CN114763315B

CN114763315B - Catalytic conversion method for preparing low-carbon olefin

Info

Publication number: CN114763315B
Application number: CN202110032112.5A
Authority: CN
Inventors: 左严芬; 许友好; 舒兴田; 杜令印; 韩月阳; 郭秀坤; 谢昕宇
Original assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Current assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Priority date: 2021-01-11
Filing date: 2021-01-11
Publication date: 2024-05-17
Anticipated expiration: 2041-01-11
Also published as: CN114763315A

Abstract

The present disclosure relates to a catalytic conversion process for producing light olefins, the process comprising: s1, a hydrocarbon oil raw material with the olefin content of more than 50 weight percent is contacted with a catalytic conversion catalyst with the temperature of more than 650 ℃ at the upper part of a downlink reaction zone of a U-shaped tube reactor, and a first catalytic conversion reaction is carried out to obtain a first material flow, and the first material flow is separated to obtain a first catalyst and a first reactant flow; s2, contacting heavy raw oil with a first catalyst at the bottom of an uplink reaction zone of the U-shaped tube reactor and performing a second catalytic conversion reaction to obtain a second reactant flow and a spent catalyst; s3, separating ethylene, propylene, butylene, an olefin-rich stream and catalytic wax oil from the first reactant stream and the second reactant stream; the olefin-rich stream is introduced into the downstream reaction zone of the U-tube reactor for continued reaction. The method can effectively improve the yield and selectivity of ethylene and propylene, and can effectively inhibit the generation of byproducts such as methane and the like.

Description

Catalytic conversion method for preparing low-carbon olefin

Technical Field

The application relates to petroleum refining and petrochemical processing, in particular to a catalytic conversion method for preparing low-carbon olefin.

Background

Ethylene and propylene are two most important basic raw materials for forming modern petrochemical industry, wherein ethylene is one of the largest chemical products in the world and accounts for more than 75% of the global overall petrochemical product yield; propylene is also an important organic chemical raw material, and is mainly used for preparing acrylonitrile, propylene oxide, acetone and the like. Ethylene and propylene are used to produce a variety of important organic chemical raw materials, synthetic resins, synthetic rubbers, a variety of fine chemicals, and the like. However, with the continuous increase of oil field production, the available yield of conventional crude oil is gradually reduced, the quality of crude oil is gradually poorer, the crude oil gradually tends to be inferior and heavy, and the light olefin demand of the market cannot be met.

The traditional steam cracking process is adopted to prepare ethylene and propylene, the demand for light hydrocarbon, naphtha and other chemical light hydrocarbons is large, and the chemical light oil is difficult to meet the demand of ethylene and propylene raw materials due to the common weight of crude oil. The research institution predicts that the annual average growth rate of global gasoline complexes will be less than 1% from 2018 to 2026, but that propylene increases by about 4%. The high-carbon olefin in the refinery process is reasonably utilized to crack and prepare ethylene and propylene, thereby not only meeting the aim of improving quality and enhancing efficiency of petrochemical enterprises, but also conforming to the time demand of energy transformation.

Chinese patent CN101092323a discloses a method for preparing ethylene and propylene by using a C4-C8 olefin mixture as a raw material, reacting at 400-600 ℃ under the absolute pressure of 0.3-1.1KPa, and recycling 30-90 wt% of C4 fraction into the reactor through a separation device for re-cracking. The method mainly improves the conversion rate of olefin through C4 fraction circulation, and the obtained ethylene and propylene are not less than 62% of the total amount of olefin in the raw material, but the ethylene/propylene is smaller, can not be flexibly regulated according to market demands, and has the advantages of low reaction selectivity, high butene content in the product, C4 separation energy consumption and the like.

Chinese patent CN101239878a discloses an olefin-rich mixture of olefins with four or more carbon atoms as a raw material, the reaction is performed under the conditions of a reaction temperature of 400-680 ℃, a reaction pressure of-0.09-1.0 MPa and a weight space velocity of 0.1-50 hours-1, the ethylene/propylene of the product is lower and lower than 0.41, and the ethylene/propylene increases with the increase of the temperature, although the yield of diene can be improved to a certain extent, the selectivity of ethylene is lower, and mass production of ethylene cannot meet market demands.

Thus, there is a need in the art for a new process to produce ethylene and propylene in high yields.

Disclosure of Invention

The purpose of the present disclosure is to provide a catalytic conversion method for producing light olefins, which improves the yields of ethylene and propylene, and achieves the effective utilization of petroleum resources.

To achieve the above object, the present disclosure provides a catalytic conversion method for producing light olefins, the method comprising:

S1, a hydrocarbon oil raw material with the olefin content of more than 50 weight percent is contacted with a catalytic conversion catalyst with the temperature of more than 650 ℃ at the upper part of a downlink reaction zone of a U-shaped tube reactor, and a first catalytic conversion reaction is carried out to obtain a first material flow, and the first material flow is subjected to oil agent separation to obtain a first catalyst and a first reactant flow;

S2, contacting heavy raw oil with the first catalyst at the bottom of an uplink reaction zone of the U-shaped tube reactor and performing a second catalytic conversion reaction to obtain a second reactant flow and a spent catalyst;

S3, separating ethylene, propylene, butylene, an olefin-rich stream and catalytic wax oil from the first reactant stream and the second reactant stream; the olefin-rich stream is introduced into the downstream reaction zone of the U-tube reactor to continue the reaction.

Optionally, the olefin-rich stream is an olefin of C5 or greater; the olefin-rich stream has a content of olefins greater than C5 of from 50 wt% to 100 wt%.

Optionally, the method further comprises: carrying out burning regeneration on the spent catalyst to obtain a regenerated catalyst; and returning the regenerated catalyst to a downstream reaction zone of the U-shaped tube reactor after preheating.

