要组成部分，在汽油池中占75％左右，而其特点是具有较高的烯烃和硫含量。采用加氢技术实现汽油的深度脱硫和降低催化裂化汽油中的烯烃含量并不困难，但由于烯烃是高辛烷值组分，其含量的大幅降低将导致汽油辛烷值的严重损失，从而影响汽油的车用性能和炼油厂的经济效益，因此在实现汽油的深度脱硫同时保持汽油辛烷值成为清洁汽油生产的热点。
目前，汽油的深度脱硫主要是采用加氢脱硫或吸附脱硫的方法。
选择性加氢脱硫是目前脱除噻吩类硫化物的主要方式之一，但烯烃饱和等反应同样大量发生，造成辛烷值损失较大。此外，恢复辛烷值的深度加氢方法也同样被人们认可，是在进行深度脱硫和烯烃饱和的同时，设置第二反应器以促进低辛烷值的烃类裂化、异构化和烷基化反应，从而达到恢复辛烷值的目的。
中国专利申请CN101845322A公开了一种降低汽油中硫和烯烃含量的方法，原料催化裂化汽油先经过预加氢反应器脱除二烯烃，然后进入分馏塔切割分馏成轻、重汽油，轻汽油在氢气存在下进行吸附脱硫，重汽油进入选择性加氢反应器加氢脱硫，反应流出物再进入加氢改质反应器进行加氢改质，降低烯烃含量，改质后的重汽油与轻汽油吸附脱硫产品调和得到满足标准要求的清洁汽油。吸附脱硫催化剂对于汽油中的硫化物虽然具有很好的脱除效果，但吸附脱硫在氢气存在下进行，会饱和催化裂化汽油中的烯烃，尤其是轻汽油进行吸附脱硫，轻汽油中的烯烃组分辛烷值较高，会造成汽油的辛烷值大量损失。
吸附法脱除燃油中的含硫化合物，是在氢气存在下采用吸附剂对轻质油进行反应吸附，生成金属硫化物或利用硫化物极性脱除硫，氢耗较低，脱硫效率高，可以生产硫含量在10μg/g以下汽油。虽然吸附工艺在低氢耗条件下实现了汽油的深度脱硫，汽油产物辛烷值仍然略有损失。尤其是当处理烯烃含量高和硫含量高的汽油原料时，仍然造成汽油辛烷值损失较大。
对于大多数催化裂化装置而言，为了增产丙烯和丁烯产量并提高汽油的辛烷值，采用含有具有MFI结构分子筛的催化剂或助剂是有效的方法。美国专利USP3758403报导了在催化裂化催化剂中添加ZSM-5分子筛的方法可以提高汽油的辛烷值和增加C3～C4烯烃的产率。但是，正如本领域技术人员所知，增产丙烯和丁烯产量是以牺牲汽油产量为代价的。
因此，本领域仍然需要一种能够有效地降低汽油中的硫和烯烃含量，同时能够保持甚至提高汽油的收率以及辛烷值的汽油处理方法。
发明内容
本发明的目的是提供一种能够降低汽油中的硫和烯烃含量，并能够同时提高汽油的辛烷值和保持高汽油收率的汽油处理方法和系统。
为了实现上述目的，一方面，本发明提供了一种汽油的处理方法，包括：
使汽油原料与混合催化剂接触并在氢气存在下进行脱硫和芳构化反应，得到脱硫-芳构化产物，
其中，所述混合催化剂包括吸附脱硫催化剂和芳构化催化剂，并且至少约50wt％、优选至少约80wt％、更优选至少约90wt％、特别优选至少约95wt％、最优选约100wt％的所述芳构化催化剂经过钝化和/或老化处理。
另一方面，本发明提供了一种汽油的处理方法，包括：
1)使汽油原料与混合催化剂接触并在氢气存在下进行脱硫和芳构化反应，得到脱硫-芳构化产物；
2)将至少一部分所得脱硫-芳构化产物进行切割，得到轻汽油馏分和重汽油馏分；
3)将至少一部分所得轻汽油馏分进行醚化处理，得到醚化油；以 及
4)任选地，将至少一部分所述醚化油与至少一部分所述重汽油馏分混合，得到汽油产品，
其中，所述混合催化剂包括吸附脱硫催化剂和芳构化催化剂，且所述芳构化催化剂选自新鲜芳构化催化剂、经钝化处理的芳构化催化剂、经老化处理的芳构化催化剂，或者它们的任意组合。
再一方面，本发明提供了一种汽油的处理系统，该系统包括脱硫-芳构化反应器、高压分离器、汽油全馏分分馏塔和醚化装置，其中所述脱硫-芳构化反应器设置有汽油原料入口、氢气入口、以及脱硫-芳构化产物出口，所述汽油全馏分分馏塔设置有汽油全馏分入口、轻汽油馏分出口和重汽油馏分出口；所述脱硫-芳构化反应器的脱硫-芳构化产物出口经由所述高压分离器与所述汽油全馏分分馏塔的汽油全馏分入口流体连通；所述汽油全馏分分馏塔的轻汽油馏分出口与醚化装置的入口流体连通，
其中汽油原料经所述汽油原料入口进入所述脱硫-芳构化反应器，在其中在氢气存在下进行脱硫和芳构化反应，得到脱硫-芳构化产物；所述脱硫-芳构化产物在所述高压分离器分离得到含氢尾气和汽油全馏分，所述汽油全馏分经汽油全馏分入口进入所述汽油全馏分分馏塔，在其中分馏得到轻汽油馏分和重汽油馏分；所述轻汽油馏分进入所述醚化装置，在其中进行醚化反应，得到醚化产物。
本发明与现有技术相比具有以下技术效果中的至少一种：
1、本发明的方法使高硫高烯烃的汽油原料与吸附脱硫催化剂和芳构化催化剂进行脱硫和芳构化反应，在降低汽油硫含量的同时，能够将汽油中的烯烃进行裂解和芳构化，生产C5以下烃类并优选切割在轻汽油馏分中进行醚化反应，从而既降低了汽油中烯烃的含量，还提高了汽油的辛烷值且保持汽油原料的高收率。
2、本发明的脱硫和芳构化反应可以采用现有吸附脱硫反应器进行，而无需对现有反应器进行改造。
3、本发明的脱硫和芳构化反应采用两种催化剂在一个反应器中进行，可以避免汽油分步处理方法(汽油先脱硫后芳构化，或汽油先芳构化后脱硫)需要单独增加芳构化反应器及附属系统的问题，也可以避免汽油脱硫和芳构化耦合催化剂需要改变现有的催化剂制备工艺 且催化剂磨损强度较低的问题，既提高了反应效率，还降低了投资成本。
4、本发明的芳构化催化剂优选经过钝化和/或老化处理，使芳构化催化剂的活性适中，从而更有利于芳构化的进行。
5、本发明优选对轻汽油馏分进行醚化处理，既能够降低其中的烯烃，还能够生产高辛烷值的醚化油，提高汽油产品的辛烷值。
6、本发明优选采用的醚化处理还能够减少汽油产品中的轻组分，降低汽油蒸汽压。
本发明的其他特征和优点将在随后的具体实施方式部分予以详细说明。
附图说明
附图是用来帮助对本发明的进一步理解，并且构成说明书的一部分，与下面的具体实施方式一起用于解释本发明，但并不构成对本发明的限制。在附图中：
图1是本发明方法和系统的一种优选实施方式的示意图。
附图标记说明
1汽油原料         2氢气               3脱硫-芳构化反应器
4脱硫-芳构化产物  5高压分离器         6含氢尾气
7汽油全馏分       8汽油全馏分分馏塔   9轻汽油馏分
10重汽油馏分      11预处理单元        12醚化前轻汽油
13含醇物流        14醚化装置          15醚化产物
16醚化产物分馏塔  17含醇尾气          18醚化油
19混合器          20汽油产品
具体实施方式
以下结合附图对本发明的具体实施方式进行详细说明。应当理解的是，此处所描述的具体实施方式仅用于说明和解释本发明，并不用于限制本发明。
在本文中所披露的任何具体数值(包括数值范围的端点)都不限于该数值的精确值，而应当理解为还涵盖了接近该精确值的值。并且， 对于所披露的数值范围而言，在该范围的端点值之间、端点值与范围内的具体点值之间，以及各具体点值之间可以任意组合而得到一个或多个新的数值范围，这些新的数值范围也应被视为在本文中具体公开。
本发明中RIPP试验方法具体可参见《石油化工分析方法》，杨翠定等编，科学出版社，1990年9月，第263-268和303-304页，ISBN：7-03-001894-X。
在第一方面，本发明提供了一种汽油的处理方法，包括：
使汽油原料与混合催化剂接触并在氢气存在下进行脱硫和芳构化反应，得到脱硫-芳构化产物，
其中，所述混合催化剂包括吸附脱硫催化剂和芳构化催化剂，并且至少约50wt％、优选至少约80wt％、更优选至少约90wt％、特别优选至少约95wt％、最优选约100wt％的所述芳构化催化剂经过钝化和/或老化处理。
