WO2009062367A1 - A process for producing propylene - Google Patents

A process for producing propylene Download PDF

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Publication number
WO2009062367A1
WO2009062367A1 PCT/CN2008/000490 CN2008000490W WO2009062367A1 WO 2009062367 A1 WO2009062367 A1 WO 2009062367A1 CN 2008000490 W CN2008000490 W CN 2008000490W WO 2009062367 A1 WO2009062367 A1 WO 2009062367A1
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Prior art keywords
reaction
ethylene
reaction zone
propylene
catalyst
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PCT/CN2008/000490
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English (en)
French (fr)
Inventor
Yue Qi
Zhongmin Liu
Jinzhe Li
Zhihui Lv
Lixin Yang
Peng Tian
Bing Li
Cuiyu Yuan
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Dalian Institute Of Chemical Physics, Chinese Academy Of Sciences
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Publication of WO2009062367A1 publication Critical patent/WO2009062367A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C1/00—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon
    • C07C1/20—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon starting from organic compounds containing only oxygen atoms as heteroatoms
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C4/00—Preparation of hydrocarbons from hydrocarbons containing a larger number of carbon atoms
    • C07C4/02—Preparation of hydrocarbons from hydrocarbons containing a larger number of carbon atoms by cracking a single hydrocarbon or a mixture of individually defined hydrocarbons or a normally gaseous hydrocarbon fraction
    • C07C4/06—Catalytic processes
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/50—Improvements relating to the production of bulk chemicals
    • Y02P20/52—Improvements relating to the production of bulk chemicals using catalysts, e.g. selective catalysts
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P30/00—Technologies relating to oil refining and petrochemical industry
    • Y02P30/20—Technologies relating to oil refining and petrochemical industry using bio-feedstock
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P30/00—Technologies relating to oil refining and petrochemical industry
    • Y02P30/40—Ethylene production

Definitions

  • the present invention relates to a process for producing propylene. Background technique
  • Propylene is an important basic raw material for petrochemicals.
  • the source of propylene has long been dependent on ethylene crackers and FCC plants.
  • ethylene crackers and FCC plants As the propylene growth rate continues to exceed the ethylene growth rate, people continue to improve the traditional propylene production process to further increase propylene yield.
  • Increasing propylene production on existing plants is limited by raw material composition, plant handling capacity, plant modification and operating costs, so the development of new propylene-producing processes is an important direction to meet the growing demand for propylene.
  • High-carbon olefins especially high-carbon olefin materials containing more olefins (such as FCC gasoline and mixed C4), which are converted into propylene-based low-carbon olefins by catalytic cracking, have received extensive attention in recent years.
  • a number of patents have been published for the production of propylene by catalytic conversion of high carbon number olefins.
  • US Patent 6,222,087 B1 discloses a process for converting a feedstock containing C4-C7 olefins and alkanes to a lower olefin, the catalyst being a P-modified ZSM-5 or/and ZSM-11 molecular sieve having a silica-alumina ratio of greater than 300 .
  • the reaction conditions are at a temperature of 5I0-704 ° C, the reaction pressure is between a negative pressure and 8 bar, and the WHSV is between UOh 1 .
  • Low carbon olefin yield is higher than 20%, up to 30%, and propylene/ethylene ratio can be above 3.0
  • EP 0109059 discloses a process for the conversion of C4-C12 olefins to propylene.
  • the catalyst used is a ZSM-5 or a ZSM-11 molecular sieve having a silica to alumina ratio of less than or equal to 300, a WHSV greater than SOhr- 1 , and a reaction temperature of 400-600 °C.
  • the total yield of ethylene and propylene is 36-44%, wherein the yield of propylene is 30-40%.
  • US Patent 5,171,921 and EP0511013 A3 disclose a technique for converting high carbon mixed hydrocarbons (containing olefins and hydrazines) to lower olefins at a reaction temperature of 500-700 ° C and a WHSV between ⁇ -100 hr ⁇ 1 .
  • the catalyst contains ZSM-5 with a silica-alumina ratio of 20-60 and is subjected to P modification and steam aging treatment.
  • U.S. Patent 5,981,819 discloses a technique for converting a material containing a C4-C7 olefin to propylene. Counter The material is mixed with steam to enter the fixed bed reactor and the molecular sieve catalyst is contacted, and the feed water/oil ratio is
  • WO 01/05909 A1 discloses a process similar to that described above for converting a C4-C8 olefin containing material to a lower olefin.
  • US Patent 2003/0139636 A1 discloses a process for converting an olefin-containing material to propylene.
  • the catalysts used were rare earth or metal modified SAPO, MeAPO, MeASPO, ELAPO and ELASPO.
  • CN 1600757 discloses a process for the production of light olefins, in particular propylene, from a hydrocarbon feedstock containing C4-C6 olefins, which is contacted with a ZSM-5/ZSM-11 zeolite catalyst having a modified silica-alumina ratio of greater than 30.
  • the selectivity of light olefins is above 60%, the yield is 40-55%, and the reaction conditions are temperature 500-650 Torr, weight space velocity 1-50, pressure 0.1-8 atm.
  • CN 1490288 discloses a process for the catalytic cracking of propylene to produce propylene from C4 and above, which mainly solves the problems of low selectivity, low yield and poor catalyst stability of propylene in the target product existing in the prior art.
  • the catalyst used was ZSM-5 with a silica to alumina ratio of 50-1000, and a certain amount of sodium halide was added during the crystallization of the molecular sieve.
  • the reaction conditions are a temperature of 400-600 ° C, a liquid space velocity of 10-50 hr-l, and a pressure of 0-0.15 MPa.
  • Another new process for the production of propylene with good application prospects is to co-feed ethylene with a thiolation reagent (such as methanol or / and dimethyl ether) to generate a thiolation reaction on the catalyst to form propylene. Hydrocarbon products within.
  • a thiolation reagent such as methanol or / and dimethyl ether
  • the above type of thiolation reaction can also occur between thiolation reagents such as olefin or / and dimethyl ether.
  • thiolation reagents such as olefin or / and dimethyl ether.
  • the reaction of ethylene with a thiolation reagent produces propylene.
  • This type of reaction provides a new way to produce propylene.
  • the advantage of this approach is that one carbon atom that produces propylene is derived from relatively inexpensive methanol or / and dimethyl ether, reducing the cost of propylene production. If a low-value ethylene feedstock such as catalytic cracking dry gas is used, the economics of the process can be further improved.
  • U.S. Patent 3,906,054 discloses a process for the oximation of olefins by contacting an olefin with a catalyst in the presence of an alkylating agent having a silica to alumina ratio of at least 12 and having a P content of at least 0.78%.
  • the olefins which can be thiolated include ethylene, propylene, butene-2 and isobutylene, and useful thiolation reagents are methanol, dimethyl ether and methyl chloride.
  • World Patent WO 2005/056504 A1 discloses a process for the efficient production of propylene starting from ethylene and methanol or/and dimethyl ether by reacting ethylene with methanol or/and dimethyl ether in the presence of a catalyst to form propylene. It is characterized in that the amount of ethylene flowing out of the reaction system is less than the amount of ethylene added to the reaction system. At the same time, the propylene yield can be up to 40 mol% or more based on the number of moles of methanol entering the reaction system or twice the number of moles of dimethyl ether.
  • Chinese Patent Application No. 200610112555.0 discloses a process for preparing propylene, which is characterized in that: a raw material containing ethylene is in the presence of a methylating agent under a specific reaction condition and a molecular sieve having a micropore diameter of 0.3-0.5 nm. The catalyst contacts to form a product containing propylene. The propylene selectivity in the product can reach more than 70%.
  • the butene formed can be further reacted with a thiolation reagent to form a C 5 or higher hydrocarbon; the formation of these high carbon hydrocarbons reduces the economics of the process:
  • the conversion of the reagent itself produces ethylene on the one hand, offsets the ethylene feedstock consumed in the alkylation reaction, and on the other hand produces propylene at a lower selectivity, reducing the propylene selectivity throughout the process.