Optionally, the method further comprises: and introducing the butene into a downlink reaction zone of the U-shaped pipe reactor for continuous reaction.

Optionally, the butene is contacted with the catalytic conversion catalyst prior to the olefin-rich stream.

Optionally, the reaction conditions for the continued reaction of butene include: the reaction temperature is 650-800 ℃, the reaction pressure is 0.05-1MPa, the reaction time is 0.01-10 seconds, and the weight ratio of the catalytic conversion catalyst to the butene is (20-200): 1, a step of; preferably, the reaction temperature is 680-780 ℃, the reaction pressure is 0.1-0.8MPa, the reaction time is 0.05-8 seconds, and the weight ratio of the catalytic conversion catalyst to the butene is (30-180): 1.

Optionally, the method further comprises: and (3) carrying out hydrotreating on the catalytic wax oil to obtain hydrogenation catalytic wax oil, and introducing the hydrogenation catalytic wax oil into the bottom of the uplink reaction zone of the U-shaped tube reactor to continue to react.

Optionally, the hydrotreating conditions include: the hydrogen partial pressure is 3.0-20.0 megapascals, the reaction temperature is 300-450 ℃, the hydrogen-oil volume ratio is 300-2000, and the volume airspeed is 0.1-3.0 hours ^-1.

Optionally, the conditions of the first catalytic conversion reaction include: the reaction temperature is 600-800 ℃, the reaction pressure is 0.05-1MPa, the reaction time is 0.01-100 seconds, and the weight ratio of the catalytic conversion catalyst to the hydrocarbon oil raw material is (1-200): 1, a step of; the conditions of the second catalytic conversion reaction include: the reaction temperature is 400-650 ℃, the reaction pressure is 0.05-1MPa, the reaction time is 0.01-100 seconds, and the weight ratio of the catalytic conversion catalyst to the heavy raw oil is (1-100): 1.

Preferably, the conditions of the first catalytic conversion reaction include: the reaction temperature is 630-780 ℃, the reaction pressure is 0.1-0.8MPa, the reaction time is 0.1-80 seconds, and the weight ratio of the catalytic conversion catalyst to the hydrocarbon oil raw material is (3-180): 1, a step of; the conditions of the second catalytic conversion reaction include: the reaction temperature is 450-600 ℃, the reaction pressure is 0.1-0.8MPa, the reaction time is 0.1-80 seconds, and the weight ratio of the catalytic conversion catalyst to the heavy raw oil is (3-70): 1.

Optionally, the olefin content in the hydrocarbon oil feedstock is 80wt% or more; preferably, the olefin content in the hydrocarbon oil feedstock is 90 wt% or more; more preferably, the hydrocarbon oil feedstock is a pure olefin feedstock.

Optionally, the heavy feedstock oil is selected from petroleum hydrocarbons and/or mineral oils; the petroleum hydrocarbon is at least one selected from vacuum gas oil, normal pressure gas oil, coker gas oil, deasphalted oil, vacuum residue, normal pressure residue and heavy aromatic raffinate oil; the mineral oil is selected from at least one of coal liquefied oil, oil sand oil and shale oil.

Optionally, the olefins in the hydrocarbon oil feedstock are derived from a C4 or more fraction produced by dehydrogenation of an alkane feedstock, a C4 or more fraction produced by a catalytic cracking unit in a refinery, a C4 or more fraction produced by a steam cracking unit in an ethylene plant, an olefin-rich fraction of C4 or more by-product MTO, and an olefin-rich fraction of C4 or more by-product MTP; the alkane feed is selected from at least one of naphtha, aromatic raffinate oil, and light hydrocarbons.

Alternatively, the catalytic conversion catalyst comprises 1 to 50 wt.% molecular sieve, 5 to 99 wt.% inorganic oxide, and 0 to 70 wt.% clay, based on the weight of the catalytic conversion catalyst; the molecular sieve comprises one or more of a large pore molecular sieve, a medium pore molecular sieve and a small pore molecular sieve. The catalytic conversion catalyst further comprises 0.1 to 3 wt% of an active metal, based on the weight of the catalytic conversion catalyst; the active metal is selected from one or more of VIII group metal, IVA group metal and rare earth metal.

Through the technical scheme, the macromolecule olefin produced in the oil-gas separation process is recycled in a specific route, so that the scheme of producing ethylene and propylene by steam cracking with high energy consumption in the prior art can be replaced, the yield and selectivity of ethylene and propylene are effectively improved, and meanwhile, the generation of byproducts such as methane can be effectively inhibited, and the purpose of efficiently utilizing petroleum resources is achieved.

Additional features and advantages of the present disclosure will be set forth in the detailed description which follows.

Drawings

The accompanying drawings are included to provide a further understanding of the disclosure, and are incorporated in and constitute a part of this specification, illustrate the disclosure and together with the description serve to explain, but do not limit the disclosure. In the drawings:

FIG. 1 is a schematic flow diagram of a first embodiment of the present disclosure;

fig. 2 is a flow diagram of a second embodiment of the present disclosure.

Description of the reference numerals

I downlink reaction zone II uplink reaction zone

1U pipe reactor 2 pipeline 3 cyclone separator

4 Line 5 line 6 line

7 Outlet section 8 settler 9 plenum

10 Pipeline 11 pipeline 12 inclined pipe

13 Regenerator 14 line 15 line

16 Pipeline 17 pipeline 18 product separator

19 Line 20 line 21 line

22 Line 23 line 24 line

25 Line 26 olefin separation plant 27 line

28 Pipeline 29 hydrotreating reactor

Detailed Description

The following describes specific embodiments of the present disclosure in detail. It should be understood that the detailed description and specific examples, while indicating and illustrating the disclosure, are not intended to limit the disclosure.

