CN114763485B

CN114763485B - Catalytic conversion method for preparing ethylene and propylene

Info

Publication number: CN114763485B
Application number: CN202110032110.6A
Authority: CN
Inventors: 许友好; 左严芬; 王新; 何鸣元; 沙有鑫; 白旭辉
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: 2023-07-11
Anticipated expiration: 2041-01-11
Also published as: CN114763485A

Abstract

The present disclosure relates to a catalytic conversion process for producing ethylene and propylene, 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 ℃ and is subjected to catalytic conversion reaction in a first reactor, so as to obtain a first reactant flow and a first spent catalyst; s2, contacting heavy raw oil with a catalytic conversion catalyst with the temperature of more than 650 ℃ and carrying out catalytic conversion reaction in a second reactor to obtain a second reactant flow and a second spent catalyst; s3, separating the first reactant flow and the second reactant flow to obtain ethylene, propylene, butylene, a flow containing olefins with more than C5 and catalytic wax oil; the butene and the stream containing olefins above C5 are respectively introduced into the first reactor for continuous reaction. The process of the present disclosure has high ethylene and propylene yields, high selectivity and low methane yields.

Description

Catalytic conversion method for preparing ethylene and propylene

Technical Field

The application relates to petroleum refining and petrochemical processing, in particular to a catalytic conversion method for preparing ethylene and propylene.

Background

Ethylene is one of the most chemical products in the world, accounting for more than 75% of the global overall petrochemical yield; bulk downstream products of ethylene are mainly polyethylene, ethylene oxide, ethylene glycol, polyvinyl chloride, styrene, vinyl acetate, and the like. Propylene is an important organic chemical raw material and is mainly used for preparing acrylonitrile, propylene oxide, acetone and the like. Ethylene and propylene are increasingly demanded as important chemical intermediates.

The traditional steam cracking process is adopted to prepare ethylene and propylene, and the demand for light hydrocarbon, naphtha and other chemical light hydrocarbons is large. 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 via 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 raw material, and the reaction temperature is 400-680 ℃, the reaction pressure is-0.09-1.0 MPa, and the weight airspeed is 0.1-50 hours ^-1 The reaction is carried out under conditions such that the product ethylene/propylene is lower, less than 0.41, and increases with increasing temperature, while hydrogen, methane and ethane increase.

Accordingly, there is a need in the art for a new process for producing ethylene and propylene in high yields to achieve efficient utilization of petroleum resources.

Disclosure of Invention

The present disclosure aims to provide a catalytic conversion process for producing ethylene and propylene, further improving the ethylene and propylene yields.

In order to achieve the above object, the present disclosure provides a catalytic conversion method for producing more ethylene and propylene, comprising the steps of:

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 ℃ and is subjected to catalytic conversion reaction in a first catalytic conversion reactor, so as to obtain a first reactant flow and a first spent catalyst;

s2, contacting heavy raw oil with a catalytic conversion catalyst with the temperature of more than 650 ℃ and carrying out catalytic conversion reaction in a second catalytic conversion reactor to obtain a second reactant flow and a second spent catalyst;

s3, separating the first reactant flow and the second reactant flow to obtain ethylene, propylene, butylene, a flow containing olefins with more than C5 and catalytic wax oil; the butenes and the stream containing olefins having more than C5 s are separately introduced into the first catalytic conversion reactor to continue the reaction.

Optionally, in step S3, the butenes are contacted with the catalytic conversion catalyst prior to the C5 or higher containing stream, and the C5 or higher containing stream is co-fed in admixture with the hydrocarbon oil feedstock.

Optionally, the content of olefins in the stream containing olefins above C5 is above 50 wt%; the olefins in the olefin-rich stream are olefins above C5.

Optionally, the method further comprises: the first spent catalyst and the second spent catalyst are sent to a shared regenerator for burning regeneration, and a regenerated catalyst is obtained; the regenerated catalyst is returned to the first catalytic conversion reactor and the second catalytic conversion reactor.