在某些优选的实施方式中，所述处理方法进一步包括：将至少一部分所得脱硫-芳构化产物进行蒸馏，得到C3-C4液态烃组分和稳定汽油馏分。优选地，所述液态烃组分送至气体分离单元作进一步处理。
任选地，所述稳定汽油馏分可以直接作为汽油产品，或者作为其他工艺过程的原料。
在某些优选的实施方式中，所述处理方法进一步包括如下步骤：
将至少一部分所得脱硫-芳构化产物进行切割，得到轻汽油馏分和重汽油馏分；以及
将至少一部分所得轻汽油馏分进行醚化处理，得到醚化油；
在某些进一步优选的实施方式中，所述处理方法进一步包括如下步骤：
将至少一部分所述醚化油与至少一部分所述重汽油馏分混合，得到汽油产品。
在第二方面，本发明提供了一种汽油的处理方法，包括：
1)使汽油原料与混合催化剂接触并在氢气存在下进行脱硫和芳构化反应，得到脱硫-芳构化产物；
2)将至少一部分所得脱硫-芳构化产物进行切割，得到轻汽油馏分和重汽油馏分；
3)将至少一部分所得轻汽油馏分进行醚化处理，得到醚化油；以及
4)任选地，将至少一部分所述醚化油与至少一部分所述重汽油馏分混合，得到汽油产品，
其中，所述混合催化剂包括吸附脱硫催化剂和芳构化催化剂，且所述芳构化催化剂选自新鲜芳构化催化剂、经钝化处理的芳构化催化剂、经老化处理的芳构化催化剂，或者它们的任意组合。
根据本发明，所述“脱硫和芳构化反应”是指汽油原料在氢气存在下在吸附脱硫催化剂和芳构化催化剂的共同作用下进行脱硫和将烯烃转化为芳烃的过程，期间可伴随有裂化反应。在某些优选的实施方式中，所述脱硫和芳构化反应在如下条件下进行：反应温度为约350-500℃，优选为约380-420℃，重时空速为约2-50小时-1，优选为约5-20小时-1，反应压力为约0.5-3.0MPa，优选为约1.5-2.5MPa，氢气与汽油原料体积比(在标准状况(STP)0℃(273K)、1.01×105Pa下)为约1-500，优选为约50-200。
本发明所用的吸附脱硫催化剂没有特别限制，可以是任何本领域技术人员所熟知的适用于对汽油进行吸附脱硫的催化剂。在某些优选的实施方式中，所述吸附脱硫催化剂含有二氧化硅、氧化铝、氧化锌和脱硫活性金属，所述脱硫活性金属为选自钴、镍、铜、铁、锰、钼、钨、银、锡和钒中的至少一种。
在某些优选的实施方式中，以所述吸附脱硫催化剂的干基重量为基准并以氧化物重量计，所述吸附脱硫催化剂中所述氧化锌占约10-90重量％，二氧化硅占约5-85重量％，氧化铝占约5-30重量％；以所述吸附脱硫催化剂的干基重量为基准并以元素重量计，所述吸附脱硫催化剂中所述脱硫活性金属的含量为约5-30重量％。
在一优选实施方式中，所述吸附脱硫催化剂还可以含有1-10重量％的类焦炭物质。工业实践表明，吸附脱硫催化剂上的碳含量对吸附脱硫催化剂的脱硫效率及汽油辛烷值损失均有影响，随着吸附脱硫催化剂碳含量的增多，吸附脱硫催化剂脱硫效率逐步下降，汽油辛烷值损失随之减小。同样地，吸附脱硫催化剂保持一定的硫含量是非常优选的。实践表明，待生吸附脱硫催化剂载硫量为约9-10重量％，再生吸附脱硫催化剂载硫量为约5-6重量％，且待生吸附脱硫催化剂与再生吸 附脱硫催化剂的硫差为约4重量％左右最为优选。
根据本发明，所述“芳构化催化剂”是指能够将汽油原料中的烃类如烯烃等转化为芳烃的催化剂，一般包括分子筛，优选包括分子筛、载体和金属。在某些优选的实施方式中，以干基计并以所述芳构化催化剂的总重量为基准，所述芳构化催化剂含有约10-30重量％的分子筛、约0-20重量％的芳构化活性金属氧化物和约50-90重量％的载体。
在某些进一步优选的实施方式中，所述分子筛包括Y分子筛和/或MFI结构分子筛，优选为五元环高硅分子筛，该高硅分子筛可以是氢型的，也可以是经过稀土和/或磷改性的，其硅铝比优选大于100，更优选大于150。
在某些进一步优选的实施方式中，所述芳构化活性金属可以发挥部分脱硫或烃类转化功能，例如可以为选自第IVB族的金属元素、第VB族的金属元素、第VIB族的金属元素、第VIII族的金属元素、第IB族的金属元素、第IIB族的金属元素和第IIIA族的金属元素中的至少一种。在某些更进一步优选的实施方式中，第IVB族的金属元素为Zr或/和Ti，第VB族的金属元素为V，第VIB族的金属元素为Mo或/和W，第VIII族的金属元素为选自Fe、Co、Ni中的一种或多种，第IB族的金属元素为Cu，第IIB族的金属元素为Zn，且第IIIA族的金属元素为Ga。在某些特别优选的实施方式中，所述芳构化活性金属为选自Fe、Zn和Ga中的至少一种，含量优选为0.5-5重量％。
在某些进一步优选的实施方式中，所述载体优选包括氧化硅和/或氧化铝。
在某些优选实施方式中，所述芳构化催化剂的粒径一般为20-120微米，与所述吸附脱硫催化剂的粒径相当。优选地，本发明将所述吸附脱硫催化剂与芳构化催化剂分别成型(例如喷雾干燥)后混合使用。
根据汽油原料中烯烃和硫含量的不同，本发明的混合催化剂中吸附脱硫催化剂和芳构化催化剂的比例可以不同。在某些优选的实施方式中，以重量计，所述芳构化催化剂占所述混合催化剂的比例为约1-30重量％，优选为约3-15重量％。
在某些优选的实施方式中，本发明所述的芳构化催化剂通过如下方法制备得到：将芳构化催化剂的制备原料和水进行混合打浆，然后将浆液进行喷雾干燥和焙烧；其中，以重量计并以所述制备原料的干 基重量为基准，所述制备原料包括约15-60重量％的天然矿物质、约10-30重量％的无机氧化物粘结剂的前驱体和约20-80重量％的含磷和负载金属的MFI结构分子筛。
在进一步优选的实施方式中，以重量计并以所述制备原料的干基重量为基准，所述制备原料包括：约20-55重量％的天然矿物质、约12-28重量％的无机氧化物粘结剂的前驱体和约35-70重量％的含磷和负载金属的MFI结构分子筛。
根据本发明，所述“天然矿物质”是指在地壳各种物质的综合作用下(称地质作用)形成的天然单质或化合物，并具有化学式表达的特有的化学成分和相对固定的化学成分，例如可以包括选自高岭土、多水高岭土、蒙脱土、硅藻土、凸凹棒石、海泡石、埃洛石、水滑石、膨润土和累托土中的至少一种。
根据本发明，所述“无机氧化物粘结剂”是指在催化剂中起到粘结各组分作用的无机氧化物，例如可以包括选自氧化硅、氧化铝、氧化锆、氧化钛和无定形硅铝中的至少一种。
根据本发明，所述“无机氧化物粘结剂的前驱体”是指通常用于在催化裂解催化剂中生成无机氧化物粘结剂的催化裂解催化剂制备原料，例如可以包括选自硅溶胶、铝溶胶、胶溶拟薄水铝石、硅铝溶胶和含磷铝溶胶中的至少一种。
根据本发明，所述“负载金属”是指通过负载方式负载到分子筛上的金属，不包括铝以及钠、钾等碱金属。本发明所用的负载金属没有特别限制，可以为锌、镓和/或铁，也可以包括其它金属。
在某些进一步优选的实施方式中，所述负载金属为锌和/或镓，所述天然矿物质包括选自高岭土、多水高岭土、蒙脱土、硅藻土、凸凹棒石、海泡石、埃洛石、水滑石、膨润土和累托土中的至少一种，所述无机氧化物粘结剂包括选自氧化硅、氧化铝、氧化锆、氧化钛和无定形硅铝中的至少一种。
在某些更进一步优选的实施方式中，所述MFI结构分子筛为选自ZSM-5、ZSM-8和ZSM-11中的至少一种。