  • a higher ratio of raw material ethylene/alkylation reagents is used to suppress these side reactions and achieve higher propylene selectivity, which requires a large amount of unconverted ethylene to be recycled repeatedly, which greatly increases the energy consumption of the process. Summary of the invention
  • the present invention provides a method for producing propylene, wherein a pore diameter of 0.3 nm to 0.5 nm is used, and an ammonia saturation adsorption amount of 0.8 mmol/g to 2.0 mmol/g at 200 ° C is used.
  • Catalyst, and at least two reaction zones are provided, including - a) in the first reaction zone, a hydrocarbon having a carbon number of not less than 4 undergoes a catalytic cracking reaction on the catalyst, and is converted into a hydrocarbon product including ethylene and propylene;
  • At least a part of ethylene is used as a raw material of the second reaction zone, and at least a part of hydrocarbons having a carbon number of not less than 4 in the product of the second reaction zone is used as a raw material of the first reaction zone.
  • the reaction condition of the first reaction zone is: the reaction temperature is 350-75 (TC, the reaction pressure is 0.01-0.8 MPa;
  • the reaction conditions of the second reaction zone are: a reaction temperature of 300-600 ° C, a reaction pressure of 0.01-0.8 MPa, and a molar ratio of ethylene/methanol or ethylene/2-fold dimethyl ether of 0.05-5.
  • the catalyst comprises at least one silica-alumina molecular sieve or a silicoaluminophosphate molecular sieve, or a product obtained by modifying a molecular sieve having the above characteristics by elements other than a skeleton constituent element, or a plurality of products satisfying the above characteristics. a mixture of molecular sieves.
  • the method wherein the catalyst has a molecular sieve content of from 10% by weight to 90% by weight.
  • the catalyst is formed by one or more of binders including silica, alumina or clay.
  • the method wherein the reactor forms of the first reaction zone and the second reaction zone each employ a fluidized bed.
  • hydrocarbon having a carbon number of not less than 4 in the first reaction zone is liquefied gas, naphtha, gasoline, condensate, light diesel oil, hydrogenated tail oil or kerosene, or in the conversion process of claim 1.
  • the ethylene-containing gas in the second reaction zone is ethylene derived from a process of hydrocarbon cracking, acetonitrile dehydrogenation or methanol conversion to olefins, or ethylene and C1-C3 hydrocarbons or carbon from the above process.
  • reaction condition of the first reaction zone is: the reaction temperature is 400-700 ° C, and the reaction pressure is
  • reaction conditions of the second reaction zone are: a reaction temperature of 350-550 ° C, a reaction pressure of 0.1-0.45 MPa, and an ethylene/methanol (or 2 times dimethyl ether) molar ratio of 0.1-5 .
  • the propylene selectivity in the product can be up to 75% or more.
  • two different conversion processes using propylene as a target product namely, high-carbon hydrocarbon catalytic cracking to produce propylene and ethylene and methanol (or / and dimethyl ether) co-feed to produce propylene, combined, each will At least a portion of the by-products obtained by the process are used as starting materials for another process, making full use of the by-products obtained in the two processes, and finally obtaining the propylene product with high selectivity.
  • Another advantage of this method is that the two processes can share some equipment.
  • the two reactors use the same catalyst and use a fluidized bed reactor.
  • the catalysts that deactivate different reaction zones can be carried out in the same regenerator. Regeneration; at the same time, product separation can be carried out in the same separation system, Significant savings in investment, lower energy consumption, and improved economics throughout the process.
  • the high carbon number cracking of propylene to produce propylene, and the co-feeding of ethylene with methanol or / and dimethyl ether to produce propylene are two different types of reactions which have different requirements for the performance of the catalyst.
  • the catalyst is first required to have a suitable pore size distribution.
  • the selectivity of the catalytic reaction often depends on the corresponding size of the molecule and the pore size. This selectivity is called shape selective catalysis.
  • shape selective catalysis In the process of high carbon number hydrocarbon cracking, by selecting a molecular sieve catalyst with a certain pore size, larger molecules in the product mixture, such as hydrocarbons above c 4 , are difficult to diffuse out of the pores of the molecular sieve catalyst, thereby improving hydrocarbon catalysis. Selectivity of ethylene and propylene in the cracking reaction.
  • the catalyst is specifically acidic.
  • the acidity of the catalyst is required to ensure sufficient reactivity.
  • this conversion process is usually accompanied by side reactions such as hydrogen transfer and cyclization. These side reactions not only produce coking of polynuclear aromatic hydrocarbons, etc.
  • a suitable acid distribution is beneficial to reduce such side reactions, inhibit the formation of coke and saturated hydrocarbons, slow down catalyst deactivation and improve the selectivity of the desired product.
  • a suitable acid distribution can reduce the direct conversion of methanol or / and dimethyl ether, which is beneficial between ethylene and methanol (or / and dimethyl ether)
  • the thiolation reaction increases the feedstock utilization rate and propylene selectivity in the product.
  • the principle is: The conversion of methanol/dimethyl ether to an olefin over an acidic catalyst occurs through a "carbon pool mechanism.”
  • the pores or cages of the catalyst are highly active polysubstituted aromatic hydrocarbons (ie, "carbon pools”). These polysubstituted aromatic hydrocarbons are rapidly methylated with methanol or/and dimethyl ether, and then further dealkylated, released. Ethylene or propylene molecules.
  • the rate and number of carbon pool formation on the catalyst determines the rate of direct conversion of methanol or / and dimethyl ether.
  • the formation of carbon pool involves hydrogen transfer, cyclization and other reactions need to occur at multiple acid centers adjacent to each other. We can reduce the formation of carbon pool by reducing the acid center density of the catalyst and increasing the distance between acid centers. Thereby direct conversion of methanol or/and dimethyl ether is inhibited.
  • the number of acid centers of the catalyst is within a certain range, can the requirements of the above two conversion processes be satisfied at the same time.
  • the amount of basic molecular adsorption under certain conditions is the characterization of molecular sieve acid.
  • the number of acid centers of the catalyst is represented by the amount of ammonia saturated adsorption per unit weight of molecular sieve at 200 °C.
  • a catalyst having a pore diameter of 0.3 nm to 0.5 nm and an ammonia saturated adsorption amount of 0.8 mmol/g to 2.0 mmol/g at 200 ° C is used, and at least two reaction zones are provided, including:
  • a hydrocarbon having a carbon number of not less than 4 undergoes a catalytic cracking reaction on the catalyst and is converted into a hydrocarbon product including ethylene and propylene;
  • At least a part of ethylene is used as a raw material of the second reaction zone, and at least a part of the hydrocarbons having a carbon number of not less than 4 in the product of the second reaction zone is used as a raw material of the first reaction zone.
  • the catalyst may contain at least one silica-alumina molecular sieve having a pore diameter of 0.3-0.5 nm and an ammonia-saturation adsorption amount of 0.8 mmol/g to 2.0 mmol/g at 200 ° C or A silica-phosphorus aluminum molecular sieve, or a product which is modified by other elements in accordance with the above-described characteristic molecular sieve, or a mixture of a plurality of molecular sieves conforming to the above characteristics.
  • the catalyst may have a molecular sieve content of from 10% by weight to 90% by weight.
  • the catalyst may be formed by one or more of binders including silica, alumina or clay.
  • both the first reaction zone and the second reaction zone may be in the form of a fluidized bed.
  • the hydrocarbon having a carbon number of not less than 4 in the first reaction zone may be liquefied gas, naphtha, gasoline, condensate, light diesel oil, hydrogenated tail oil or kerosene, or the conversion process of the present invention.