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.

optionally, the reaction conditions under which the butene is introduced into the first catalytic conversion reactor to continue the reaction 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.

preferably, the reaction conditions under which the butene is introduced into the first catalytic conversion reactor for continuous reaction include: 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) hydrotreating the catalytic wax oil to obtain hydrogenation catalytic wax oil, and mixing the hydrogenation catalytic wax oil with the heavy raw oil and then jointly entering the second catalytic conversion reactor.

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 olefin content in the hydrocarbon oil feedstock is above 80 wt%; preferably, the olefin content in the hydrocarbon oil feedstock is above 90 wt.%; more preferably, the hydrocarbon oil feedstock is a pure olefin feedstock; 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 higher fraction produced by dehydrogenation of an alkane feedstock, a C4 or higher fraction produced by a catalytic cracking unit of a refinery, a C4 or higher fraction of a steam cracking unit in an ethylene plant, an olefin-rich fraction of a C4 or higher byproduct of MTO, and an olefin-rich fraction of a C4 or higher byproduct of MTP.

Optionally, the paraffinic feedstock is selected from at least one of naphtha, aromatic raffinate, and light hydrocarbons.

Alternatively, the catalytic conversion catalyst comprises 1 to 50 weight percent molecular sieve, 5 to 99 weight percent inorganic oxide, and 0 to 70 weight percent clay, based on the weight of the catalytic conversion catalyst;

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

optionally, the catalytic conversion catalyst further comprises 0.1 to 3 wt% 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.

By the technical scheme, the method disclosed by the invention is used for respectively cracking the raw material containing olefin and the heavy raw material oil at high temperature, and returning the olefin in the separated product to the device for continuous reaction. The macromolecule olefin produced in the oil-gas separation process is recycled in a specific route, so that the methane yield is reduced, the purpose of high-efficiency utilization of petroleum resources is achieved, the traditional scheme of producing ethylene and propylene by high-energy-consumption steam cracking can be replaced, and meanwhile, the method disclosed by the invention has the advantages of high ethylene and propylene yield, selectivity and low methane yield. And the present disclosure uses the shared regenerator to burn and regenerate the first spent catalyst and the second spent catalyst, and then send the first spent catalyst and the second spent catalyst back to the first catalytic conversion reactor and the second catalytic conversion reactor respectively, a large amount of coke with higher graphitization degree is inevitably generated when heavy raw oil is cracked on the acid catalyst, and the heat brought by the coke combustion can be used for providing energy for olefin high-temperature cracking, thereby not only realizing the cyclic utilization of the catalytic conversion catalyst, but also further improving the utilization rate of petroleum resources, and simultaneously reducing the energy consumption of the device.

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 flow diagram of one embodiment of the present disclosure.

Description of the reference numerals

A first catalytic conversion reactor B second catalytic conversion reactor

1 line 2 line 3 line

5 pipeline 6 stripping section 7 outlet section

8 settler 9 plenum 10 pipeline

12 inclined tube 13 regenerator 14 pipeline

15 hydrotreatment reactor 16 line 17 line

18 line 19 inclined line 20 line

21 line 22 line 23 line

24 line 25 line 26 stripping section

27 outlet section 28 settler 29 plenum

30 line 31 product separation device 32 line

33 line 34 line 35 line

36 line 37 line 38 line

39 line 40 olefin separation plant 41 line

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 the production of high yields of ethylene and propylene, comprising the steps of:

s3, separating the first reactant flow and the second reactant flow to obtain ethylene, propylene, butylene, a flow containing olefins with more than C5 and catalytic wax oil; the butenes and the stream containing the olefins above C5 are respectively introduced into the first catalytic conversion 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. Therefore, the method disclosed by the invention refines the low-added-value olefin produced in the chemical process in a specific route, so that the yields of ethylene and propylene are effectively improved, and the effective utilization of petroleum resources is realized.

In a preferred embodiment of the present disclosure, in step S3, the butene is contacted with a catalytic conversion catalyst prior to a C5 or higher olefin-containing stream that is co-fed in admixture with the hydrocarbon oil feedstock. 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 material flow containing olefins with 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 olefins, and the efficient utilization of resources is realized.