在某些优选的实施方式中，所述芳构化催化剂中所述MFI结构分子筛的n(SiO2)/n(Al2O3)大于约100；以P2O5计并以分子筛的干基重量为基准，所述分子筛的磷含量为约0.1-5重量％；以负载金属的氧化物 计并以分子筛的干基重量为基准，所述分子筛的负载金属含量为约0.5-5重量％；所述分子筛的Al分布参数D(Al)满足：0.6≤D(Al)≤0.85，其中，D(Al)＝Al(S)/Al(C)，Al(S)表示采用TEM-EDS(其中，TEM为透射电镜，EDS为X射线能谱，TEM-EDS即配有X射线能谱的透射电子显微镜)方法测定的分子筛晶粒的晶面边沿向内H距离内任意大于100平方纳米区域的铝含量，Al(C)表示采用TEM-EDS方法测定的分子筛晶粒所述晶面的几何中心向外H距离内任意大于100平方纳米区域的铝含量，其中所述H为所述晶面边沿某点到该晶面几何中心距离的10％；所述分子筛的负载金属分布参数D(M)满足：2≤D(M)≤10，其中，D(M)＝M(S)/M(C)，M(S)表示采用TEM-EDS方法测定的分子筛晶粒的晶面边沿向内H距离内任意大于100平方纳米区域的负载金属含量，M(C)表示采用TEM-EDS方法测定的分子筛晶粒所述晶面的几何中心向外H距离内任意大于100平方纳米区域的负载金属含量；所述分子筛的中孔体积占总孔体积的比例为约40-80体％，孔径为约2纳米至约20纳米的中孔体积优选占总中孔体积的比例大于约90体％；所述分子筛的强酸酸量占总酸量的比例为约60-80％，Bronsted酸(B酸)酸量与Lewis酸(L酸)酸量之比为约15-80。
在某些进一步优选的实施方式中，所述MFI结构分子筛的n(SiO2)/n(Al2O3)大于约120；以P2O5计并以分子筛的干基重量为基准，所述分子筛的磷含量为约0.2-4重量％；以负载金属的氧化物计并以分子筛的干基重量为基准，所述分子筛的负载金属含量为约0.5-3重量％；所述分子筛的Al分布参数D(Al)满足：0.65≤D(Al)≤0.8；所述分子筛的负载金属分布参数D(M)满足：3≤D(M)≤6；所述分子筛的中孔体积占总孔体积的比例为约50-70体％，孔径为约2纳米至约20纳米的中孔体积占总中孔体积的比例大于约92体％；所述分子筛的强酸酸量占总酸量的比例为约65-75％，且B酸酸量与L酸酸量之比为约20-50。
本发明用于测定分子筛的铝含量和负载金属含量的TEM-EDS方法是本领域技术人员所熟知的，其中所述几何中心的确定也是本领域技术人员所熟知的，可以根据公式计算得到，本发明不再赘述。一般对称图形的几何中心为各相对顶点连线的交点，例如，常规六方片形 ZSM-5分子筛的六边形晶面的几何中心在三个相对顶点连线的交点处，所述的晶面为规则晶粒的一个面，所述向内和向外的方向均指在所述晶面上的向内和向外的方向。
根据本发明，所述分子筛的中孔体积占总孔体积的比例采用氮吸附BET比表面积方法进行测量，所述中孔体积为孔径大于约2纳米小于约100纳米的孔体积；所述分子筛的强酸酸量占总酸量的比例采用NH3-TPD方法进行测量，所述强酸的酸中心为NH3脱附温度大于300℃所对应的酸中心；所述B酸酸量与L酸酸量之比采用吡啶吸附红外酸性方法进行测量。
在某些优选的实施方式中，所述含磷和负载金属的MFI结构分子筛可以通过如下步骤制备得到：
a、将晶化所得MFI结构分子筛浆液进行过滤和洗涤后，得到水洗分子筛；其中，以氧化钠计并以水洗分子筛的总干基重量为基准，所述水洗分子筛的钠含量小于3重量％；
b、将步骤a中所得水洗分子筛在碱溶液中进行脱硅处理，并进行过滤和洗涤后，得到脱硅分子筛；
c、将步骤b中所得脱硅分子筛进行铵交换处理，得到铵交换分子筛；其中，以氧化钠计并以铵交换分子筛的总干基重量为基准，所述铵交换分子筛的钠含量小于0.2重量％；
d、将步骤c中所得铵交换分子筛在由氟硅酸、有机酸和无机酸组成的复合酸脱铝剂溶液中进行脱铝处理，并进行过滤和洗涤后，得到脱铝分子筛；
e、将步骤d中所得的脱铝分子筛进行磷改性处理和负载金属的负载处理后，得到改性分子筛；以及
f、将步骤e中所得改性分子筛进行水热焙烧处理，得到所述含磷和负载金属的MFI结构分子筛。
本发明步骤a中所用的“晶化所得MFI结构分子筛浆液”可以根据本领域技术人员所熟知的方法得到，在此不再赘述。此外，所述MFI结构分子筛也是本领域技术人员所熟知的，可以是无胺晶化所得，也可以是由模板剂法制备的分子筛，其中无胺晶化所得的分子筛不需焙烧，由模板剂法制备的分子筛需干燥后在空气中焙烧，ZSM-5分子筛的硅铝比一般小于100。
本发明步骤b中采用碱溶液进行的脱硅处理是本领域技术人员所熟知的，例如，步骤b中所述碱溶液可以为选自氢氧化钠溶液和/或氢氧化钾溶液，优选为氢氧化钠溶液，所述脱硅处理的条件可以包括：以干基重量计的分子筛与碱溶液中的碱的重量比为约1∶(0.1-1)，优选为约1∶(0.15-0.4)；脱硅处理的温度为室温至约100℃，优选为约50-85℃，时间为约15分钟至约8小时，优选为约30分钟至约4小时。
本发明步骤c中进行的铵交换处理是本领域技术人员所熟知的，例如，步骤c中可以将碱处理后的脱硅分子筛按照分子筛∶铵盐∶H2O为约1∶(0.1-1)∶(5-10)的重量比在室温至约100℃下交换约0.5-2小时后过滤，使沸石上的Na2O含量小于约0.2重％。所述的铵盐可以为常用的无机铵盐，例如，选自氯化铵、硫酸铵和硝酸铵中的至少一种。
本发明步骤d中涉及的“脱铝处理”本身是本领域技术人员所熟知的，但未报道过如本发明步骤d中那样将无机酸、有机酸和氟硅酸一起用于脱铝处理。所述脱铝处理可以一次或分多次进行，可以先将有机酸与所述铵交换分子筛混合，然后将氟硅酸和无机酸与所述铵交换分子筛混合，即可以为先将有机酸加入铵交换分子筛中，然后将氟硅酸和无机酸慢速并流加入，或先加入氟硅酸再加入无机酸，优选为氟硅酸和无机酸慢速并流加入。所述脱铝处理可在如下条件下进行：以干基重量计的分子筛、氟硅酸、有机酸和无机酸的重量之比为约1∶(0.02-0.5)∶(0.05-0.5)∶(0.05-0.5)，优选为约1∶(0.05-0.3)∶(0.1-0.3)∶(0.1-0.3)；处理温度为约25-100℃，处理时间为约0.5-6小时。
实验证明，本发明采用复合酸体系进行脱铝处理，在三种酸的协同作用下，能够在保证分子筛晶体结构和介孔孔道结构完整性的前提下有效提高分子筛硅铝比，调节铝分布，并改善酸性分布。
本发明步骤d中采用的有机酸和无机酸可以是本领域中常用的那些，例如，所述有机酸可以为选自乙二胺四乙酸、草酸、柠檬酸和磺基水杨酸中的至少一种，优选为草酸；无机酸可以为选自盐酸、硫酸和硝酸中的至少一种，优选为盐酸。
本发明步骤d中进行的所述洗涤是本领域技术人员所熟知的，例如，其方式可以为：将约5-10倍约30-60℃的水对过滤后分子筛进行淋洗。
本发明步骤e中进行的磷改性处理和负载金属的负载处理是本领域技术人员所熟知的，例如，步骤e中所述磷改性处理可以包括：将选自磷酸、磷酸氢铵、磷酸二氢铵和磷酸铵中的至少一种含磷化合物对分子筛进行浸渍和/或离子交换；步骤e中所述负载金属的负载处理可以包括：将含有选自锌和镓中至少一种负载金属的可溶性盐溶解于去离子水中，用氨水调节pH值使负载金属以氢氧化物的形态沉淀出来，然后将所得沉淀物与分子筛混合均匀。
本发明步骤f中进行的分子筛的水热焙烧处理是本领域技术人员所熟知的，例如，步骤f中所述水热焙烧处理可在如下条件下进行：焙烧处理的气氛为水蒸汽气氛；焙烧温度为约400-800℃，焙烧时间为约0.5-8小时。
在某些优选的实施方式中，本发明的发明人意外地发现，所述芳构化催化剂的微反活性在约20-55时具有更好的效果，所述微反活性采用RIPP 92-90催化裂化工业平衡催化剂的微反活性试验法进行测定，具体测定过程可参见实施例部分。通常，自生产商处商购的或者自制的合格的新鲜芳构化催化剂活性一般在60以上，活性较高，裂化性能较强。因此，在用于所述脱硫和芳构化反应之前，优选对这类新鲜芳构化催化剂进行处理，以减少酸量，并增加酸强度，从而有利于减少氢转移反应的发生。
在某些优选的实施方式中，在用于所述脱硫和芳构化反应之前，所述芳构化催化剂经过钝化处理。优选地，所述钝化处理包括：使芳构化催化剂，如新鲜芳构化催化剂，与含碳、硫和/或氮的化合物接触并进行钝化反应。