The present disclosure provides a catalytic conversion process for producing light olefins, the process comprising:

S3, separating ethylene, propylene, butene, an olefin-rich stream and catalytic wax oil from the first reactant stream and the second reactant stream; the olefin-rich stream is introduced into the downstream reaction zone of the U-tube reactor to continue the reaction.

The inventors of the present disclosure have found through a number of experiments that there is a difference in the distribution of products formed by the reaction of alkanes and alkenes over the catalyst. When the olefin is catalyzed on the high-temperature catalyst, the yields of hydrogen, methane and ethane in the product are lower, the yields of ethylene and propylene are higher, and the selectivity of ethylene and propylene is obviously improved. Thus, the method of the present disclosure can effectively improve the yields and selectivity of ethylene and propylene by continuing the reaction of macromolecular olefins produced in the oil-gas separation process in a specific route.

According to the present disclosure, the olefins in the olefin-rich stream are C5 or greater olefins; the olefin-rich stream has a content of olefins greater than C5 of from 50 wt% to 100 wt%.

As a preferred embodiment of the present disclosure, the method may further comprise: carrying out burning regeneration on the spent catalyst to obtain a regenerated catalyst; and returning the regenerated catalyst to a downstream reaction zone of the U-shaped tube reactor after preheating.

As a second preferred embodiment of the present disclosure, the method may further include: and introducing the butene into a downlink reaction zone of the U-shaped pipe reactor for continuous reaction.

Preferably, the butene is contacted with the catalytic conversion catalyst prior to the olefin-rich stream. The difficulty of hydrocarbon cracking is increased along with the decrease of the carbon number, and the energy required by butene cracking is higher, so that if the butene is preferably contacted with a high-temperature catalytic conversion catalyst, the olefin containing more than C5 is contacted with the catalytic conversion catalyst, the butene conversion rate and the selectivity of ethylene and propylene of products can be improved, more byproducts are prevented from being generated by simultaneous feeding of the olefin, and the high-efficiency utilization of resources is realized.

According to the present disclosure, the reaction conditions for the continued reaction of butenes may include: the reaction temperature is 650-800 ℃, the reaction pressure is 0.05-1MPa, the reaction time is 0.01-10s, and the weight ratio of the catalytic conversion catalyst to the butene is (20-200): 1, a step of; preferably, the reaction temperature is 680-780 ℃, the reaction pressure is 0.1-0.8MPa, the reaction time is 0.05-8s, and the weight ratio of the catalytic conversion catalyst to the butene is (30-180): 1.

As a third preferred embodiment of the present disclosure, the method may further include: and (3) carrying out hydrotreating on the catalytic wax oil to obtain hydrogenation catalytic wax oil, and introducing the hydrogenation catalytic wax oil into the bottom of the uplink reaction zone of the U-shaped tube reactor to continue to react. The method of the present disclosure continuously reacts the hydrogenation catalytic wax oil, further improves the yields of ethylene and propylene, and realizes the effective utilization of petroleum resources.

According to the present disclosure, the hydrotreating conditions may include: the hydrogen partial pressure is 3.0-20.0 megapascals, the reaction temperature is 300-450 ℃, the hydrogen-oil volume ratio is 300-2000, and the volume airspeed is 0.1-3.0 hours ^-1.

In the present disclosure, the catalyst used in the hydrotreatment may be well known to those skilled in the art, and may include, for example, a support, which is alumina and/or amorphous silica alumina, and a metal component, which is a group VIB metal and/or a group VIII metal, and an optional additive, which is at least one selected from fluorine, phosphorus, titanium, and platinum, supported on the support. Specifically, the VIB group metal is Mo or/and W, and the VIII group metal is Co or/and Ni; the content of the additive is 0-10 wt%, the content of the VIB group metal is 12-39 wt% and the content of the VIII group metal is 1-9 wt% based on the weight of the hydrotreated catalyst.

According to the present disclosure, the conditions of the first catalytic conversion reaction may include: the reaction temperature is 600-800 ℃, the reaction pressure is 0.05-1MPa, the reaction time is 0.01-100 seconds, and the weight ratio of the catalytic conversion catalyst to the hydrocarbon oil raw material is (1-200): 1, a step of; preferably, the reaction temperature is 630-780 ℃, the reaction pressure is 0.1-0.8MPa, the reaction time is 0.1-80 seconds, and the weight ratio of the catalytic conversion catalyst to the hydrocarbon oil raw material is (3-180): 1, a step of; more preferably, the reaction temperature is 650-750 ℃, the reaction pressure is 0.2-0.5MPa, the reaction time is 0.2-70 seconds, and the weight ratio of the catalytic conversion catalyst to the hydrocarbon oil raw material is (4-150): 1. the conditions of the second catalytic conversion reaction may include: the reaction temperature is 400-650 ℃, the reaction pressure is 0.05-1MPa, the reaction time is 0.01-100 seconds, and the weight ratio of the catalytic conversion catalyst to the heavy raw oil is (1-100): 1, a step of; preferably, the reaction temperature is 450-600 ℃, the reaction pressure is 0.1-0.8MPa, the reaction time is 0.1-80 seconds, and the weight ratio of the catalytic conversion catalyst to the heavy raw oil is (3-70): 1, a step of; further preferably, the reaction temperature is 480-580 ℃, the reaction pressure is 0.2-0.5MPa, the reaction time is 0.2-70 seconds, and the weight ratio of the catalytic conversion catalyst to the heavy raw oil is (4-30): 1.