According to the present disclosure, the content of olefins in the C5 or higher olefin containing stream may be above 50 wt%.

In a preferred embodiment of the present disclosure, the method may further include: the first spent catalyst and the second spent catalyst are sent to a shared regenerator for burning regeneration, and a regenerated catalyst is obtained; the regenerated catalyst is returned to the first catalytic conversion reactor and the second catalytic conversion reactor. The method uses the common regenerator to burn and regenerate the first spent catalyst and the second spent catalyst, and when the heavy raw oil is cracked on the acid catalyst, a large amount of coke with higher graphitization degree is inevitably generated, so that heat brought by burning the coke can be used for providing energy for high-temperature cracking of olefin, and further petroleum resources are efficiently utilized.

In the present disclosure, the first catalytic conversion reactor and the second catalytic conversion reactor are each independently selected from one or two of a riser, an equal linear velocity fluidized bed, an equal diameter fluidized bed, an upstream conveying line and a downstream conveying line, which are in series combination, wherein the riser is an equal diameter riser reactor or a variable radial fluidized bed reactor.

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; more 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.

according to the present disclosure, the reaction conditions under which the butene is introduced into the first catalytic conversion reactor to continue the reaction may 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.

further preferably, the reaction conditions under which the butene is introduced into the first catalytic conversion reactor to continue the reaction may include: 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.

as a preferred embodiment of the present disclosure, the method may further include: and (3) hydrotreating the catalytic wax oil to obtain hydrogenation catalytic wax oil, and mixing the hydrogenation catalytic wax oil with the heavy raw oil and then jointly entering the second catalytic conversion reactor. The method disclosed by the invention carries out hydrogenation treatment on the catalytic wax oil and then continues the reaction, so that the side reaction for generating micromolecular alkane and coke is further reduced, the yield of ethylene and propylene is improved, and the effective utilization of carbon atoms is realized.

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 . The catalyst for hydrotreatment comprises a carrier, a metal component and an optional additive, wherein the metal component is supported on the carrier, the carrier is alumina and/or amorphous silicon aluminum, the metal component is VIB group metal and/or VIII group metal, and the additive is at least one of fluorine, phosphorus, titanium and platinum. 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 olefin content in the hydrocarbon oil feedstock may be above 80 wt%; preferably, the olefin content in the hydrocarbon oil feedstock is above 90 wt.%; more preferably, the hydrocarbon oil feedstock is a pure olefin feedstock; the olefin in the hydrocarbon oil raw material can come from corresponding alkane raw material dehydrogenation, a fraction above C4 generated by a catalytic cracking device of an oil refinery, a fraction above C4 of a steam cracking device or an olefin-rich fraction above C4 of an MTO byproduct in an ethylene plant, and an olefin-rich fraction above C4 of an MTP byproduct; 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. The paraffinic feedstock may be at least one of light hydrocarbons from naphtha, aromatic raffinate, and other units.

In a further embodiment, the method for preparing olefin by dehydrogenating alkane comprises the steps of carrying out contact reaction on alkane and a dehydrogenation catalyst, wherein the inlet temperature of the reactor is 400-700 ℃, and the volume space velocity of the alkane is 200-5000h ^-1 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 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 can comprise one or more of a large pore molecular sieve, a medium pore molecular sieve and a small pore molecular sieve; as a preferred embodiment of the present disclosure, 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.

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.

In one embodiment of the present disclosure, the small pore molecular sieve may be a SAPO molecular sieve, further, the SAPO molecular sieve may be selected from one or more of SAPO-34, SAPO-11, SAPO-47.

In one embodiment of the present disclosure, the large pore molecular sieve may be selected from one or more of rare earth Y molecular sieves, rare earth hydrogen Y molecular sieves, ultrastable Y molecular sieves, high silicon Y molecular sieves, beta molecular sieves, and other molecular sieves of similar structure.