在某些进一步优选的实施方式中，经钝化处理的所述芳构化催化剂包含约0.1-5.0重量％的钝化物，其中所述钝化物包括选自碳、硫和氮中的至少一种元素。
在某些优选的实施方式中，所述钝化处理可以在位于脱硫-芳构化反应器外的反应器中进行，也可以在用作脱硫-芳构化反应器的流态化反应器的预提升段中进行；钝化处理所用的含碳、硫和/或氮的化合物可以为汽油、硫化氢、二硫化碳、氨气、苯胺类、吡啶类和喹啉类等，该汽油既可以为所述汽油原料，也可以为来自外部的其它汽油，例如催化裂化汽油、蒸汽裂解汽油或其它含烯烃的汽油。
在某些优选的实施方式中，在用于所述脱硫和芳构化反应之前，所述芳构化催化剂经过老化处理。在本发明中，所述“老化处理”是指在水蒸汽存在的条件下对芳构化催化剂，如新鲜芳构化催化剂，进行高温处理，以降低其活性。在某些进一步优选的实施方式中，所述老化处理在如下条件下进行：温度为约500-800℃，优选约600-800℃，更优选约700-800℃；时间为约1-360小时，优选约2-48小时，更优选约4-24小时；老化气氛为包含水蒸汽的气氛，优选为约100％的水蒸汽气氛。
在所述老化处理的某些优选实施方式中，将新鲜芳构化催化剂与水蒸汽或含水蒸汽的老化介质接触，在温度为约500℃-800℃的水热环境下，老化约1小时至360小时后，得到经老化的芳构化催化剂。在某些进一步优选的实施方式中，所述老化介质包括空气、干气、再生烟气、空气与干气燃烧后的气体或空气与燃烧油燃烧后的气体、或其他气体如氮气。优选地，所述含水蒸汽的老化介质中的水蒸汽与老化介质的重量比为约0.2-0.9，更优选为约0.40-0.60。优选地，所述再生烟气可以来自本发明的催化剂再生装置，也可以来自其他过程的再生装置。
在某些优选的实施方式中，所述老化处理在老化反应器，优选密相流化床中进行。在某些进一步优选的实施方式中，所述老化处理包括将新鲜芳构化催化剂在800℃和100％水蒸汽条件下老化4h至17h的时间。
在某些优选的实施方式中，至少约50wt％、优选至少约80wt％、更优选至少约90wt％、特别优选至少约95wt％、最优选约100wt％的所述芳构化催化剂经过钝化和/或老化处理。
本发明的脱硫和芳构化反应优选在流态化反应器中进行，以利于芳构化催化剂的快速再生。在某些优选的实施方式中，所述流态化反应器为选自流化床、提升管、下行式输送线反应器、由提升管与流化床构成的复合反应器、由提升管与下行式输送线构成的复合反应器、由两个或两个以上的提升管构成的复合反应器、由两个或两个以上的流化床构成的复合反应器和由两个或两个以上的下行式输送线构成的复合反应器，优选为提升管反应器、流化床反应器或其组合，优选地，上述每种反应器可以分成两个或两个以上的反应区。在某些优选的实 施方式中，所述流化床反应器为选自固定流化床、散式流化床、鼓泡床、湍动床、快速床、输送床和密相流化床中的一种或多种；所述提升管反应器为选自等直径提升管、等线速提升管和各种变直径提升管中的一种或多种。在特别优选的实施方式中，所述流态化反应器为密相流态化反应器。
本发明所用的汽油原料可以是任何本领域常用的常规汽油原料。在某些优选的实施方式中，所述汽油原料可以为选自催化裂化汽油、催化裂解汽油、焦化汽油、热裂化汽油和直馏汽油中的至少一种或其部分馏分。在某些优选的实施方式中，所述汽油原料为高烯烃和高硫汽油，其烯烃体积分数一般大于约20体％，优选大于约30体％，更优选大于约40体％，进一步优选大于约50体％；硫含量一般在约10μg/g以上，优选大于约50μg/g，更优选大于约100μg/g，进一步优选大于约500μg/g，更进一步优选大于约1000μg/g。所述汽油原料中的有机硫化物没有特别的限制，例如可以为硫醇、硫醚、噻吩、烷基噻吩、苯并噻吩和/或甲基苯并噻吩等。
根据需要，本发明中所述轻汽油馏分和重汽油馏分的切割点可以改变。在某些优选的实施方式中，所述轻汽油馏分和重汽油馏分的切割点为60-80℃，更优选为约65-70℃。在某些进一步优选的实施方式中，在分馏塔中按照馏程从低到高对所述汽油全馏分进行切割，更优选地，汽油切割的分馏塔的操作条件为：塔顶温度为约60-80℃，塔底温度为约120-160℃，操作压力为约0.05-0.3MPa。
在本发明中，所述“醚化处理”是指使轻汽油馏分中C5以下的低碳烃类(例如异戊烯和环戊烯)与醇类发生醚化反应，以生成高辛烷值的醚化油。在某些优选的实施方式中，所述醚化处理包括如下步骤：将所述轻汽油馏分与醇类接触，使所述轻汽油馏分中的烯烃在醚化催化剂的作用下与醇类发生醚化反应，得到所述醚化油；其中，所述醚化反应的温度为约20-200℃，压力为约0.1-5MPa，重时空速为约0.1-20小时-1，所述醇类与轻汽油馏分的摩尔比为约1∶(0.1-100)。在某些进一步优选的实施方式中，所述醚化催化剂包括选自树脂、分子筛和杂多酸中的至少一种。在某些更进一步优选的实施方式中，所述醇类为选自甲醇、乙醇和丙醇中的至少一种。
在所述醚化处理的某些优选实施方式中，将强酸性阳离子交换树 脂催化剂装填于一段醚化和/或二段醚化固定床反应器中，将经过脱硫和脱二烯烃等预处理的轻汽油馏分通入该醚化固定床反应器中，在反应温度为约50-90℃，液时空速为约1.0-3.0h-1，甲醇与轻汽油馏分中活性烯烃的摩尔比为约1-2的条件下进行醚化反应，醚化产物送入精馏塔分离，在塔底得到醚化油，未反应的轻烃和甲醇循环利用。在本发明中，所述“活性烯烃”是指在叔碳原子上有一个双键的烯烃。
在某些进一步优选的实施方式中，醚化过程的反应温度为入口温度约55-60℃，出口温度小于约90℃，空速为约1-2h-1，甲醇和轻汽油馏分中活性烯烃的摩尔比为约1.2-1.4。在某些实施方式中，一段醚化烯烃含量较高，适宜采用混相床反应器，二段醚化烯烃含量较低，适宜采用绝热固定床反应器。在某些优选的实施方式中，异构化单元也可以应用到轻汽油醚化过程中。轻汽油醚化具有降低汽油烯烃含量、提高辛烷值、降低蒸汽压以及提高附加值和强化调和效益等诸多优点，醚化油可以作为汽油辛烷值调和组分进行使用，也可以与重汽油馏分混合作为全馏分汽油产品。
在某些优选的实施方式中，所述轻汽油馏分在进行醚化反应之前经过预处理，以脱除硫化物和/或二烯烃等杂质，从而延长醚化催化剂的寿命。在某些进一步优选的实施方式中，所述预处理为选自碱液抽提处理、硫醇转化处理和选择性加氢处理中的至少一种。在某些更进一步优选的实施方式中，所述碱液抽提处理使用碱液将轻汽油馏分的硫醇抽提到碱液中而脱除；所述硫醇转化处理将小分子硫醇转化为其它硫化物而脱除，其可以采用常规的无碱脱臭工艺和预加氢等方式进行，其中所用的催化剂和助催化剂均可以为本领域常用的催化剂。所述选择性加氢处理是本领域技术人员所熟知的，用于脱除汽油中二烯烃，并可以将3-甲基-1-丁烯异构化为2-甲基-1-丁烯。
在本发明中，由于所述轻汽油馏分在醚化前已经过脱硫和芳构化处理，因此上述预处理过程并不是必需的，在某些实施方式中可以省略。
在第三方面，本发明提供了一种汽油的处理系统，该系统包括脱硫-芳构化反应器、高压分离器、汽油全馏分分馏塔和醚化装置，其中所述脱硫-芳构化反应器设置有汽油原料入口、氢气入口、以及脱硫-芳构化产物出口，所述汽油全馏分分馏塔设置有汽油全馏分入口、轻 汽油馏分出口和重汽油馏分出口；所述脱硫-芳构化反应器的脱硫-芳构化产物出口经由所述高压分离器与所述汽油全馏分分馏塔的汽油全馏分入口流体连通；所述汽油全馏分分馏塔的轻汽油馏分出口与醚化装置的入口流体连通，
其中汽油原料经所述汽油原料入口进入所述脱硫-芳构化反应器，在其中在氢气存在下进行脱硫和芳构化反应，得到脱硫-芳构化产物；所述脱硫-芳构化产物在所述高压分离器分离得到含氢尾气和汽油全馏分，所述汽油全馏分经汽油全馏分入口进入所述汽油全馏分分馏塔，在其中分馏得到轻汽油馏分和重汽油馏分；所述轻汽油馏分进入所述醚化装置，在其中进行醚化反应，得到醚化产物。