  • the gas containing ethylene in the second reaction zone may be ethylene produced by a process such as hydrocarbon cracking, acetonitrile dehydrogenation or methanol conversion to olefins, or ethylene and C1-C3 hydrocarbons from the above process. Or a mixture of carbon oxides, or a product containing ethylene during the conversion of the present invention.
  • the hydrocarbons of the first reaction zone having a carbon number of not less than 4 or/and the ethylene of the second reaction zone are respectively derived from the products of the conversion process of the present invention.
  • the reaction condition of the first reaction zone may be: the reaction temperature is 350-750 ° C, preferably 400-700 ° C, and the reaction pressure is 0.01-0.8 MPa, preferably 0.1. -0.45 MPa.
  • the reaction conditions of the second reaction zone may be: a reaction temperature of 300-600 ° C, preferably 350-550 ° C, a reaction pressure of 0.01-0.8 MPa, preferably 0.1-0.45 MPa. , ethylene / methanol (or 2 times dimethyl ether) molar ratio of 0.05-10, preferably 0.1-5;
  • the total selectivity of the product propylene can reach more than 75%.
  • Catalyst A used SAPO-34 molecular sieve (Dalian Institute of Chemical Physics, Chinese Academy of Sciences, microporous pore size about 0.4nm, ammonia saturated adsorption capacity of 1.36mmol / gram at 200'C) and clay, aluminum sol and silica sol (both purchased from Zhejiang Yuda Chemical Co., Ltd.) mixes and disperses into a slurry in water, and is spray-molded into microspheres with a particle size distribution of 20-100 microns. The above microspheres were calcined at 600 ° C for 4 hours to be the catalyst A. The SAPO-34 content in the catalyst was 30% by weight.
  • the ammonia saturation adsorption measurement procedure of the above SAPO-34 molecular sieve at 200 Torr is as follows:
  • the instruments used are Microchem's Autochem 2910 chemisorption analyzer and the Swiss PFeiffer Omnistar 300 online mass spectrometer.
  • the amount of ammonia removed by the integration is the ammonia saturated adsorption amount of the molecular sieve at 200 ⁇ .
  • the butene catalytic cracking reaction is carried out in a micro fluidized bed reactor.
  • the reaction conditions are as follows: The catalyst loading is 10g, the reaction temperature is 450 °C, and the raw material is Futeng Petrochemical Company butene-2 (purity 98%, cis, anti-butene-2 each 50% by weight), feed airspeed lO h 1 , the reaction pressure is 0.1 MPa, using water vapor as the reaction diluent gas, and the feed ratio of water to butene-2 is 1.5:1 (weight ratio).
  • the reaction product was analyzed by Varian CP-3800 gas chromatography, Plot column and hydrogen flame detector at a sampling time of 6 minutes.
  • the co-feed reaction of ethylene and methanol is carried out in a microfluidizer reactor.
  • the reaction conditions were as follows: The catalyst loading was 10 g, the reaction temperature was 400 ° C, and the raw materials were methanol (analytically pure, Shenyang Federal Reagent Factory) and ethylene (purity 99.5%, Ministry of Chemical Industry Bright Special Gas Research Institute) mixture.
  • the reaction pressure was 0.1 MPa.
  • the reaction product was analyzed by Varian CP-3800 gas chromatography, Plot column and hydrogen flame detector at a sampling time of 6 minutes. The results of the reaction are shown in Table 2. Under the above reaction conditions, the ethylene conversion was 19.27%, the methanol conversion was 100%, and the yield of propylene in the product was 62.61% (% by mole, based on methanol).
  • Catalyst B was made of SAPO-34 molecular sieve (Dalian Institute of Chemical Physics, Chinese Academy of Sciences, with a micropore diameter of about 0.4 nm and an ammonia adsorption capacity of 1.28 mmol/g at 200 ° C), using silica sol (purchased from Zhejiang Yuda Chemical Co., Ltd.). The company is formed as a binder and calcined at 550 ° C for 4 hours. The content of SAPO 34 in the catalyst after molding is 80.
  • the ammonia saturation adsorption amount measurement step of the SAPO-34 molecular sieve at 200 ° C is the same as in the first embodiment.
  • the butene catalytic cracking reaction is carried out in a fixed bed microreactor.
  • the reaction conditions are as follows: the catalyst loading is lg, the reaction temperature is 450 ° C, and the raw material is Futeng Petrochemical Company butene 2 (purity 98%, cis, anti-butene 2 each accounted for 50% by weight), feed airspeed l .O hr' 1 , the reaction pressure is O. lMPa, using water vapor as the reaction diluent gas, the feed ratio of water to butene-2 is 1.5:1 (weight ratio).
  • the reaction product was analyzed by Varian CP-3800 gas chromatography, Plot column and hydrogen flame detector at a sampling time of 6 minutes.
  • the ethylene and methanol co-feed reaction is carried out in a fixed bed microreactor.
  • the reaction conditions are as follows: The catalyst loading is lg, the reaction temperature is 400 ° C, and the raw materials are methanol (analytically pure, Shenyang Federal Reagent Factory) and ethylene (purity 99.5 %, Ministry of Chemical Industry Bright Special Gas Research Institute) mixture, mixed mode is ethylene
  • the methanol vapor is carried by a bubble saturator.
  • the reaction product was analyzed by Varian CP-3800 gas chromatography, Plot column and hydrogen flame detector at a sampling time of 6 minutes.
  • Catalyst C was prepared by using SAPO-34 molecular sieve (Dalian Institute of Chemical Physics, Chinese Academy of Sciences, microporous pore size of about 0.4 nm, ammonia adsorption capacity of 2.7 mmol/g at 200 ° C), using silica sol (purchased from Zhejiang Yuda Chemical Co., Ltd.). The company is formed as a binder and calcined at 550 Torr for 4 hours. After molding, the content of SAPO-34 in the catalyst is 80.
  • ammonia saturation adsorption amount measurement step, reaction conditions and analysis method of the SAPO-34 molecular sieve at 200 ° C are the same as in the second embodiment.
  • Catalyst D uses ZSM-5 molecular sieve (Fushun Petrochemical Company catalyst plant, micropore diameter 0.53nmX0.56nm, 20 (TC saturated ammonia adsorption capacity 1.40 mmol / gram), with clay, aluminum sol and silica sol (both purchased from Zhejiang Yuda Chemical Co., Ltd. mixes and disperses into a slurry in water, and is spray-molded into microspheres with a particle size distribution of 20-100 ⁇ m. The above microspheres are calcined by 60 (TC for 4 hours, which is the catalyst D.
  • the ZSM-5 content is 30 weight 0 / 0 .
  • ammonia saturation adsorption amount measurement step, reaction conditions and analysis method of the ZSM-5 molecular sieve at 200 ° C are the same as in the first embodiment.
  • a scheme for producing propylene from ethylene and methanol is designed in the form of a fluidized bed comprising two reaction zones and a common regenerator, and a catalyst A is used.
  • the contact time in each reaction zone was substantially the same as in Example 1, so that the raw material conversion rate and product selectivity were calculated in accordance with Example 1, and the coke yield was ignored.
  • a flow rate of 100 tons of I hour of mixed olefinic feedstock having a carbon number of not less than 4 (48 tons of I hour from the product of the second reaction zone, 52 tons of hourly use of unconverted feedstock by separation cycle)
  • the product having a carbon number of not less than 4 in the reaction zone is contacted with the catalyst.
  • the reaction temperature is 450 ⁇
  • the feed space velocity is 0.8-1.2 hr
  • the reaction pressure is 0.25 MPa.
  • the water vapor is used as the reaction diluent gas
  • the feed ratio of water to raw material is 1.5:1.
  • the material flowing out of the reaction zone is separated to obtain 32 tons of 1 hour of propylene, 5 tons of 1 hour of ethylene, 4 tons of 1 hour of C1 to C3 alkane, and 59 tons of hourly hydrocarbon having a carbon number of not less than 4 (including unconverted).