In one embodiment of the present disclosure, as shown in fig. 1, a pre-lift medium enters from the bottom of the first catalytic conversion reactor a through a line 1, a hydrocarbon oil feedstock with an olefin content of 50% or more is injected into the bottom of the first catalytic conversion reactor a through a line 3 together with atomized steam from a line 2, and moves upward along the first catalytic conversion reactor a with a thermally regenerated catalytic conversion catalyst from a line 17 under the lift action of the pre-lift medium, and reacts. The generated first reactant flow and the first spent catalyst enter a cyclone separator in a settler 8 through an outlet section 7, separation of the first reactant flow and the first spent catalyst is realized, the first reactant flow enters a gas collection chamber 9, and the first spent catalyst fine powder returns to the settler through a dipleg. The first spent catalyst in the settler flows to stripping section 6 where it is contacted with stripping steam from line 5. The oil gas stripped from the first spent catalyst enters the gas collection chamber 9 after passing through the cyclone separator, and the stripped first spent catalyst enters the regenerator 13 through the inclined pipe 12.

The pre-lift medium enters from the bottom of the second catalytic conversion reactor B through a line 21, the thermally regenerated catalytic conversion catalyst from the line 18 is accelerated upward along the second catalytic conversion reactor B by the lift action of the pre-lift medium, the heavy feedstock is injected into the bottom of the second catalytic conversion reactor B through a line 23 together with the atomized steam from the line 22, and the heavy feedstock reacts on the thermally regenerated catalytic conversion catalyst and is accelerated upward. The generated second reactant flow and the second spent catalyst enter a cyclone separator in a settler 28 through an outlet section 27 to realize separation of the second reactant flow and the second spent catalyst, the second reactant flow and the second spent catalyst enter a gas collection chamber 29, and the second spent catalyst fines are returned to the settler from a dipleg. The second spent catalyst in the settler flows to stripping section 26 where it is contacted with stripping steam from line 25. The oil gas stripped from the second spent catalyst enters the gas collection chamber 29 after passing through the cyclone separator, and the stripped second spent catalyst enters the regenerator 13 through the inclined pipe 19.

In the regenerator 13, main wind enters the regenerator through a pipeline 14 to burn coke on the first spent catalyst and the second spent catalyst, so that the deactivated first spent catalyst and the deactivated second spent catalyst are regenerated. The flue gas enters the extractor through line 16. Regenerated catalyst is fed via

lines

17 and 18 to the risers of the first catalytic conversion reactor a and the second catalytic conversion reactor B, respectively.

The first reactant flow and the second reactant flow enter a subsequent product separation device 31 through a large oil gas pipeline 10 and a pipeline 30 respectively, hydrogen, methane and ethane obtained through separation are led out through a pipeline 32, ethylene is led out through a pipeline 33, propylene is led out through a pipeline 34, butylene is led out through a pipeline 35 to the bottom of a first catalytic conversion reactor A for continuous reaction, propane and butane are led out through a pipeline 36, catalytic wax oil is led out through a pipeline 38 to a hydrotreating reactor 15, light components after hydrotreating are led out through a pipeline 39, the catalytic wax oil is led out through a pipeline 20 to the bottom of a second catalytic conversion reactor B for continuous reaction, an olefin-containing crude material is led into an olefin separation device 40 through a pipeline 37, a hydrocarbon flow obtained through separation is led out through a pipeline 41, and an olefin-rich material flow is led into the first catalytic conversion reactor A for continuous reaction through a pipeline 24.

The present disclosure is further illustrated in detail by the following examples. The starting materials used in the examples are all available commercially. The raw materials a and b used in the examples are heavy raw oil, and the properties are shown in Table 1; feedstock c is a catalytically cracked gasoline light fraction, the composition of which is shown in table 2.

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%, 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 is kept at 2-4, the temperature is lowered to normal temperature, and 5000 g of slurry prepared in advance is added, wherein 1600g of medium pore ZSM-5 molecular sieve and macroporous 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, washing free Na ⁺ Obtaining the catalyst. The resulting catalyst was aged at 800 ℃ and 100% steam, and the aged catalyst was designated as catalyst a, and the properties of catalyst a are shown in table 3.