在某些优选的实施方式中，所述系统进一步包括醚化产物分馏塔和混合器，其中所述醚化产物分馏塔设置有醚化产物入口、含醇尾气出口和醚化油出口；所述醚化装置的出口与所述醚化产物分馏塔的醚化产物入口流体连通，所述醚化产物分馏塔的醚化油出口与所述混合器的入口流体连通；且所述汽油全馏分分馏塔的重汽油馏分出口与所述混合器的入口流体连通，
其中所述醚化产物在所述醚化产物分馏塔中分馏得到含醇尾气和醚化油，所述醚化油和所述重汽油馏分在所述混合器中混合得到汽油产品。
在某些优选的实施方式中，所述脱硫-芳构化反应器的汽油原料入口和氢气入口为同一入口。
下面将结合附图描述本发明方法和系统的一种优选实施方式，但是并不因此而限制本发明。
如图1所示，在一优选实施方案中，本发明的汽油处理系统包括：脱硫-芳构化反应器3、高压分离器5、汽油全馏分分馏塔8、预处理单元11、醚化装置14、醚化产物分馏塔16和混合器19，其中所述脱硫-芳构化反应器3设置有汽油原料入口、氢气入口、以及脱硫-芳构化产物出口，所述汽油全馏分分馏塔8设置有汽油全馏分入口、轻汽油馏分出口和重汽油馏分出口，所述醚化产物分馏塔16设置有醚化产物入口、含醇尾气出口和醚化油出口；所述脱硫-芳构化反应器3的脱硫和芳构化产物出口经由所述高压分离器5与所述汽油全馏分分馏塔8的汽油全馏分入口流体连通；所述汽油全馏分分馏塔8的轻汽油馏分出口依次通过所述预处理单元11和醚化装置14与所述醚化产物分馏塔 16的醚化产物入口流体连通，所述醚化产物分馏塔16的醚化油出口与所述混合器19的入口流体连通；所述汽油全馏分分馏塔8的重汽油馏分出口与所述混合器19的入口流体连通。
在图1所示的优选实施方式中，本发明方法包括将将高烯烃高硫的汽油原料1与氢气2混合后送入脱硫-芳构化反应器3，在其中与吸附脱硫催化剂和芳构化催化剂接触进行吸附脱硫和芳构化反应，反应后脱硫-芳构化产物4进入高压分离器5，得到低硫的汽油全馏分7与含氢尾气6。汽油全馏分7进入汽油全馏分分馏塔8进行切割分馏，得到轻汽油馏分9和低硫的重汽油馏分10。轻汽油馏分9经预处理单元11进行脱硫醇等预处理得到醚化前轻汽油12，然后与含醇物流13混合进入醚化装置14进行反应，醚化产物15经醚化产物分馏塔16分馏得到醚化油18与含醇尾气17。重汽油馏分10与醚化处理后的醚化油18在混合器19中混合得到高辛烷值清洁汽油产品20。
实施例
下面将结合实施例将对本发明予以进一步说明，但并不因此而限制本发明。
测量方法
本发明中，结晶度采用ASTM D5758-2001(2011)e1的标准方法进行测定。
本发明中，n(SiO2)/n(Al2O3)，即硅铝比通过氧化硅和氧化铝的含量计算得到，氧化硅和氧化铝的含量采用GB/T 30905-2014标准方法进行测定。
本发明中，磷含量采用GB/T 30905-2014标准方法进行测定，负载金属的含量采用GB/T 30905-2014标准方法进行测定，钠含量采用GB/T 30905-2014标准方法进行测定。
本发明所用的TEM-EDS测定方法参见“固体催化剂的研究方法”，薛用芳，石油化工，29(3)，2000：227-235页。
本发明中，总比表面积(SBET)、中孔孔体积、总孔体积、2-20纳米的中孔孔体积的测定方法如下所示：
采用Quantachrome仪器公司生产的AS-3，AS-6静态氮吸附仪测定。
仪器参数：将样品置于样品处理系统，在300℃下抽真空至 1.33×10-2Pa，保温保压4h，净化样品。在液氮温度-196℃下，测试净化样品在不同比压P/P0条件下对氮气的吸附量和脱附量，获得N2吸附-脱附等温曲线。然后利用两参数BET公式计算总比表面积、微孔比表面积和中孔比表面积，取比压P/P0＝0.98以下的吸附量为样品的总孔体积，利用BJH公式计算中孔部分的孔径分布，并采用积分法计算中孔孔体积(2-100纳米)和2-20纳米的中孔孔体积。
本发明中，B酸酸量和L酸酸量的测定方法如下所示：
采用美国BIO-RAD公司生产的FTS3000型傅里叶红外光谱仪。
测试条件：将样品压制成片后置于红外光谱仪的原位池中密封，在350℃下抽真空至10-3Pa，保持1h，使样品表面的气体分子脱附干净，冷却至室温。向原位池中导入压力为2.67Pa的吡啶蒸汽，平衡30min后，升温至200℃，再次抽真空至10-3Pa，保持30min，冷却至室温，在1400-1700cm-1波数范围内扫描，记录下200℃吡啶吸附的红外光谱谱图。再将红外吸收池中的样品移至热处理区，升温至350℃，抽真空至10-3Pa，保持30min，冷至室温，记录下350℃吡啶吸附的红外谱图。仪器自动积分得到B酸酸量和L酸酸量。
本发明中，总酸量和强酸酸量的测定方法如下所示：
采用美国麦克公司Autochem II 2920程序升温脱附仪。
测试条件：称取0.2g待测样品装入样品管，置于热导池加热炉，He气为载气(50mL/min)，以20℃/min的速率升温至600℃，吹扫60min驱除催化剂表面吸附的杂质。然后降温至100℃，恒温30min，切换成NH3-He混合气(10.02％NH3+89.98％He)吸附30min，再继续以He气吹扫90min至基线平稳，以脱附物理吸附的氨气。以10℃/min升温速率升温至600℃进行脱附，保持30min，脱附结束。采用TCD检测器检测气体组分变化，仪器自动积分得到总酸量和强酸酸量，强酸的酸中心为NH3脱附温度大于300℃所对应的酸中心。
本发明中，所述D值的计算方法如下：
在透射电镜中选取一个晶粒以及该晶粒的某个晶面所形成一个多边形，该多边形存在几何中心、边沿以及几何中心到边沿某点的10％距离H(不同的边沿点，H值不同)，分别选取该晶面边沿向内H距离内的任意一块大于100平方纳米区域以及晶面几何中心向外H距离内的任意一块大于100平方纳米区域，测定铝含量，即为Al(S1)和 Al(C1)，并计算D(Al)1＝Al(S1)/Al(C1)，分别选取不同的晶粒测定5次，计算平均值即为D(Al)，D(M)的测定方法与D(Al)的测定方法类似。
本发明中，干基重量的测试方法为：
将待测分子筛或催化剂置于马弗炉中在600℃下的空气气氛中进行焙烧3小时，将得到的焙烧产物在密闭的干燥皿中冷却至室温，然后称重。
下面的实施例和对比例中所使用的吸附脱硫催化剂由中国石油化工股份有限公司催化剂分公司生产，商品牌号为FCAS，所使用的芳构化催化剂包括实验室自制催化剂，编号为OTAZ-C-3，和商购芳构化催化剂，商品牌号为MP051，由中国石油化工股份有限公司催化剂分公司生产。实施例中采用的各催化剂的性质列于表1。
表1实施例中采用的各催化剂的物性数据
催化剂 FCAS OTAZ-C-3 MP051
化学组成，重量％      
氧化铝 13 50.3 50.8
氧化钠 / / /
氧化镍 21 / /
氧化锌 52 / /
氧化镓 / 1.5 /
氧化硅 14 48.14 48.7
表观密度，kg/m3 1010 800 790
孔体积，mL/g / 0.27 0.23
比表面积，m2/g / 218 215
磨损指数，重量％时-1 / 1.5 1.4
筛分组成，重量％      
0～40微米 14.5 13.9 15
40～80微米 51.9 49.5 50.5
＞80微米 33.6 36.6 34.5
微反活性 / 80 80
OTAZ-C-3芳构化催化剂的制备方法如下所示：
所用原材料的性质如下：高岭土(苏州中国高岭土公司，固含量75重量％)，拟薄水铝石(山东铝业公司，固含量65重量％，使用时先用浓度31重量％盐酸胶溶，所述的盐酸与以氧化铝计的拟薄水铝石的摩尔比为0.20)。