The downstream reaction zone of the U-tube reactor in the present disclosure may further be provided with a separator for performing preliminary separation on the reaction product in step S1, further separating oil gas obtained by the preliminary separation, and sending the catalyst obtained by the preliminary separation to the upstream reaction zone of the U-tube reactor to perform catalytic cracking reaction with heavy feedstock oil.

According to the present disclosure, the olefin content in the hydrocarbon oil feedstock may be 80 wt% or more; preferably, the olefin content in the hydrocarbon oil feedstock is 90 wt% or more; more preferably, the hydrocarbon oil feedstock is a pure olefin feedstock. The heavy feedstock oil may be selected from petroleum hydrocarbons and/or mineral oils; the petroleum hydrocarbon may be at least one selected from the group consisting of vacuum gas oil, atmospheric gas oil, coker gas oil, deasphalted oil, vacuum residue, atmospheric residue, and heavy aromatic raffinate oil; the mineral oil may be selected from at least one of coal liquefied oil, oil sand oil, and shale oil.

In one embodiment of the present disclosure, the olefins in the hydrocarbon oil feedstock may be from a C5 or higher fraction produced by dehydrogenation of an alkane feedstock, a C5 or higher fraction produced by a catalytic cracker in a refinery, a C5 or higher fraction produced by a steam cracker in an ethylene plant, an olefin-rich fraction of C5 or higher by MTO by-products, an olefin-rich fraction of C5 or higher by-products of MTP; the paraffinic feedstock is selected from at least one of naphtha, aromatic raffinate, and light hydrocarbons of other units.

Further, in the butene product recycle embodiment, the olefins in the hydrocarbon oil feedstock are derived from a C4 or higher fraction produced by dehydrogenation of an alkane feedstock, a C4 or higher fraction produced by a catalytic cracker in a refinery, a C4 or higher fraction produced by a steam cracker in an ethylene plant, an olefin-rich fraction of C4 or higher by an MTO byproduct, an olefin-rich fraction of C4 or higher by an MTP byproduct; the paraffinic feedstock is selected from at least one of naphtha, aromatic raffinate, and light hydrocarbons of other units.

In one embodiment, the method for preparing olefin by alkane dehydrogenation comprises the steps of carrying out contact reaction on alkane and a dehydrogenation catalyst, wherein the inlet temperature of a reactor is 400-700 ℃, the volume space velocity of the alkane is 200-5000h ^-1, and the pressure of the contact reaction is 0-1.0MPa. The dehydrogenation catalyst consists of a carrier, an active component and an auxiliary agent, wherein the active component and the auxiliary agent are loaded on the carrier; the content of the carrier is 60-90 wt%, the content of the active component is 8-35 wt% and the content of the auxiliary agent is 0.1-5 wt% based on 100% of the total weight of the catalyst; the carrier is alumina containing a modifier; the content of the modifier is 0.1-2 wt% of the total weight of the catalyst, and the modifier is La or Ce; the active component is platinum or chromium; the auxiliary agent is bismuth and alkali metal components or bismuth and alkaline earth metal components; the molar ratio of bismuth to the active component is 1 (5-50); the molar ratio of bismuth to alkali metal component is 1: (0.1-5); the molar ratio of bismuth to alkaline earth metal component is 1: (0.1-5); the alkali metal component is one or more of Li, na and K; the alkaline earth metal component is one or more of Mg, ca and Ba.

According to the present disclosure, the catalytic conversion catalyst may include 1 to 50 wt% of a molecular sieve, 5 to 99 wt% of an inorganic oxide, and 0 to 70 wt% of clay, based on the weight of the catalytic conversion catalyst; the molecular sieve may include one or more of a large pore molecular sieve, a medium pore molecular sieve, and a small pore molecular sieve. The catalytic conversion catalyst further comprises 0.1 to 3 wt% of an active metal, based on the weight of the catalytic conversion catalyst; the active metal is selected from one or more of VIII group metal, IVA group metal and rare earth metal. The inorganic oxide in the present disclosure as a binder may be selected from silica (SiO ₂) and/or alumina (Al ₂O₃); the clay in the present disclosure may be selected from kaolin and/or halloysite as a matrix. In a specific embodiment of the present disclosure, the catalytic cracking catalyst may further support a metal ion, the metal ion is selected from at least one of a non-metal element, a transition metal element and a rare earth metal element, wherein the non-metal element may be phosphorus, the transition metal element may be selected from iron, cobalt and nickel, and the weight of the modified element is 0.1% -3% of the weight of the catalytic cracking catalyst.

In one embodiment of the present disclosure, the mesoporous molecular sieve may be a ZSM molecular sieve, further, the ZSM molecular sieve may be one or more selected from ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48. The small pore molecular sieve can be a SAPO molecular sieve, and further, the SAPO molecular sieve can be one or more selected from SAPO-34, SAPO-11 and SAPO-47. The macroporous molecular sieve can be selected from one or more than one of rare earth Y molecular sieve, rare earth hydrogen Y molecular sieve, ultrastable Y molecular sieve, high silicon Y molecular sieve, beta molecular sieve and other molecular sieves with similar structures.