  • the raw material and the hydrocarbons produced in the reaction zone have a carbon number of not less than 4). All of the ethylene enters the second reaction zone, and 52 tons of hydrocarbons having a carbon number of not less than 4 are returned to the raw materials of the reaction zone.
  • the flow rate is 104 tons of ethylene per hour (80 tons of I hour is the unconverted raw material used in the separation cycle, 5 tons / hour from the first reaction zone, 19 tons / hour is the additional supplemental ethylene
  • the raw material is contacted with the catalyst together with 220 tons/hour of methanol.
  • the reaction temperature is 400'C, the feed space velocity in terms of methanol 0.8 -.
  • the overall material balance of the device is as follows: Inflow device 19 tons / hour of ethylene and 220 tons / hour of methanol, outflow device 92 tons / hour of propylene, 7 tons of I hour carbon number of not less than 4 hydrocarbons, 4 tons of I hour ethylene and 12 tons of I H1 C3 - C3 hydrocarbons.
  • the propylene yield in the whole process was 80% by carbon.
  • a scheme of producing propylene from a mixed olefin having dimethyl ether and a carbon number of not less than 4 is used in the same manner as in the embodiment 3, and the flow rates of the inflow and outflow materials in the respective reaction zones are different.
  • the mixed olefin having a carbon number of not less than 4 may be liquefied gas, naphtha, gasoline, condensate, light diesel oil, hydrogenated tail oil, kerosene or the like, and the flow rate thereof is based on the olefin contained in the raw material.
  • the conversion of dimethyl ether on the catalyst was the same as that of methanol except that the amount of water produced during the process was different.
  • the reaction conditions in each reaction zone were substantially the same as in Example 1, so that the conversion of the raw materials and the selectivity of the product were calculated in accordance with Example 1, and the coke yield was ignored.
  • a flow rate of 240 tons / hour of mixed olefinic feedstock with a carbon number of not less than 4 (24 tons of I hour from the product of the second reaction zone, 124.8 tons of I hour for the separation of recycled unconverted raw materials)
  • the product in the reaction zone having a carbon number of not less than 4, and 91.2 ton / hr as an additional supplementary material) is contacted with the catalyst.
  • the reaction temperature was 450 ° C
  • the feed space velocity was 0.8-1.2 hr
  • the reaction pressure was 0.25 MPa
  • water vapor was used as the reaction diluent gas.
  • the feed ratio of water to raw material was 1.5:1.
  • the material flowing out of the reaction zone is separated to obtain 76.8 tons/hour of propylene, 12 tons/hour of ethylene, 9.6 tons/hour of CI-C3 alkane, and 141.6 tons/hour of hydrocarbons having a carbon number of not less than 4 (including unconverted).
  • the raw material and the hydrocarbons produced in the reaction zone have a carbon number of not less than 4). All of the ethylene enters the second reaction zone, and 124.8 tons of hydrocarbons having an I-hour carbon number of not less than 4 are returned to the raw materials of the reaction zone.
  • a flow rate of 52 tons / hour of ethylene (40 tons / hour is the unconverted raw material used in the separation cycle, 12 tons / hour from the first reaction zone) and 79 tons / hour of dimethyl ether Contact with the catalyst.
  • the reaction temperature was 400 ° C