将晶化好的ZSM-5分子筛(催化剂齐鲁分公司生产，无胺法合成，n(SiO2)/n(Al2O3)＝27)过滤掉母液，水洗至Na2O含量低于3.0重量％，过滤得滤饼；取上述分子筛100g(干基)加入1000g 2.0％的NaOH溶液中，升温至65℃，反应30min后，快速冷却至室温后，过滤，洗涤至滤液中性。然后，将滤饼加入800g水打浆，加入40g NH4Cl，升温至75℃，交换处理1h后，至Na2O含量低于0.2重量％，过滤，洗涤，得到分子筛滤饼；取上述分子筛50g(干基)加水配制成固含量10重％的分子筛浆液，搅拌中加入草酸11g，然后将110g盐酸(质量分数10％)和92g氟硅酸(质量分数3％)并流加入，加入时间30min；升温至65℃恒温搅拌1h，过滤水洗至滤液中性；将滤饼加水打浆得固含量为45重量％的分子筛浆液；将1.2g H3PO4(浓度85重量％)和3.3gZn(NO3)2·6H2O溶于10g水中，加入氨水调节pH＝6，然后加入分子筛浆液中混合均匀、烘干、100％水蒸汽气氛下550℃焙烧处理2h，得分子筛A，性质列于表2。
将拟薄水铝石与高岭土混合，并用脱阳离子水将其配制成固含量为30重量％的浆液，搅拌均匀，用盐酸将浆液pH调至2.5，保持该pH值，于50℃下静置老化1小时后搅拌1小时形成胶体，加入所制备的分子筛A和水，形成催化剂浆液(固含量为35％重量记)。继续搅拌后喷雾干燥制成微球催化剂。然后将微球催化剂在500℃焙烧1小时，得到本发明所用的OTAZ-C-3芳构化催化剂，其干基原料配比为25重量％高岭土、25重量％拟薄水铝石和50重量％分子筛A。
表2分子筛A的物性数据
项目 分子筛A
结晶度/％ 90
n(SiO2)/n(Al2O3) 110
P2O5含量/％ 1.5
负载金属氧化物含量/％ 1.6
SBET/(m2/g) 440
(V中孔/V总孔)/％ 60
(V2nm-20nm/V中孔)/％ 99
(强酸酸量/总酸量)/％ 70
B酸酸量/L酸酸量 35
D(Al分布) 0.77
D(M分布) 3.6
芳构化催化剂的钝化处理如下所示：
1、将编号为OTAZ-C-3的新鲜芳构化催化剂装入密相流化床反应器，在钝化温度410℃条件下与稳定汽油A接触反应，连续钝化2小时，得到经钝化的OTAZ-C-3催化剂，钝化物含量为0.5重量％，微反活性为35。
2、将牌号为MP051的新鲜芳构化催化剂装入密相流化床反应器，在钝化温度410℃条件下与稳定汽油A接触反应，连续钝化2小时，得到经钝化的MP051催化剂，钝化物含量为0.5重量％，微反活性为35。
芳构化催化剂的老化处理如下所示：
1、将牌号为OTAZ-C-3的新鲜芳构化催化剂装入密相流化床反应器，在老化温度780℃、100％水蒸汽条件下连续老化8小时，得到经老化的OTAZ-C-3催化剂，微反活性为40。
2、将牌号为MP051的新鲜芳构化催化剂装入密相流化床反应器，在老化温度780℃、100％水蒸汽条件下连续老化8小时，得到经老化的MP051催化剂，微反活性为40。
本发明各实施例和对比例中，催化剂中Na2O、NiO、ZnO、Ga2O3、Al2O3、SiO2的含量用X射线荧光法测定，具体测试方法参见采用GB/T 30905-2014。
本发明各实施例和对比例中，钝化芳构化催化剂中的钝化物含量采用RIPP 107-90方法进行测定，具体步骤概述如下：试样在高频炉中燃烧，生成CO2、CO气体，气体经过滤干燥器进入红外鉴定器，给出总炭的百分含量。由于催化剂本身无感应效果，需加感应剂，助熔剂，促进样品燃烧，又因样品系多孔物质，具有吸水性，测试前需进行预处理。具体测试试验步骤包括：
1、试验准备：按照CS-46型碳硫测试仪(美国LECO公司生产)的操作规程进行预热、清洗，试漏并检查电路。催化剂样品在110℃烘2-3小时，至于干燥器内冷却到室温。
2、测试：燃烧标准铜样，检查C％值，调节读数至误差范围之内。将坩埚放在电子天平上，去皮回零。加已烘过的催化剂样品0.3-0.4克，将样品重量自动输入测试仪。取下坩埚加一片锡箔放回到电子天平上，去皮回零。加铜助熔剂(山东冶金研究所产品)1.5-1.6克，取下坩埚后加盖，用坩埚钳将坩埚放在燃烧座上，燃烧30秒钟后，读数指示灯亮，记录C％值，每个样品重复测定两次。按同样步骤测定本次试验的碳空白值。按下式计算催化剂炭含量：C％＝样品C％-空白C％。
本发明各实施例和对比例中，芳构化催化剂的微反活性测试方法采用RIPP 92-90方法，其中采用与ASTM D3907-2013相同的设备和测试流程，在如下原料油和测试反应条件下进行，原料油：大港直馏柴油，相对密度d4 20为0.8419，初馏点235℃，干点337℃；反应条件：进油量1.56克，反应温度460℃，剂油质量比3.2，原料油重时空速16h-1，进料时间70秒。
本发明各实施例和对比例中，汽油的辛烷值RON和MON分别采用GB/T 5487-1995和GB/T 503-1995标准方法进行测定，抗爆指数＝(MON+RON)/2，汽油PONA采用模拟蒸馏和汽油单体烃分析(分别采用ASTM D2887和ASTM D6733-01(2011)测试方法进行测试)，汽油硫含量采用SH/T0689-2000进行测定。
以下的实施例和对比例中采用的汽油原料为稳定汽油A，其性质列于下表3中：
表3实施例采用的汽油原料的物性数据
Figure PCTCN2017000634-appb-000001
以下实施例I-1至实施例I-10为采用各种类型的催化剂对上述汽油原料进行脱硫和芳构化反应的实施例，反应结果列于表4中。
实施例I-1
汽油原料在小型连续流化床反应器内与包括吸附脱硫催化剂 FCAS和经钝化的OTAZ-C-3催化剂的混合催化剂接触，进行吸附脱硫和芳构化反应。其中，所述OTAZ-C-3催化剂占所述混合催化剂总重量的7重量％。
操作条件为：反应温度为400℃，反应器压力为2.1MPa，汽油原料的重时空速为6小时-1，氢气与汽油原料的体积比为75。由反应器顶部得到的脱硫-芳构化产物经冷却到10℃左右、分离得到气体产物和汽油全馏分，汽油全馏分不经切割和醚化作为汽油产品，其性质列于表4。混合催化剂的再生温度为550℃，再生后的混合催化剂返回反应器内循环使用。
实施例I-2
与实施例I-1的操作基本相同，不同之处在于，所述经钝化的OTAZ-C-3催化剂采用同等重量的经老化的OTAZ-C-3催化剂替换，汽油全馏分不经切割和醚化作为汽油产品，其性质列于表4。
实施例I-3
与实施例I-1的操作基本相同，不同之处在于，所述经钝化的OTAZ-C-3催化剂采用同等重量的新鲜的OTAZ-C-3催化剂替换，汽油全馏分不经切割和醚化作为汽油产品，其性质列于表4。
实施例I-4
与实施例I-1的操作基本相同，不同之处在于，所述经钝化的OTAZ-C-3催化剂采用同等重量的混合OTAZ-C-3催化剂(包括5重量％的钝化催化剂和2重量％的新鲜催化剂，以混合的吸附脱硫催化剂和芳构化催化剂的总重量计)替换，汽油全馏分不经切割和醚化作为汽油产品，其性质列于表4。
实施例I-5
与实施例I-1的操作基本相同，不同之处在于，所述经钝化的OTAZ-C-3催化剂采用同等重量的混合OTAZ-C-3催化剂(包括5重量％的老化催化剂和2重量％的新鲜催化剂，以混合的吸附脱硫催化剂和芳构化催化剂的总重量计)替换，汽油全馏分不经切割和醚化作为汽油产品，其性质列于表4。
实施例I-6
与实施例I-1的操作基本相同，不同之处在于，所述经钝化的OTAZ-C-3催化剂采用同等重量的经钝化的MP051催化剂替换，汽油 全馏分不经切割和醚化作为汽油产品，其性质列于表4。