In a specific embodiment of the disclosure, as shown in fig. 1, a hydrocarbon oil raw material with an olefin content of more than 50wt% enters the upper part of a downlink reaction zone i of a U-shaped tube reactor 1 through a pipeline 2, contacts and reacts with a thermocatalytic conversion catalyst from a pipeline 17, a reaction product is separated from the bottom of the downlink reaction zone i through a cyclone separator 3, separated oil gas is introduced into a separation system through a pipeline 4, the separated catalyst continuously participates in the reaction along with the downlink of the U-shaped tube reactor 1, and heavy raw material oil enters the bottom of an uplink reaction zone ii of the U-shaped tube reactor 1 through a pipeline 5, contacts and reacts with an existing material flow in the U-shaped tube reactor 1. The generated reactant flow and the deactivated spent catalyst enter a cyclone separator in a settler 8 through an outlet section 7 to realize the separation of the spent catalyst and the reactant flow, the reactant flow enters a gas collection chamber 9, the spent catalyst fine powder returns to the settler through a dipleg, and the oil gas stripped from the spent catalyst enters the gas collection chamber 9 after passing through the cyclone separator. The stripped spent catalyst enters a regenerator 13 through a diagonal pipe 12, and main air enters the regenerator through a pipeline 14 to burn coke on the spent catalyst and regenerate the deactivated spent catalyst. The flue gas enters the extractor through line 16. The preheated regenerated catalyst is passed via line 17 to the downstream reaction zone I. The reactant flow enters a subsequent product separation device 18 through a large oil-gas pipeline 10, the separated hydrogen, methane and ethane are led out through a pipeline 19, the ethylene is led out through a pipeline 20, the propylene is led out through a pipeline 21, the butylene is led into the top of a downlink reaction zone I of the U-shaped pipe reactor 1 for continuous reaction, the propane and butane are led out through a pipeline 23, the crude material containing the olefins is led into an olefin separation device 26 through a pipeline 24, the separated material flow rich in the olefins is led out through a pipeline 27, the material flow rich in the olefins is led into the upper part of the downlink reaction zone I of the U-shaped pipe reactor 1 for continuous reaction through a pipeline 28, the catalytic wax oil is led into a hydrotreating reactor 29 through a pipeline 25, the hydrotreated light component is led out through a pipeline 11, and the hydrocatalytic wax oil is led into the bottom of the uplink reaction zone II of the U-shaped pipe reactor 1 through a pipeline 6 for continuous reaction.

In a second specific embodiment of the disclosure, as shown in fig. 2, a hydrocarbon oil feedstock with an olefin content of more than 50wt% enters the upper part of the downstream reaction zone i of the U-tube reactor 1 through a pipeline 2, contacts and reacts with a thermocatalytic conversion catalyst from a pipeline 17, the material flow obtained by the reaction is separated from the bottom of the downstream reaction zone i through a cyclone separator 3, the separated material flow is introduced into a separation system through a pipeline 4, and the separated catalyst continues to participate in the reaction along with the downstream reaction zone of the U-tube reactor 1. Heavy raw oil enters the bottom of an uplink reaction zone II of the U-shaped pipe reactor 1 through a pipeline 5, contacts and reacts with the existing material flow containing the catalyst. The generated reactant flow and the deactivated spent catalyst enter a cyclone separator in a settler 8 through an outlet section 7 to realize the separation of the spent catalyst and the reactant flow, the reactant flow enters a gas collection chamber 9, the spent catalyst fine powder returns to the settler through a dipleg, and the oil gas stripped from the spent catalyst enters the gas collection chamber 9 after passing through the cyclone separator. The stripped spent catalyst enters a regenerator 13 through a diagonal pipe 12, and main air enters the regenerator through a pipeline 14 to burn coke on the spent catalyst and regenerate the deactivated spent catalyst. The flue gas enters the extractor through line 16. The preheated regenerated catalyst enters the downstream reaction zone I of the U-tube reactor 1 via line 17. The reactant flow enters a subsequent product separation device 18 through a large oil-gas pipeline 10, hydrogen, methane and ethane obtained through separation are led out through a pipeline 19, ethylene is led out through a pipeline 20, propylene is led out through a pipeline 21, butylene is led out through a pipeline 22, propane and butane are led out through a pipeline 23, an olefin-containing crude material flow is led into an olefin separation device 26 through a pipeline 24, a hydrocarbon flow obtained through separation is led out through a pipeline 27, an olefin-rich flow is led into the upper part of a downlink reaction zone I of the U-shaped tube reactor 1 through a pipeline 28 for continuous reaction, catalytic wax oil is led into a hydrotreating reactor 29 through a pipeline 25, light components obtained after hydrotreating are led out through a pipeline 11, and hydrogenation catalytic wax oil is led into the bottom of an uplink reaction zone II of the U-shaped tube reactor 1 through a pipeline 6 for continuous reaction.

The present disclosure is further illustrated in detail by the following examples. The starting materials used in the examples are all available commercially.

The materials a and b used in the examples are heavy raw oil, and the material c is a light fraction of catalytically cracked gasoline, and the properties thereof are shown in table 1.

The preparation of the catalytic conversion catalyst A used in the examples is briefly described below: 969 g of halloysite (product of China Kaolin Co., ltd., solid content: 73%) is beaten by 4300 g of decationized water, 781 g of pseudo-boehmite (product of Shandong Zibo aluminum stone Co., ltd., solid content: 64%) and 144 ml of hydrochloric acid (concentration: 30% and specific gravity: 1.56) are added to be stirred uniformly, the mixture is left to stand and age at 60 ℃ for 1 hour, the pH value is kept at 2-4, the mixture is cooled to normal temperature, and 5000 g of slurry prepared in advance is added, wherein 1600g of medium pore ZSM-5 molecular sieve and large pore Y-type molecular sieve (product of China Oldham catalyst Co., ltd.) are added, and the weight ratio of the two is 9:1. Stirring uniformly, spray drying, and washing free Na ⁺ to obtain the catalyst. The resulting catalyst was aged at 800 ℃ and 100% steam, and the aged catalyst was referred to as catalyst a, the properties of which are shown in table 3.

The catalytic conversion catalyst B used in the examples was sold under the trade name CEP-1, the catalytic conversion catalyst C was sold under the trade name CHP-1, both of which were industrial products produced by Qilu division, china petrochemical catalyst, and the properties of the catalytic conversion catalyst B and the catalytic conversion catalyst C are shown in Table 3.