  • the feed space velocity was 0.6-0.9 hr -1 in terms of dimethyl ether
  • the reaction pressure was 0.25 MPa.
  • the material flowing out of the reaction zone is separated to obtain 30 tons of I-hour propylene, 42 tons of 1 hour of unconverted raw material ethylene, 4 tons/hour of C1-C3 alkane, and 24 tons/hour of hydrocarbons having a carbon number of not less than 4.
  • 40 tons/hour of ethylene are all returned to the raw materials in the reaction zone, and all of the hydrocarbons having a carbon number of not less than 4 enter the first reaction zone.
  • the overall material balance of the device is as follows: 91.2 tons/hour of hydrocarbons with a carbon number of not less than 4 and 79 tons of I-hour dimethyl ether, 106.8 tons of I-hour propylene, 16.8 tons of hydrocarbons with a carbon number of not less than 4 2 tons of 1 hour ethylene and 13.6 tons of 1 hour C1 to C3 hydrocarbon.
  • the propylene yield in the whole process was 76.7 carbon %.

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一种制取丙烯的方法 技术领域
本发明涉及一种制取丙烯的方法。 背景技术
丙烯是一种重要的石油化工基础原材料。 长期以来丙烯来源依赖于乙烯裂解装置 和 FCC装置， 由于丙烯增长率持续高于乙烯增长率， 人们不断改进传统的丙烯生产工 艺， 进一步提高丙烯产率。 而在现有装置上增产丙烯受到原料组成、 装置处理能力、 装置改造和操作费用的限制， 因此开发新的增产丙烯的工艺过程是满足日益增长的丙 烯需求的重要方向。 近年来， 人们开发了采用不同原料制取丙烯新工艺， 如丙垸脱氢 制丙烯、乙烯和丁烯反歧化制丙烯、高碳数烃类裂解制低碳烯烃、甲醇制烯烃（ΜΤΟ)、 甲醇制丙烯 （ΜΤΡ)、 乙烯与烷基化试剂共进料制丙烯等。
以高碳数烯烃、 特别是含有较多烯烃的高碳数烯烃物料 （如 FCC汽油和混合 C4 等） 为原料、 通过催化裂解转化为以丙烯为主的低碳烯烃， 是近年来受到广泛重视、 具有商业化潜力的新过程。这一转化过程的优点在于， 可以利用的原料种类来源丰富、 价值低， 在产物中丙烯 /乙烯比例高， 整个过程具有较好的经济性。 已有多项专利公 布了采用高碳数烯烃催化转化生产丙烯的方法。
U.S. Patent 6,222,087 B1公布了一种将含有 C4-C7烯烃和烷烃的原料转化为低碳 烯烃的方法， 催化剂为 P改性的 ZSM-5或 /和 ZSM-11分子筛， 分子筛的硅铝比大于 300。 采用密相流化床工艺， 反应条件为温度在 5I0-704°C , 反应压力在负压到 8bar之 间, WHSV在 UOh 1之间。 低碳烯烃收率高于 20%， 最高可达 30%以上， 丙烯 /乙烯 比可达 3.0以上
EP 0109059公布了一项将 C4-C12烯烃转化为丙烯的方法。 釆用的催化剂为硅铝 比小于或等于 300的 ZSM-5或，和 ZSM-11分子筛， WHSV大于 SOhr—1 , 反应温度为 400-600°C。 乙烯和丙烯的总收率为 36-44%, 其中丙烯的收率为 30-40%。
U.S. Patent 5,171,921和 EP0511013 A3公布了一种将高碳数混合烃（含有烯烃和垸 烃） 转化为低碳烯烃的技术， 反应温度为 500-700'C， WHSV在 ΙΟ-lOOOhr·1之间， 催 化剂中含有硅铝比为 20-60的 ZSM-5, 并经过 P改性和水蒸气老化处理。
U.S. Patent 5,981,819公布了一种将含有 C4-C7烯烃的物料转化为丙烯的技术。 反 应物料与水蒸汽混合进入固定床反应器与分子筛催化剂接触反应， 进料水 /油比为
0.5:1-3:1 ,反应温度为 380-500°C。分子筛的 Si/Al原子比为 10-200。原料中烯烃的 60% 以上转化为丙烯和丁烯。 WO 01/05909 A1公开了与上述过程相似的工艺，将含有 C4-C8 烯烃的物料转化为低碳烯烃。
US Patent 2003/0139636 A1公开了一种将含烯烃物料转化为丙烯的方法。 采用的 催化剂为稀土或金属改性的 SAPO、 MeAPO、 MeASPO、 ELAPO和 ELASPO。
CN 1600757公开了一种由含有 C4-C6烯烃的烃原料生产轻质烯烃特别是丙烯的 方法， 该方法将烯烃原料与改性的硅铝比大于 30的 ZSM-5/ZSM-11沸石催化剂接触， 以产生轻质烯烃流出物， 轻质烯烃的选择性在 60%以上， 收率为 40-55%, 反应条件为 温度 500-650Ό、 重量空速 1-50， 压力 0.1-8atm。
CN 1490288公开了一种 C4及以上烯烃催化裂解生产丙烯的方法， 主要解决以往 技术中存在的目的产物中丙烯的选择性低、 收率低、 催化剂稳定性差的问题。 采用的 催化剂为硅铝比为 50-1000 的 ZSM-5,并在分子筛的晶化过程中添加一定量的卤素钠 盐。 反应条件为温度 400-600°C、 液体空速 10-50 hr-l， 压力 0-0.15MPa。
另一种有良好应用前景的丙烯生产的新工艺是， 将乙烯与垸基化试剂 （如甲醇或 /和二甲醚等）共进料， 在催化剂上发生垸基化反应， 生成包括丙烯在内的烃类产品。
研究发现， 烯烃可与甲醇之间发生烷基化反应， 使得烯烃的碳数增加 （Svelle等， J. Catal. 224 (2004) , 115-123, J. Catal. 234 (2005 ) ，385— 400):
CH3OH + CnH2n = Cn+lH2n+2 +H20
以上类型的垸基化反应， 也可以在烯烃或 /和二甲醚等垸基化试剂之间发生。 特 别地， 乙烯与垸基化试剂的反应可生成丙烯。 这种类型的反应为丙烯的生产提供了一 个新的途径。 这一途径的优点在于： 生成丙烯的一个碳原子来自于相对便宜的甲醇或 /和二甲醚， 降低了丙烯生产的成本。 如果采用催化裂解干气等低价值乙烯原料， 则 该方法的经济性可进一步提高。
美国专利 US3906054公开了一种烯烃垸基化的工艺， 将烯烃在烷基化试剂存在下 与催化剂接触，催化剂为硅铝比至少为 12的沸石，采用 P改性， P含量最低为 0.78%。 可进行垸基化的烯烃包括乙烯、 丙烯、 丁烯一 2和异丁烯， 可用的垸基化试剂为甲醇、 二甲醚和氯甲烷。
世界专利 WO2005/056504 A1公开了一种从乙烯和甲醇或 /和二甲醚出发， 高效 制备丙烯的方法， 将乙烯和甲醇或 /和二甲醚在催化剂存在下进行反应而生成丙烯。 其特征在于， 由反应体系中流出的乙烯量少于向反应体系中加入的乙烯量。 同时， 以 进入反应体系的甲醇的摩尔数或 2倍的二甲醚摩尔数计算， 丙烯收率可达 40mol %以 上。
中国专利申请 200610112555.0公开了一种制取丙烯的方法， 该方法的特征在于： 含有乙烯的原料在甲基化试剂存在下， 在特定的反应条件下与含有微孔孔径为 0.3-0.5nm的分子筛的催化剂接触， 生成含有丙烯的产物。产物中丙烯选择性可达 70% 以上。