实施例I-7
与实施例I-1的操作基本相同，不同之处在于，所述经钝化的OTAZ-C-3催化剂采用同等重量的经老化的MP051催化剂替换，汽油全馏分不经切割和醚化作为汽油产品，其性质列于表4。
实施例I-8
与实施例I-1的操作基本相同，不同之处在于，所述经钝化的OTAZ-C-3催化剂采用同等重量的新鲜的MP051催化剂替换，汽油全馏分不经切割和醚化作为汽油产品，其性质列于表4。
实施例I-9
与实施例I-1的操作基本相同，不同之处在于，所述经钝化的OTAZ-C-3催化剂采用同等重量的混合MP051催化剂(包括5％重量的钝化催化剂和2重量％的新鲜催化剂，以混合的吸附脱硫催化剂和芳构化催化剂的总重量计)替换，汽油全馏分不经切割和醚化作为汽油产品，其性质列于表4。
实施例I-10
与实施例I-1的操作基本相同，不同之处在于，所述经钝化的OTAZ-C-3催化剂采用同等重量的混合MP051催化剂(包括5％重量的老化催化剂和2重量％的新鲜催化剂，以混合的吸附脱硫催化剂和芳构化催化剂的总重量计)替换，汽油全馏分不经切割和醚化作为汽油产品，其性质列于表4。
对比例I-1
与实施例I-1的操作基本相同，不同之处在于，所述经钝化的OTAZ-C-3催化剂采用同等重量的吸附脱硫催化剂FCAS替换，汽油全馏分不经切割和醚化作为汽油产品，其性质列于表4。
表4实施例I-1至I-10和对比例I-1的结果数据
Figure PCTCN2017000634-appb-000002
表4(续)实施例I-1至I-10和对比例I-1的结果数据
Figure PCTCN2017000634-appb-000003
如表4所示，在汽油全馏分不经切割和醚化作为汽油产品的情况下，仅采用吸附脱硫催化剂FCAS对汽油原料进行处理时，即仅进行吸附脱硫反应，对比例I-1所得的汽油产品与汽油原料相比辛烷值(抗爆指数)大幅降低；采用包括吸附脱硫催化剂FCAS和芳构化催化剂MP051的混合催化剂对汽油原料进行处理时，即进行吸附脱硫和芳构化反应，实施例I-6至I-10与对比例I-1相比所得汽油产品的辛烷值得到了显著改善，特别是采用包含经钝化或经老化的MP051催化剂的芳构化催化剂时的改善效果更为明显；进一步，采用包括吸附脱硫催化剂FCAS和芳构化催化剂OTAZ-C-3的混合催化剂对所述重汽油馏分进行处理时，实施例I-1至I-5与实施例I-6至I-10相比所得汽油产品的辛烷值得到了进一步改善，且采用包含经钝化或经老化的OTAZ-C-3催化剂的芳构化催化剂的改善效果更为明显。
以下实施例II-1至实施例II-3为采用各种类型的催化剂对汽油原料进行脱硫和芳构化反应，随后对汽油全馏分进行分割，并对轻汽油馏分进行醚化的实施例，反应结果列于表6中。
实施例II-1
汽油原料在小型连续流化床反应器内与包括吸附脱硫催化剂FCAS和经钝化的OTAZ-C-3催化剂的混合催化剂接触，进行吸附脱硫和芳构化反应。其中，所述OTAZ-C-3催化剂占所述混合催化剂总重量的7重量％。
操作条件为：反应温度为400℃，反应器压力为2.1MPa，汽油原料的重时空速为6小时-1，氢气与汽油原料的体积比为75。由反应器顶部得到的脱硫-芳构化产物经冷却到10℃左右、分离得到尾气和汽油全馏分。混合催化剂的再生温度为550℃，再生后的混合催化剂返回反应器内循环使用。
将上述汽油全馏分在分馏塔中进行蒸馏，切割为轻汽油馏分和低硫重汽油馏分，控制轻馏分终馏点为65-70℃(根据ASTM D86标准进行)。其中，蒸馏得到的轻汽油馏分记为LCN-A-1，重汽油馏分记为HCN-A-1，蒸馏得到的轻汽油馏分和重汽油馏分的性质列于表5。
轻汽油馏分经精制反应器在氢气存在下进行脱硫、脱二烯烃预处理，使轻汽油馏分中硫含量和二烯烃含量均降至10ppm以下，然后与甲醇(分析纯)混合进入醚化反应器进行醚化反应，醚化反应条件为 反应温度55-80℃，空速1.2h-1，甲醇与轻汽油馏分中活性烯烃的摩尔比1.2，醚化催化剂为强酸性阳离子交换树脂(732型苯乙烯系)，反应得到的醚化后油气进入醚化分馏塔分离，塔顶气相为包括剩余C5和甲醇的含甲醇尾气，塔底得到醚化油，醚化分馏塔的塔顶温度为60-80℃，塔底温度为110-140℃。相应的醚化油记为LCN-A-M-1，其性质同样列于表5。
将得到的吸附脱硫后的重汽油馏分HCN-A-1与醚化油LCN-A-M-1混合得到高辛烷值清洁汽油产品，性质列于表6。
实施例II-2
汽油原料在小型连续流化床反应器内与包括吸附脱硫催化剂FCAS和经老化的OTAZ-C-3催化剂的混合催化剂接触，进行吸附脱硫和芳构化反应。其中，所述OTAZ-C-3催化剂占所述混合催化剂总重量的7重量％。
操作条件为：反应温度为400℃，反应器压力为2.1MPa，汽油原料的重时空速为6小时-1，氢气与汽油原料的体积比为75。由反应器顶部得到的脱硫-芳构化产物经冷却10℃左右、分离得到尾气和汽油全馏分。混合催化剂的再生温度为550℃，再生后的混合催化剂返回反应器内循环使用。
将上述汽油全馏分采用与实施例II-1相同的方式进行处理，所得到的轻汽油馏分记为LCN-A-2，重汽油馏分记为HCN-A-2，醚化油记为LCN-A-M-2，性质列于表5；得到的醚化油与重汽油馏分混合得到高辛烷值清洁汽油产品，性质列于表6。
实施例II-3
汽油原料在小型连续流化床反应器内与包括吸附脱硫催化剂FCAS和新鲜的OTAZ-C-3催化剂的混合催化剂接触，进行吸附脱硫和芳构化反应。其中，所述OTAZ-C-3催化剂占所述混合催化剂总重量的7重量％。
操作条件为：反应温度为400℃，反应器压力为2.1MPa，汽油原料的重时空速为6小时-1，氢气与汽油原料的体积比为75。由反应器顶部得到的脱硫-芳构化产物经冷却10℃左右、分离得到尾气和汽油全馏分。混合催化剂的再生温度为550℃，再生后的混合催化剂返回反应器内循环使用。
将上述汽油全馏分采用与实施例II-1相同的方式进行处理，所得到的轻汽油馏分记为LCN-A-3，重汽油馏分记为HCN-A-3，醚化油记为LCN-A-M-3，性质列于表5；得到的醚化油与重汽油馏分混合得到高辛烷值清洁汽油产品，性质列于表6。
对比例II-1
与实施例II-1的操作基本相同，不同之处在于，所述经钝化的OTAZ-C-3催化剂采用同等重量的吸附脱硫催化剂FCAS替换。所得汽油产品的性质列于表6。
表5实施例II-1至II-3所得的各馏分和醚化油的物性数据
Figure PCTCN2017000634-appb-000004
表6实施例II-1至II-3和对比例II-1的结果数据
Figure PCTCN2017000634-appb-000005
如表6所示，在对汽油全馏分进行分割、并将轻汽油馏分醚化所得的醚化油与重汽油馏分混合得到汽油产品的情况下，仅采用吸附脱硫催化剂FCAS的对比例II-1所得的汽油产品与原料汽油相比辛烷值(抗爆指数)显著降低；采用包括吸附脱硫催化剂FCAS和芳构化催化剂OTAZ-C-3的混合催化剂的实施例II-1至II-3与对比例II-1相比所得汽油产品的辛烷值得到了明显改善，特别是采用包含经钝化或经老化的OTAZ-C-3催化剂的芳构化催化剂时的改善效果更为明显。
此外，对比表4和表6所示的结果数据可以发现，对汽油全馏分进行分割并对轻汽油馏分进行醚化与不进行分割相比，所得汽油产品的辛烷值显著提高、烯烃含量显著降低，同时蒸汽压得到有效控制，特别是在采用新鲜芳构化催化剂的情况下，汽油收率也得到一定程度的提高。由此，在结合使用混合催化剂和醚化处理的情况下，可以高收率得到同时具有低硫含量、低烯烃含量、低蒸汽压和高辛烷值的汽油产品。