The preparation of hydrotreating catalyst D used in the examples is briefly described as follows: ammonium metatungstate ((NH ₄)₂W₄O₁₃·18H₂ O, chemically pure) and nickel nitrate (Ni (NO ₃)₂·18H₂ O, chemically pure)) were weighed and made up into 200 ml of solution with water, the solution was added to 50 g of alumina carrier, immersed for 3 hours at room temperature, the immersed solution was treated with ultrasonic waves for 30 minutes during the immersion, cooled, filtered, and dried in a microwave oven for about 15 minutes, the composition of the catalyst was 30.0 wt% WO ₃, 3.l wt% NiO and the balance alumina, denoted as catalyst D.

The hydrodesulfurization catalyst E used in the examples was prepared as follows: 1000 g of pseudo-boehmite produced by China petrochemical catalyst Chang Ling is weighed, 1000 ml of aqueous solution containing 10ml of nitric acid (chemical purity) is added, the mixture is extruded and molded on a double-screw extruder, and the mixture is dried at 120 ℃ for 4 hours and baked at 800 ℃ for 4 hours to obtain the catalyst carrier. Immersing with 900 ml of aqueous solution containing 120 g of ammonium fluoride for 2 hours, drying at 120 ℃ for 3 hours, and roasting at 600 ℃ for 3 hours; after cooling to room temperature, the catalyst was immersed in 950 ml of an aqueous solution containing 133 g of ammonium meta-molybdate for 3 hours, dried at 120℃for 3 hours, calcined at 600℃for 3 hours, cooled to room temperature, immersed in 900 ml of an aqueous solution containing 180 g of nickel nitrate, 320 g of ammonium meta-tungstate for 4 hours, and immersed in a mixed aqueous solution of 0.1% by weight of ammonium meta-molybdate (chemically pure) and 0.1% by weight of nickel nitrate (chemically pure) relative to the catalyst carrier for 4 hours, and dried at 120℃for 3 hours, and calcined at 600℃for 4 hours, to obtain catalyst E.

TABLE 1-1

TABLE 1-2

Feed Properties	b
		Density (20 ℃ C.)/(kg/m ³)	901.5
Kangshi carbon residue, weight percent	4.9
		H, wt%	12.86
S, weight percent	0.16
		N, wt%	0.26
Ni, microgram/gram	6.2
		Group composition, weight percent
Saturated hydrocarbons	54.8
		Aromatic hydrocarbons	28.4
Colloid	16.0
		Asphaltenes	0.8

TABLE 2

Carbon number	Alkane, weight percent	Olefins, wt%	Naphthene, weight percent	Aromatic hydrocarbon, weight percent
					5	1.80	4.40	0.00	0.00
6	3.91	30.20	0.67	0.87
					7	3.32	33.57	1.81	2.39
8	2.79	10.68	1.75	0.59
					9	0.40	0.85	0.00	0.00

TABLE 3 Table 3

Example 1

This example was carried out according to the procedure of fig. 1, with tests carried out on a medium-sized apparatus of a U-tube reactor. The 1-pentene and the high-temperature catalytic conversion catalyst A are contacted at the top of a downstream reaction zone, the reaction pressure is 0.1MPa at the reaction temperature of 700 ℃, the reaction time is 5s, and the weight ratio of the catalyst to the raw materials is 45:1, the heavy raw oil a and the catalytic conversion catalyst A are contacted at the bottom of an upstream reaction zone, the reaction temperature is 530 ℃, the reaction pressure is 0.1MPa, the reaction time is 6s, and the weight ratio of the catalyst to the raw materials is 5: the catalytic conversion reaction is carried out under the condition 1, and the ratio of 1-pentene to heavy raw materials is 1:9. separating the reaction product and the spent catalyst, introducing the spent catalyst into a regenerator for burning and regenerating, and introducing the obtained reaction product into a combined separation system to obtain products including ethylene, propylene, butylene, olefin-rich stream (boiling point is less than 250 ℃) and catalytic cracking wax oil (boiling point is more than 250 ℃), and the like. The catalytic cracking wax oil and the hydrogenation catalyst D react at 350 ℃, the hydrogen partial pressure is 18MPa, the volume space velocity is 1.5 hours ^-1, and the hydrogen-oil volume ratio is 1500 to obtain the hydrogenation catalytic cracking wax oil. Introducing the obtained butene into the top of a downlink reaction zone for recracking, wherein the reaction temperature is 740 ℃, and the weight ratio of the catalyst to the raw materials is 100:1, the reaction time is 0.2s; introducing the olefin-rich material flow into the top of a downstream reaction zone for recracking, wherein the reaction temperature is 700 ℃ and the reaction time is 5s; the hydrocracking wax oil is mixed with heavy raw oil and then returned to the bottom of the uplink reaction zone for reaction. The reaction conditions and product distribution are shown in Table 4.

Comparative example 1

This comparative example was run on a medium-sized unit in a U-tube reactor using a similar process flow to example 1, except that the olefin-rich stream was not returned to the U-tube reactor for continued reaction. The heavy raw oil a and the catalytic conversion catalyst A are contacted at the bottom of an upstream reaction zone, the reaction temperature is 530 ℃, the reaction pressure is 0.1MPa, the reaction time is 6s, and the weight ratio of the catalyst to the raw materials is 5:1, a catalytic conversion reaction occurs. Separating the reaction product and the spent catalyst, introducing the spent catalyst into a regenerator for burning and regenerating, and introducing the reaction product into a combined separation system to obtain products including ethylene, propylene, butylene, olefin-rich material flow, catalytic cracking wax oil and the like. The catalytic cracking wax oil and the hydrogenation catalyst D react at 350 ℃, the hydrogen partial pressure is 18MPa, the volume space velocity is 1.5 hours ^-1, and the hydrogen-oil volume ratio is 1500 to obtain the hydrogenation catalytic cracking wax oil. Introducing the obtained butene into the top of a downlink reaction zone for recracking, wherein the reaction temperature is 740 ℃, and the weight ratio of the catalyst to the raw materials is 100:1, the reaction time is 0.2s; the hydrocracking wax oil is mixed with heavy raw oil and then returned to the bottom of the uplink reaction zone for reaction. The reaction conditions and product distribution are shown in Table 4.