上述两种方法， 在生成丙烯的同时也生成其他副产物。 高碳数烃类裂解制低碳烯 烃的过程生成乙烯副产物， 乙烯的分离和提纯只能依靠深冷装置， 涉及到巨大的建设 投资， 从而大大降低了整个过程的经济性。 乙烯与甲醇或 /和二甲醚等烷基化试剂可 在酸性催化剂表面发生烷基化反应产生丙烯， 但是， 在同样的催化剂上也可以发生其 他多种反应， 如， 产物丙烯也可以与垸基化试剂反应而生成丁烯， 同样， 生成的丁烯 又可以进一步与垸基化试剂反应而生成 C5以上烃类； 这些高碳数烃类的生成， 降低了 过程的经济性： 垸基化试剂自身转化一方面产生乙烯， 抵消了在烷基化反应中消耗的 乙烯原料， 另一方面以较低选择性生成丙烯， 降低了全过程的丙烯选择性。 通常， 采 用较高的原料乙烯 /烷基化试剂比例， 才能抑制这些副反应、实现较高的丙烯选择性， 这样需要大量未转化的乙烯反复循环反应， 大大增加了过程的能耗。 发明内容
本发明的目的在于提供一种制取丙烯的方法。
为实现上述目的， 本发明提供的制取丙烯的方法， 其中， 采用微孔孔径为 0.3nm-0.5nm、 且在 200'C下氨饱和吸附量为 0.8毫摩尔 /克一 2.0毫摩尔 /克的催化剂， 并设置至少两个反应区， 包括- a)在第一个反应区， 碳数不小于 4的烃类在催化剂上发生催化裂解反应， 转化为 包括乙烯和丙烯的烃类产物；
b)在第二个反应区， 甲醇（或 /和二甲醚）和含有乙烯的气体在与 a)相同的催化 剂上反应， 转化为含有丙烯和更高碳数烃类的产物；
c)第一反应区的产物中， 至少一部分乙烯用作第二反应区的原料， 第二反应区的 产物中， 至少一部分碳数不小于 4的烃类用作第一反应区的原料。
其中第一反应区的反应条件为： 反应温度为 350-75(TC， 反应压力为 0.01-0.8MPa; 其中第二反应区的反应条件为： 反应温度为 300-600°C， 反应压力为 0.01-0.8MPa, 乙烯 /甲醇、 或乙烯 /2倍的二甲醚摩尔比为 0.05-5。
所述的方法， 其中， 所述的催化剂含有至少一种硅铝分子筛或硅磷铝分子筛， 或 符合上述特征的分子筛经骨架组成元素以外的元素改性得到的产物， 或多种符合上述 特征的分子筛的混合物。
所述的方法， 其中催化剂的分子筛含量为 10重量％—90重量％。
所述的方法， 其中的催化剂采用包括氧化硅、 氧化铝或粘土等粘合剂中的一种或 几种粘结成型。
所述的方法， 其中第一反应区和第二反应区的反应器形式均采用流化床。
所述的方法， 其中第一反应区碳数不小于 4的烃类为液化气、 石脑油、 汽油、 凝 析油、 轻柴油、 加氢尾油或煤油， 或权利要求 1的转化过程中碳数不小于 4的烃类产 物。
所述的方法， 其中第二反应区中含有乙烯的气体为来自烃类裂解、 乙垸脱氢或甲 醇转化制烯烃等过程产生的乙烯、 或来自上述过程的乙烯与 C1-C3烃类或碳氧化物的 混合物、 或权利要求 1的转化过程中含有乙烯的产物。
所述的方法， 其中第一反应区碳数不小于 4的烃类或 /和第二反应区的乙烯分别 或全部来自权利要求 1的转化过程中的产物。
所述的方法， 其中第一反应区的反应条件为： 反应温度为 400-700°C , 反应压力为
Figure imgf000006_0001
所述的方法， 其中第二反应区的反应条件为： 反应温度为 350-550'C， 反应压力为 0.1-0.45MPa, 乙烯 /甲醇 (或 2倍的二甲醚) 摩尔比为 0.1-5。
采用本发明所述的方法， 其产物中丙烯选择性可达 75%以上。 具体实施方式
依照本发明， 将以丙烯为目的产物的两个不同转化过程， 即高碳数烃类催化裂解 制丙烯及乙烯与甲醇 （或 /和二甲醚） 共进料制丙烯， 结合起来， 将每个过程得到的 副产物中至少一部分用作另一过程的原料， 充分利用这两个过程中得到的副产物， 最 终高选择性地得到丙烯产物。 这一方法的另一个优势在于， 两个过程可以共用某些设 备， 特别是， 两个反应区使用同一催化剂并采用流化床反应器， 可以将不同反应区失 活的催化剂在同一再生器进行再生； 同时产物分离可以在同一套分离系统中进行， 可 大大节省投资、 降低能耗、 提高整个过程的经济性。
高碳数烃类裂解制丙烯、 和乙烯与甲醇或 /和二甲醚共进料制丙烯是两种不同类 型的反应， 它们对催化剂的性能有不同的要求。 为了在同一催化剂上实现两种不同类 型的催化转化， 需要将催化剂的孔径分布和酸性控制在特定范围， 使之能够同时满足 上述两种转化过程的需要。
同时满足上述两种转化过程的需要， 首先要求催化剂具备合适的孔径分布。 研究 表明， 当反应物和产物的分子线度与晶内的孔径相接近时， 催化反应的选择性常取决 于分子与孔径的相应大小。这种选择性称之为择形催化。在高碳数烃类裂解的过程中， 通过选择一定孔道尺寸的分子筛催化剂，使产物混合物中较大的分子，如 c4以上烃类， 难于从分子筛催化剂的孔道扩散出来， 从而提高烃类催化裂解反应中乙烯和丙烯的选 择性。 基于同样的原理， 在乙烯和甲醇或 /和二甲醚共进料的反应中， 通过选择一定 孔道尺寸的分子筛催化剂， 可以减少产物丙烯进一步垸基化而生成更高碳数烃类， 提 高丙烯的选择性。
在同一催化剂上同时实现上述两种转化过程的另一关键之处是， 催化剂要有特定 的酸性。在烃类催化裂解反应中， 一方面要求催化剂的酸性能够保证足够的反应活性， 另一方面， 这一转化过程通常伴有氢转移、 环化等副反应， 这些副反应不但产生多核 芳烃等结焦物种， 使催化剂由于酸中心覆盖或发生堵孔而失活， 另一方面生成甲烷、 乙垸和丙垸等饱和烃类， 降低了目标产物的选择性。 选择合适的酸性分布， 有利于减 少这类副反应、 抑制焦炭和饱和烃类的生成、 可以减缓催化剂失活并提高目的产物选 择性。 在甲醇或 /和二甲醚与乙烯共进料制丙烯过程中， 合适的酸性分布可以减少甲 醇或 /和二甲醚的直接转化， 有利于乙烯与甲醇（或 /和二甲醚）之间的垸基化反应， 提高原料利用率和产物中丙烯选择性。 其原理在于： 甲醇 /二甲醚在酸性催化剂上转 化为烯烃的过程是通过"碳池机理"发生的。 催化剂的孔道或笼中先生成高活性的多取 代芳烃（即"碳池 ")， 这些多取代芳烃快速地与甲醇或 /和二甲醚发生甲基化反应， 然 后进一步脱烷基， 释放出乙烯或丙烯分子。 催化剂上碳池的生成速率与数目决定了甲 醇或 /和二甲醚的直接转化速率。 碳池的生成涉及到氢转移、 环化等反应需要在位置 相邻的多个酸中心上发生， 我们通过降低催化剂的酸中心密度， 增加酸中心之间的距 离， 可以减少碳池的生成， 从而抑制甲醇或 /和二甲醚的直接转化。
研究发现， 只有当催化剂的酸中心数目在一定范围内时， 才能同时满足上述两种 转化过程的需要。 在分子筛研究领域， 一定条件下碱性分子吸附量， 是表征分子筛酸 中心数目的有效指标。 本发明中， 催化剂的酸中心数目由单位重量分子筛在 200'C的 氨饱和吸附量表示。
依据本发明， 采用微孔孔径为 0.3nm-0.5nm、 且在 200'C下氨饱和吸附量为 0.8毫 摩尔 /克一 2.0毫摩尔 /克的催化剂， 并设置至少两个反应区， 包括：
a) 在第一个反应区， 碳数不小于 4的烃类在催化剂上发生催化裂解反应， 转化为 包括乙烯和丙烯的烃类产物；
b)在第二个反应区， 甲醇（或 /和二甲醚）和含有乙烯的气体在与 a)相同的催化 剂上反应， 转化为含有丙烯和更高碳数烃类的产物；
c) 第一反应区的产物中， 至少一部分乙烯用作第二反应区的原料， 第二反应区的 产物中， 至少一部分碳数不小于 4的烃类用作第一反应区的原料。
在所述的方法中， 上述催化剂可含有至少一种微孔孔径为 0.3-0.5nm、 且在 200°C 下氨饱和吸附量为 0.8毫摩尔 /克一 2.0毫摩尔 /克的硅铝分子筛或硅磷铝分子筛， 或符 合上述特征分子筛经其它元素改性的产物， 或多种符合上述特征的分子筛的混合物。
在所述的方法中， 催化剂的分子筛含量可为 10重量％—90重量％。
在所述的方法中， 催化剂可采用包括氧化硅、 氧化铝或粘土等粘合剂中的一种或 几种粘结成型。
'在所述的方法中， 第一反应区和第二反应区的反应器形式均可以采用流化床。 在所述的方法中， 第一反应区碳数不小于 4的烃类可以为液化气、 石脑油、 汽油、 凝析油、 轻柴油、 加氢尾油或煤油， 或本发明的转化过程中碳数不小于 4的烃类产物。