在上文的说明书中，已经参照特定的实施方式描述了本发明的构思。然而，本领域技术人员可以理解，在不脱离所附的权利要求中限定的本发明范围的情况下可以做出各种修改和变更。因此，说明书和附图应认为是说明性的，而不是限制性的，并且所有这类修改和变更应当涵盖在本发明的范围之内。
可以理解，本文为清楚起见以独立的多个实施方式的形式描述的某些特征也可以作为组合提供在单一的实施方式中。相反，为简要起见以单一实施方式的形式描述的多个不同特征也可以单独地或以任何子组合的形式提供。

Claims (18)

  1. 一种汽油的处理方法，该处理方法包括：
    使汽油原料与混合催化剂接触并在氢气存在下进行脱硫和芳构化反应，得到脱硫-芳构化产物，
    其中，所述混合催化剂包括吸附脱硫催化剂和芳构化催化剂，并且至少约50wt％、优选至少约80wt％、更优选至少约90wt％、特别优选至少约95wt％、最优选约100wt％的所述芳构化催化剂经过钝化和/或老化处理。
  2. 一种汽油的处理方法，该处理方法包括：
    1)使汽油原料与混合催化剂接触并在氢气存在下进行脱硫和芳构化反应，得到脱硫-芳构化产物；
    2)将至少一部分所得脱硫-芳构化产物进行切割，得到轻汽油馏分和重汽油馏分；
    3)将至少一部分所得轻汽油馏分进行醚化处理，得到醚化油；以及
    4)任选地，将至少一部分所述醚化油与至少一部分所述重汽油馏分混合，得到汽油产品，
    其中，所述混合催化剂包括吸附脱硫催化剂和芳构化催化剂，且所述芳构化催化剂选自新鲜芳构化催化剂、经钝化处理的芳构化催化剂、经老化处理的芳构化催化剂，或者它们的任意组合。
  3. 根据权利要求1或2所述的方法，其中，所述钝化处理包括：在用于所述脱硫和芳构化反应之前，使所述芳构化催化剂与含碳、硫和/或氮的化合物接触进行钝化反应。
  4. 根据权利要求1或2所述的方法，其中，所述老化处理包括在用于所述脱硫和芳构化反应之前，在如下条件下对所述芳构化催化剂进行处理：温度为约500-800℃，时间为约1-360小时，老化气氛为包含水蒸汽的气氛。
  5. 根据权利要求1或2所述的方法，其中，所述芳构化催化剂的微反活性为约20-55。
  6. 根据权利要求1或2所述的方法，其中，所述汽油原料具有以下特性中的至少一种：
    1)烯烃体积分数大于约20体％；
    2)硫含量在约10μg/g以上；以及
    3)选自催化裂化汽油、催化裂解汽油、焦化汽油、热裂化汽油和直馏汽油中的至少一种或其部分馏分。
  7. 根据权利要求2所述的方法，其中，所述轻汽油馏分和重汽油馏分的切割点为约60-80℃。
  8. 根据权利要求2所述的方法，其中，所述醚化处理包括如下步骤：
    将所述轻汽油馏分与醇类接触，使所述轻汽油馏分中的烯烃在醚化催化剂的作用下与醇类发生醚化反应，得到所述醚化油；
    其中，所述醚化反应的温度为约20-200℃，压力为约0.1-5MPa，重时空速为约0.1-20小时-1，所述醇类与轻汽油馏分的摩尔比为约1∶(0.1-100)，所述醚化催化剂包括选自树脂、分子筛和杂多酸中的至少一种。
  9. 根据权利要求1或2所述的方法，其中，所述脱硫和芳构化反应在流态化反应器中进行，其中所述流态化反应器选自提升管反应器和密相流化床反应器中的至少一种。
  10. 根据权利要求1或2所述的方法，其中，所述吸附脱硫催化剂含有二氧化硅、氧化铝、氧化锌和脱硫活性金属，所述脱硫活性金属为选自钴、镍、铜、铁、锰、钼、钨、银、锡和钒中的至少一种。
  11. 根据权利要求10所述的方法，其中，以所述吸附脱硫催化剂的干基重量为基准并以氧化物重量计，所述吸附脱硫催化剂中所述氧化锌占约10-90重量％，二氧化硅占约5-85重量％，氧化铝占约5-30重量％；以所述吸附脱硫催化剂的干基重量为基准并以元素重量计，所述吸附脱硫催化剂中所述脱硫活性金属的含量为约5-30重量％。
  12. 根据权利要求1或2所述的方法，其中，以干基计并以所述芳构化催化剂的总重量为基准，所述芳构化催化剂含有约10-30重量％的分子筛、约0-20重量％的芳构化活性金属氧化物和约50-90重量％的载体；其中，所述分子筛包括Y分子筛和/或MFI结构分子筛，所述芳构化活性金属为选自第IVB族的金属元素、第VB族的金属元素、第VIB族的金属元素、第VIII族的金属元素、第IB族的金属元素、第IIB族的金属元素和第IIIA族的金属元素中的至少一种，所述载体包 括氧化硅和/或氧化铝。
  13. 根据权利要求1或2所述的方法，其中，所述芳构化催化剂由以下原料制备得到：以所述原料的干基重量为基准，约15-60重量％的天然矿物质、约10-30重量％的无机氧化物粘结剂的前驱体和约20-80重量％的含磷和负载金属的MFI结构分子筛，其中所述MFI结构分子筛的特征在于：
    n(SiO2)/n(Al2O3)大于约100；以P2O5计并以分子筛的干基重量为基准，磷含量为约0.1-5重％；以负载金属的氧化物计并以分子筛的干基重量为基准，负载金属含量为约0.5-5重％；Al分布参数D(Al)满足：0.6≤D(Al)≤0.85；负载金属分布参数D(M)满足：2≤D(M)≤10；中孔体积占总孔体积的比例为约40-80体％；强酸酸量占总酸量的比例为约60-80％；Bronsted酸酸量与Lewis酸酸量之比为约15-80。
  14. 根据权利要求13所述的方法，其中，所述负载金属为锌，镓和铁中的至少一种。
  15. 根据权利要求1或2所述的方法，其中，以重量计，所述芳构化催化剂占所述混合催化剂的比例为约1-30重量％。
  16. 根据权利要求1或2所述的方法，其中，所述脱硫和芳构化反应在如下条件下进行：反应温度为约350-500℃，重时空速为约2-50小时-1，反应压力为约0.5-3.0MPa，氢气与汽油原料体积比为约1-500。
  17. 一种汽油的处理系统，该系统包括脱硫-芳构化反应器(3)、高压分离器(5)、汽油全馏分分馏塔(8)和醚化装置(14)，其中所述脱硫-芳构化反应器(3)设置有汽油原料入口、氢气入口、以及脱硫-芳构化产物出口，所述汽油全馏分分馏塔(8)设置有汽油全馏分入口、轻汽油馏分出口和重汽油馏分出口；所述脱硫-芳构化反应器(3)的脱硫-芳构化产物出口经由所述高压分离器(5)与所述汽油全馏分分馏塔(8)的汽油全馏分入口流体连通；所述汽油全馏分分馏塔(8)的轻汽油馏分出口与醚化装置(14)的入口流体连通，
    其中汽油原料(1)经所述汽油原料入口进入所述脱硫-芳构化反应器(3)，在其中在氢气存在下进行脱硫和芳构化反应，得到脱硫-芳构化产物(4)；所述脱硫-芳构化产物(4)在所述高压分离器(5)分离得到含氢尾气(6)和汽油全馏分(7)，所述汽油全馏分(7)经汽油 全馏分入口进入所述汽油全馏分分馏塔(8)，在其中分馏得到轻汽油馏分(9)和重汽油馏分(10)；所述轻汽油馏分(9)进入所述醚化装置(14)，在其中进行醚化反应，得到醚化产物(15)。
  18. 根据权利要求17所述的处理系统，其中，所述系统进一步包括醚化产物分馏塔(16)和混合器(19)，其中所述醚化产物分馏塔(16)设置有醚化产物入口、含醇尾气出口和醚化油出口；所述醚化装置(14)的出口与所述醚化产物分馏塔(16)的醚化产物入口流体连通，所述醚化产物分馏塔(16)的醚化油出口与所述混合器(19)的入口流体连通；且所述汽油全馏分分馏塔(8)的重汽油馏分出口与所述混合器(19)的入口流体连通，
    其中所述醚化产物(15)在所述醚化产物分馏塔(16)中分馏得到含醇尾气(17)和醚化油(18)，所述醚化油(18)和所述重汽油馏分(10)在所述混合器(19)中混合得到汽油产品。
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