Example 2

This example was performed according to the process flow of fig. 1, with tests performed on a medium-sized apparatus of a U-tube reactor. The heavy raw oil a and the catalytic conversion catalyst A are contacted at the bottom of an upstream reaction zone, the reaction temperature is 530 ℃, the reaction pressure is 0.1MPa, the reaction time is 6s, and the weight ratio of the catalyst to the raw materials is 5:1, separating the reaction product and the spent catalyst, introducing the spent catalyst into a regenerator together for burning regeneration, and introducing the obtained reaction product into a combined separation system together to obtain products including ethylene, propylene, butylene, olefin-rich material flow, catalytic cracking wax oil and the like. The catalytic cracking wax oil and the hydrogenation catalyst D react at 350 ℃, the hydrogen partial pressure is 18MPa, the volume space velocity is 1.5 hours ^-1, and the hydrogen-oil volume ratio is 1500 to obtain the hydrogenation wax oil. Introducing the obtained butene into the top of a downlink reaction zone for recracking, wherein the reaction temperature is 740 ℃, and the weight ratio of the catalyst to the raw materials is 100:1, the reaction time is 0.2s; introducing the olefin-rich material flow into the top of a downstream reaction zone for recracking, wherein the reaction temperature is 700 ℃ and the reaction time is 5s; the hydrocracking wax oil is mixed with heavy raw oil and then returned to the bottom of the uplink reaction zone for reaction. The reaction conditions and product distribution are shown in Table 4.

Comparative example 2

The test was carried out on a medium-sized apparatus of a riser reactor, with heavy feedstock oil a and catalytic conversion catalyst B in contact at the bottom of the riser, at a reaction temperature of 610 ℃, a weight ratio of catalyst to feedstock of 16.9:1, the reaction pressure is 0.1MPa, catalytic conversion reaction is carried out under the reaction time of 6s, and the product is not hydrotreated and returns to the device to continue the reaction. The reaction conditions and product distribution are shown in Table 4.

Example 3

This example uses a similar process flow to example 2, except that a heavier feed b is used. The catalytic cracking wax oil is not subjected to deep hydrogenation treatment, contacts with a hydrodesulfurization catalyst E in a hydrodesulfurization reactor, and reacts under the conditions of a reaction pressure of 6.0MPa, a reaction temperature of 350 ℃, a hydrogen-oil volume ratio of 350 and a volume space velocity of 2.0 hours ^-1 to obtain low-sulfur hydrogenated distillate oil as a light oil component. The reaction conditions and product distribution are shown in Table 4.

Comparative example 3

The test was carried out on a medium-sized apparatus of a riser reactor, with heavy feedstock oil b and catalytic conversion catalyst C in contact at the bottom of the riser, at a reaction temperature of 530 ℃, with a weight ratio of catalyst to feedstock of 5:1, the reaction pressure is 0.1MPa, the catalytic conversion reaction is carried out under the reaction time of 6s, the hydrotreating is basically the same as in the example 3, and the product is not returned to the device for continuous reaction. The reaction conditions and product distribution are shown in Table 4.

Example 4

This example employs a similar process flow to example 1, except that the first reactor feed is catalytic light gasoline fraction C. The reaction conditions and product distribution are shown in Table 4.

TABLE 4 Table 4

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As can be seen from table 4: when the olefin is cracked at high temperature, the yield of ethylene and propylene is higher, the content of ethylene and propylene in the obtained reaction product is about 56%, and the effect is better when the content of the olefin in the raw material is higher. In addition, the yields of benzene, toluene and xylene from olefin cracking are obviously increased.

Example 5

This example was carried out according to the procedure of fig. 2, with tests carried out on a medium-sized apparatus of a U-tube reactor. The 1-pentene and the high-temperature catalytic conversion catalyst A are contacted at the top of a downstream reaction zone, the reaction pressure is 0.1MPa at the reaction temperature of 700 ℃, the reaction time is 5s, and the weight ratio of the catalyst to the raw materials is 45:1, the heavy raw oil a and the catalytic conversion catalyst A are contacted at the bottom of an upstream reaction zone, the reaction temperature is 530 ℃, the reaction pressure is 0.1MPa, the reaction time is 6s, and the weight ratio of the catalyst to the raw materials is 5: the catalytic conversion reaction is carried out under the condition 1, and the ratio of 1-pentene to heavy raw materials is 1:9. separating the reaction product and the spent catalyst, introducing the spent catalyst into a regenerator for burning and regenerating, and introducing the reaction product into a combined separation system to obtain products including ethylene, propylene, butylene, olefin-rich material flow, catalytic cracking wax oil and the like. The catalytic cracking wax oil and the hydrogenation catalyst D react at 350 ℃, the hydrogen partial pressure is 18MPa, the volume space velocity is 1.5 hours ^-1, and the hydrogen-oil volume ratio is 1500 to obtain the hydrogenation catalytic cracking wax oil. Introducing the obtained olefin-rich stream into the top of a downstream reaction zone for recracked, wherein the reaction temperature is 700 ℃ and the reaction time is 5s; the hydrocracking wax oil is mixed with heavy raw oil and then returned to the bottom of the uplink reaction zone for reaction. The ethylene yield in the product is 6.48%, the propylene yield is 21.46%, the butene yield is 24.96%, and the total triene yield is up to 52.90%.