在所述的方法中， 第二反应区中含有乙烯的气体可以为来自烃类裂解、 乙垸脱氢 或甲醇转化制烯烃等过程产生的乙烯、 或来自上述过程的乙烯与 C1-C3烃类或碳氧化 物的混合物、 或本发明的转化过程中含有乙烯的产物。
在所述的方法中， 第一反应区碳数不小于 4的烃类或 /和第二反应区的乙烯分别 或全部来自本发明的转化过程中的产物。
在所述的方法中， 第一反应区的反应条件可以为： 反应温度为反应温度为 350-750 °C , 最好为 400-700°C， 反应压力为 0.01-0.8MPa， 最好为 0.1-0.45MPa。
在所述的方法中， 第二反应区的反应条件可以为： 反应温度为 300-600°C , 最好为 350-550°C , 反应压力为 0.01-0.8MPa, 最好为 0.1-0.45MPa， 乙烯 /甲醇（或 2倍的二甲 醚） 摩尔比为 0.05-10， 最好为 0.1-5;
采用所述的方法， 产物丙烯总选择性可达 75%以上。 以下通过实施例对本发明作出详细描述， 但本发明并不局限于这些实施例。
实施例 1
催化剂 A采用 SAPO— 34分子筛（中国科学院大连化学物理研究所， 微孔孔径约 0.4nm, 200'C下氨饱和吸附量为 1.36毫摩尔 /克）与粘土、 铝溶胶和硅溶胶（均购自浙 江宇达化工有限公司） 混合并在水中分散成浆料， 喷雾成型后为粒径分布为 20— 100 微米的微球。 上述微球经 600°C焙烧 4小时， 即为催化剂 A。 催化剂中 SAPO— 34含 量为 30重量％。
上述 SAPO-34分子筛在 200Ό下氨饱和吸附量测量步骤如下： 使用的仪器为美国 Micrometric 公司的 Autochem2910 化学吸附分析仪和瑞士 PFeiffer 公司的 Omnistar 300在线质谱仪。 催化剂 0.2 g， 在 600'C下 40 ml/min 的 He气氛下活化 30 min， 然后降 温至 200Ό吸附氨气至饱和， 吹扫 30 min， 然后以 10°C/min 的速率升温脱附至 600 'C， TCD和质谱同时检测升温过程中催化剂释放的氨气， 经积分得到的脱除氨气量即为该 分子筛在 200Ό下氨饱和吸附量。
丁烯催化裂解反应在微型流化床反应装置内进行。 反应条件如下： 催化剂装填量 为 10g, 反应温度为 450°C， 原料采用抚顺石化公司丁烯一 2 (纯度 98% , 顺、 反丁 烯— 2各占 50重量％), 进料空速 l.O h 1 , 反应压力为 O.lMPa, 采用水蒸气为反应稀 释气， 水与丁烯— 2的进料比例为 1.5: 1 (重量比）。 反应产物采用 Varian CP-3800气 相色谱、 Plot柱和氢焰检测器分析， 取样时间点为 6分钟。
反应结果如表 1所示， 在上述反应条件下， 丁烯转化率为 52.63%， 产物中丙烯的 选择性为 60.92重量％。
表 1 : 实施例 1中的丁烯催化裂解反应结果
选择性 （重量％) CH4 C2H4 C2H6 C3H6 C3H8 C5 C6 +
0.13 8.98 0.08 60.92 5.15 22.62 2.13 丁烯转化率 （％) 52.63
乙烯与甲醇共进料反应在微型流化床反应装置内进行。 反应条件如下： 催化剂装 填量为 10g， 反应温度为 400'C， 原料采用甲醇（分析纯，沈阳联邦试剂厂）和乙烯（纯 度 99.5% , 化工部光明特种气体研究所）混合物。 原料组成为乙烯： 甲醇 =0.52: 0.48 (碳数比）， 进料空速以甲醇计为 1.0 ΙΙΓΛ 反应压力为 0.1MPa。 反应产物采用 Varian CP-3800气相色谱、 Plot柱和氢焰检测器分析， 取样时间点为 6分钟。 反应结果如表 2所示， 在上述反应条件下， 乙烯转化率为 19.27% , 甲醇转化率为 100% , 产物中丙烯的收率为 62.61 % (〇数％, 以甲醇计）。
表 2: 实施例 I中乙烯与甲醇共进料的反应结果
收率 （。数％， 以甲醇 CH4 C2H6 C3H6 C3Hg C4 C5 C6+
计）
0.56 1.08 62.61 6.98 36.90 9.28 3.58 乙烯转化率 （％) 19.27
甲醇转化率 （％) 100
实施例 2
催化剂 B由采用 SAPO— 34分子筛 （中国科学院大连化学物理研究所， 微孔孔径 约 0.4nm， 200°C下氨吸附量为 1.28毫摩尔 /克），采用硅溶胶 (购自浙江宇达化工有限公 司)作为粘结剂成型， 并经 550'C焙烧 4小时， 成型后催化剂中 SAPO— 34的含量为 80
%。
SAPO-34分子筛在 200'C下氨饱和吸附量测量步骤与实施例 1相同。
丁烯催化裂解反应在固定床微反装置内进行。反应条件如下：催化剂装填量为 lg， 反应温度为 450°C， 原料采用抚顺石化公司丁烯一 2 (纯度 98%， 顺、 反丁烯一 2各 占 50重量％)， 进料空速 l .O hr'1 , 反应压力为 O. lMPa, 采用水蒸气为反应稀释气， 水 与丁烯一 2的进料比例为 1.5: 1 (重量比）。 反应产物采用 Varian CP-3800气相色谱、 Plot柱和氢焰检测器分析， 取样时间点为 6分钟。
丁烯裂解反应结果如表 3所示， 在上述反应条件下， 丁烯转化率为 56.95 %， 产物 中丙烯的选择性为 57.02重量％。
表 3: 实施例 2中丁烯催化裂解的反应结果
选择性 （重量％) CH4 C2H4 C2H6 C3H6 C3H8 C C6 +
0.16 9.74 0.12 57.02 6.67 23.13 3.16 丁烯转化率 （％) 56.95
乙烯与甲醇共进料反应在固定床微反装置内进行。 反应条件如下： 催化剂装填量 为 lg， 反应温度为 400°C， 原料采用甲醇 （分析纯， 沈阳联邦试剂厂）和乙烯 （纯度 99.5 % , 化工部光明特种气体研究所） 混合物， 混合方式为乙烯通过鼓泡饱和器携带 甲醇蒸汽。原料组成为乙烯： 甲醇 =0.52: 0.48 (碳数％),进料空速以甲醇计为 1.0 hr—1 , 反应压力为 0.1MPa。 反应产物采用 Varian CP-3800气相色谱、 Plot柱和氢焰检测器分 析， 取样时间点为 6分钟。
反应结果如表 4所示， 在上述反应条件下， 乙烯转化率为 29.35 % , 甲醇转化率为 100% , 产物中丙烯的收率为 63.27% (。数％, 以甲醇计）。
表 4: 实施例 2中乙烯与甲醇共进料的反应结果
收率（。数％， 以甲醇 CH4 C2H6 C3H6 C3H8 C4 C5 C6+ 计）
0.59 1.20 63.27 9.83 40.26 11.68 5.13 乙烯转化率 （％) 29.35
甲醇转化率 （％) 100
对比例 1
催化剂 C由采用 SAPO— 34分子筛（中国科学院大连化学物理研究所， 微孔孔径 约 0.4nm， 200°C下氨吸附量为 2.7毫摩尔 /克）， 采用硅溶胶 (购自浙江宇达化工有限公 司)作为粘结剂成型， 并经 550Ό焙烧 4小时， 成型后催化剂中 SAPO— 34的含量为 80
%。
SAPO-34分子筛在 200'C下氨饱和吸附量测量步骤、 反应条件和分析方法与实施 例 2相同。
丁烯裂解反应结果如表 5所示。在上述反应条件下， 丁烯转化率为 61.02% , 产物 中丙烯的选择性为 42.56重量％。
表 5: 对比例 1中丁烯催化裂解的反应结果 选择性 （重量％) CH4 C2 C2¾ C3H C3Hg C5 C6 +
0.27 8.94 0.26 42.56 10.88 27.07 10.02 丁烯转化率 （％) 61.02
乙烯与甲醇共进料反应结果如表 6所示。在上述反应条件下， 乙烯转化率为 43.37 % , 甲醇转化率为 100% , 产物中丙烯的收率为 55.94% (。数％， 以甲醇计）。
表 6: 对比例 1乙烯与甲醇共进料的反应结果
收率（。数％， 以甲醇 CH4 C2H6 C3H6 C3H8 C4 C5 C6 + 计）
0.86 1.70 55.94 20.91 43.07 16.85 7.89 乙烯转化率 （％) 43.37
甲醇转化率 （％) 100
对比例 2
催化剂 D采用 ZSM-5分子筛 (抚顺石化公司催化剂厂，微孔孔径 0.53nmX0.56nm， 20(TC下氨饱和吸附量为 1.40毫摩尔 /克）， 与粘土、铝溶胶和硅溶胶（均购自浙江宇达 化工有限公司） 混合并在水中分散成浆料， 喷雾成型后为粒径分布为 20— 100微米的 微球。 上述微球经 60(TC焙烧 4小时， 即为催化剂 D。 催化剂中 ZSM-5含量为 30 重 量0 /0。