The preferred embodiments of the present disclosure have been described in detail above, but the present disclosure is not limited to the specific details of the above embodiments, and various simple modifications may be made to the technical solutions of the present disclosure within the scope of the technical concept of the present disclosure, and all the simple modifications belong to the protection scope of the present disclosure.

In addition, the specific features described in the above embodiments may be combined in any suitable manner without contradiction. The various possible combinations are not described further in this disclosure in order to avoid unnecessary repetition.

Moreover, any combination between the various embodiments of the present disclosure is possible as long as it does not depart from the spirit of the present disclosure, which should also be construed as the disclosure of the present disclosure.

Claims

1. A catalytic conversion process for producing light olefins, the process comprising:

S3, separating ethylene, propylene, butene, an olefin-rich stream and catalytic wax oil from the first reactant stream and the second reactant stream; introducing the olefin-rich stream into a downstream reaction zone of the U-tube reactor for continued reaction;

the olefins in the olefin-rich stream are olefins above C5; the olefin-rich stream has a content of olefins greater than C5 of from 50 wt% to 100 wt%;

the method further comprises the steps of: introducing the butene into a downlink reaction zone of the U-shaped pipe reactor for continuous reaction; contacting said butene with said catalytic conversion catalyst prior to said olefin-rich stream; the reaction conditions for the continued reaction of the butene include: the reaction temperature is 650-800 ℃, the reaction pressure is 0.05-1MPa, the reaction time is 0.01-10 seconds, and the weight ratio of the catalytic conversion catalyst to the butene is (20-200): 1, a step of;

The conditions of the first catalytic conversion reaction include: the reaction temperature is 600-800 ℃, the reaction pressure is 0.05-1MPa, the reaction time is 0.01-100 seconds, and the weight ratio of the catalytic conversion catalyst to the hydrocarbon oil raw material is (1-200): 1, a step of; the conditions of the second catalytic conversion reaction include: the reaction temperature is 400-650 ℃, the reaction pressure is 0.05-1MPa, the reaction time is 0.01-100 seconds, and the weight ratio of the catalytic conversion catalyst to the heavy raw oil is (1-100): 1, a step of;

The heavy feedstock oil is selected from petroleum hydrocarbons and/or mineral oils; the petroleum hydrocarbon is at least one selected from vacuum gas oil, normal pressure gas oil, coker gas oil, deasphalted oil, vacuum residue, normal pressure residue and heavy aromatic raffinate oil; the mineral oil is selected from at least one of coal liquefied oil, oil sand oil and shale oil; the olefins in the hydrocarbon oil raw material come from C4 above fraction generated by dehydrogenation of alkane raw material, C4 above fraction generated by catalytic cracking device of oil refinery, C4 above fraction generated by steam cracking device of ethylene plant, olefin-rich fraction of C4 above MTO byproduct and olefin-rich fraction of C4 above MTP byproduct; the alkane feed is selected from at least one of naphtha, aromatic raffinate oil, and light hydrocarbons.

2. The method of claim 1, wherein the method further comprises: carrying out burning regeneration on the spent catalyst to obtain a regenerated catalyst; and

And returning the regenerated catalyst to a downstream reaction zone of the U-shaped pipe reactor.

3. The method of claim 1, wherein the reaction conditions under which the butene continues to react comprise: the reaction temperature is 680-780 ℃, the reaction pressure is 0.1-0.8MPa, the reaction time is 0.05-8 seconds, and the weight ratio of the catalytic conversion catalyst to the butene is (30-180): 1.

4. The method of claim 1, wherein the method further comprises: and (3) carrying out hydrotreating on the catalytic wax oil to obtain hydrogenation catalytic wax oil, and introducing the hydrogenation catalytic wax oil into the bottom of the uplink reaction zone of the U-shaped tube reactor to continue to react.

5. The method of claim 4, wherein,

The hydrotreating conditions include: the hydrogen partial pressure is 3.0-20.0 megapascals, the reaction temperature is 300-450 ℃, the hydrogen-oil volume ratio is 300-2000, and the volume airspeed is 0.1-3.0 hours ^-1.

6. The method of claim 1, wherein the conditions of the first catalytic conversion reaction comprise: the reaction temperature is 630-780 ℃, the reaction pressure is 0.1-0.8MPa, the reaction time is 0.1-80 seconds, and the weight ratio of the catalytic conversion catalyst to the hydrocarbon oil raw material is (3-180): 1, a step of;

The conditions of the second catalytic conversion reaction include: the reaction temperature is 450-600 ℃, the reaction pressure is 0.1-0.8MPa, the reaction time is 0.1-80 seconds, and the weight ratio of the catalytic conversion catalyst to the heavy raw oil is (3-70): 1.

7. The method of claim 1, wherein the olefin content in the hydrocarbon oil feedstock is 80 wt% or more.

8. The method according to claim 7, wherein the olefin content in the hydrocarbon oil feedstock is 90 wt% or more.

9. The method of claim 8, wherein the hydrocarbon oil feedstock is a pure olefin feedstock.

10. The process of claim 1, wherein the catalytic conversion catalyst comprises 1-50 wt.% molecular sieve, 5-99 wt.% inorganic oxide, and 0-70 wt.% clay, based on the weight of the catalytic conversion catalyst;

the molecular sieve comprises one or more of a large pore molecular sieve, a medium pore molecular sieve and a small pore molecular sieve;

the catalytic conversion catalyst further comprises 0.1 to 3 wt% of an active metal, based on the weight of the catalytic conversion catalyst; the active metal is selected from one or more of VIII group metal, IVA group metal and rare earth metal.