ZSM-5分子筛在 200°C下氨饱和吸附量测量步骤、反应条件和分析方法与实施例 1 相同。
丁烯裂解反应结果如表 7所示。在上述反应条件下， 丁烯转化率为 60.22% , 产物 中丙烯的选择性为 35.11重量％。
表 7: 对比例 2中丁烯催化裂解的反应结果
选择性 （重量％) CH4 C2H4 C2H6 C3H8 C5 C6 +
1.06 8.62 0.25 8.85 33.95 12.16 丁烯转化率 （％) 60.22
乙烯与甲醇共进料反应结果如表 8所示， 在上述反应条件下， 乙烯转化率为 81.6 % , 甲醇转化率为 100% , 产物中丙烯的收率为 95.64% (。数％， 以甲醇计）， 产物 中烷烃和 C4以上产物较高。
表 8: 对比例 2乙烯与甲醇共进料的反应结果
Figure imgf000013_0001
实施例 3
采用乙烯和甲醇为原料制丙烯的方案， 其装置形式设计为流化床， 包括两个反应 区和一个共用的再生器， 并采用催化剂 A。 在各个反应区内接触时间与实施例 1基本 一致， 因此其原料转化率和产物选择性依实施例 1计算， 并忽略焦炭产率。
在第一个反应区， 流量为 100吨 I小时的碳数不小于 4的混合烯烃原料 （其中 48 吨 I小时来自第二个反应区的产物， 52吨 I小时为经分离循环使用未转化原料和本反 应区碳数不小于 4的产物） 与催化剂接触。 反应温度为 450Ό , 进料空速 0.8-1.2 hr 反应压力为 0.25MPa， 采用水蒸气为反应稀释气， 水与原料的进料比例为 1.5: 1。 流 出该反应区的物料分离后 得到 32吨 I小时的丙烯、 5吨 I小时的乙烯、 4吨 I小时的 C1一 C3烷烃， 以及 59吨 /小时碳数不小于 4的烃类 （包括未转化原料和本反应区生 成的碳数不小于 4的烃类）。其中乙烯全部进入第二反应区， 52吨 /小时碳数不小于 4 的烃类返回到本反应区的原料中。
在第二个反应区， 流量为 104吨 I小时的乙烯（其中 80吨 I小时为经分离循环使 用的未转化原料、 5吨 /小时来自第一反应区、 19吨 /小时为额外补充的乙烯原料） 与 220吨 /小时甲醇共同与催化剂接触。 反应温度为 400'C， 进料空速以甲醇计为 0.8 - 1.2 hr"1 , 反应压力为 0.25 MPa。 流出该反应区的物料分离后 得到 60吨 /小时的丙 烯、 84吨 /小时的未转化的原料乙烯、 8吨 /小时的 CI— C3烷烃， 以及 48吨 /小时 碳数不小于 4的烃类。其中 80吨 /小时乙烯全部返回到本反应区的原料中， 碳数不小 于 4的烃类全部进入第一反应区。
装置整体物料平衡如下： 流入装置 19吨 /小时乙烯和 220吨 /小时甲醇， 流出装 置 92吨 /小时丙烯、 7吨 I小时碳数不小于 4的烃类、 4吨 I小时乙烯和 12吨 I小时 C1一 C3垸烃。 全过程丙烯收率为 80碳数％。 实施例 4
采用二甲醚和碳数不小于 4的混合烯烃为原料制丙烯的方案， 其装置形式与使用 的催化剂与实施例 3相同， 而各反应区流入和流出物料的流量不同。 碳数不小于 4的 混合烯烃可以为液化气、 石脑油、 汽油、 凝析油、 轻柴油、 加氢尾油、 煤油等， 其流 量以原料中含有的烯烃为准。 同样条件下， 除过程中生成水量不同， 二甲醚在催化剂 上的转化过程与甲醇相同。 各个反应区内反应条件与实施例 1基本一致， 因此其原料 转化率和产物选择性依实施例 1计算， 并忽略焦炭产率。
在第一个反应区， 流量为 240吨 /小时的碳数不小于 4的混合烯烃原料 （其中 24 吨 I小时来自第二个反应区的产物， 124.8吨 I小时为经分离循环使用未转化原料和本 反应区碳数不小于 4的产物、 91.2吨 /小时为额外补充的原料） 与催化剂接触。 反应 温度为 450°C，进料空速 0.8-1.2 hr 反应压力为 0.25MPa，采用水蒸气为反应稀释气， 水与原料的进料比例为 1.5 : 1。 流出该反应区的物料分离后 得到 76.8吨 /小时的丙 烯、 12吨 /小时的乙烯、 9.6吨 /小时的 CI— C3烷烃， 以及 141.6吨 /小时碳数不小 于 4的烃类（包括未转化原料和本反应区生成的碳数不小于 4的烃类）。其中乙烯全部 进入第二反应区， 124.8吨 I小时碳数不小于 4的烃类返回到本反应区的原料中。
在第二个反应区， 流量为 52吨 /小时的乙烯 （其中 40吨 /小时为经分离循环使 用的未转化原料、 12吨 /小时来自第一反应区） 与 79吨 /小时二甲醚共同与催化剂 接触。反应温度为 400°C ,进料空速以二甲醚计为 0.6— 0.9 hr—1 , 反应压力为 0.25 MPa。 流出该反应区的物料分离后 得到 30吨 I小时的丙烯、 42吨 I小时的未转化的原料乙 烯、 4吨 /小时的 C1—C3烷烃， 以及 24吨 /小时碳数不小于 4的烃类。其中 40吨 / 小时乙烯全部返回到本反应区的原料中， 碳数不小于 4的烃类全部进入第一反应区。
装置整体物料平衡如下： 流入装置 91.2吨 /小时碳数不小于 4的烃类和 79吨 I 小时二甲醚， 流出装置 106.8吨 I小时丙烯、 16.8吨 I小时碳数不小于 4的烃类、 2吨 I小时乙烯和 13.6吨 I小时 C1一 C3垸烃。 全过程丙烯收率为 76.7碳数％。

Claims

权 利 要 求
1、 一种制取丙烯的方法， 其中， 采用微孔孔径为 0.3nm-0.5nm、 且在 200'C下氨 饱和吸附量为 0.8毫摩尔 /克一 2.0毫摩尔 /克的催化剂， 并设置至少两个反应区， 包括: a)在第一个反应区， 碳数不小于 4的烃类在催化剂上发生催化裂解反应， 转化为 包括乙烯和丙烯的烃类产物；
b)在第二个反应区， 甲醇（或 /和二甲醚）和含有乙烯的气体在与 a)相同的催化 剂上反应， 转化为含有丙烯和更高碳数经类的产物；
c )第一反应区的产物中， 至少一部分乙烯用作第二反应区的原料， 第二反应区的 产物中， 至少一部分碳数不小于 4的烃类用作第一反应区的原料。
其中第一反应区的反应条件为： 反应温度为 350-750°C， 反应压力为 0.01-0.8MPa; 其中第二反应区的反应条件为： 反应温度为 300-600°C， 反应压力为 0.01-0.8MPa， 乙烯 /甲醇、 或乙烯 /2倍的二甲醚摩尔比为 0.05-5。
2、 权利要求 1所述的方法， 其中， 所述的催化剂含有至少一种硅铝分子筛或硅磷 铝分子筛， 或符合上述特征的分子筛经骨架组成元素以外的元素改性得到的产物， 或 多种符合上述特征的分子筛的混合物。
3、 权利要求 2所述的方法， 其中催化剂的分子筛含量为 10重量％—90重量％。
4、权利要求 1所述的方法， 其中的催化剂采用包括氧化硅、 氧化铝或粘土粘合剂 中的一种或几种粘结成型。
5、权利要求 1所述的方法， 其中第一反应区和第二反应区的反应器形式均采用流 化床。
6、 权利要求 1所述的方法， 其中第一反应区碳数不小于 4的烃类为液化气、 石脑 油、 汽油、 凝析油、 轻柴油、 加氢尾油或煤油， 或权利要求 1的转化过程中碳数不小 于 4的烃类产物。
7、 权利要求 1所述的方法， 其中第二反应区中含有乙烯的气体为来自烃类裂解、 乙垸脱氢或甲醇转化制烯烃等过程产生的乙烯、 或来自上述过程的乙烯与 C1-C3烃类 或碳氧化物的混合物、 或权利要求 1的转化过程中含有乙烯的产物。
8、权利要求 1所述的方法， 其中第一反应区碳数不小于 4的烃类或 I和第二反应 区的乙烯分别或全部来自权利要求 1的转化过程中的产物。
9、权利要求 1所述的方法，其中第一反应区的反应条件为：反应温度为 400-700°C， 反应压力为 0.1-0.45MPa。
10、 权利要求 1 所述的方法， 其中第二反应区的反应条件为： 反应温度为 350-550 , 反应压力为 0.1-0.45MPa， 乙烯 /甲醇 （或 2倍的二甲醚） 摩尔比 ¾ 0.1-5。
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