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

The preparation of hydrotreating catalyst D used in the examples is briefly described as follows: weighing ammonium metatungstate ((NH) ₄ ) ₂ W ₄ O ₁₃ ·18H ₂ O, chemically pure) and nickel nitrate (Ni (NO) ₃ ) ₂ ·18H ₂ O, chemically pure) was 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, and the immersed solution was treated with ultrasonic waves for 30 minutes during the immersing, cooled, filtered, and dried in a microwave oven for about 15 minutes. The composition of the catalyst is as follows: 30.0 wt% WO ₃ 3.l wt% NiO and the balance alumina are designated as hydrotreating catalyst D.

The hydrodesulfurization catalyst E used in the examples was prepared as follows: 1000 g of pseudo-boehmite produced by China petrochemical catalyst, namely Changling Co., ltd., is weighed, 1000 ml of aqueous solution containing 10 ml 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, thus obtaining 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 hydrodesulfurization catalyst E.

TABLE 1-1

Feed Properties	a
		Density (20 ℃ C.)/(kg/m) ³ )	859.7
Kangshi carbon residue, weight percent	0.07
		C, wt%	85.63
H, wt%	13.45
		S, weight percent	0.077
N, wt%	0.058
		Fe, microgram/gram	2.3
Na, micrograms/gram	0.6
		Ni, microgram/gram	4.9
V, micrograms/gram	0.4
		Group composition, weight percent
Saturated hydrocarbons	58.1
		Aromatic hydrocarbons	26.3
Colloid	15.3
		Asphaltenes	0.3

TABLE 1-2

TABLE 2

Carbon number	Alkane weight percent	Olefin weight%	Naphthene weight%	Aromatic hydrocarbon weight%
					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

Catalyst	A	B	C
				Chemical composition/wt%
Al ₂ O ₃	49.2	46.3	26.5
				Na ₂ O	0.07	0.04	0.19
Physical Properties
				Specific surface area/(m) ² ·g ^-1 )	/	153	132
Bulk density/(g cm) ^-3 )	0.79	0.86	0.45
				Wear index/(% & h ^-1 )	1.1	1	4.2
Sieving composition/wt%
				0-40μm	14.2	17.9	7.3
40-80μm	53.8	41.4	43.7
				＞80μm	320	40.7	49.0

Example 1

The process flow of this example is shown in FIG. 1 and is tested on a medium-sized apparatus in a riser reactor. The 1-pentene and the high-temperature catalytic conversion catalyst A are contacted at the bottom of the first catalytic conversion reactor, the reaction pressure is 0.1MPa at the reaction temperature of 700 ℃, the reaction time is 5s, and the weight ratio of the catalytic conversion catalyst A to the raw materials is 45:1, the heavy raw oil A and the catalytic conversion catalyst A are contacted at the bottom of a second catalytic conversion reactor, the reaction temperature is 530 ℃, the reaction pressure is 0.1MPa, the reaction time is 6s, and the weight ratio of the catalytic conversion catalyst A 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 (with the final distillation point of 250 ℃) and catalytic cracking wax oil (with the initial distillation point of 250 ℃), and the like. Catalytic cracking wax oil and hydrogenation catalyst D at a temperature of 350 ℃, a hydrogen partial pressure of 18MPa and a volume space velocity of 1.5 hours ^-1 And (3) reacting under the condition of hydrogen oil volume ratio of 1500 to obtain the hydrogenation catalytic cracking wax oil. Introducing the obtained butene into the bottom of a first catalytic conversion reactor for recracking, wherein the reaction temperature is 740 ℃, and the weight ratio of the catalytic conversion catalyst to the butene is 100:1 reaction time0.2s; introducing the olefin-rich stream into the bottom of the first catalytic conversion reactor for recracked, wherein the reaction temperature is 700 ℃, and the reaction time is 5s; the hydrocracked wax oil is mixed with heavy raw oil and then returned to the second catalytic conversion reactor 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 riser reactor using a similar process scheme to example 1, except that the olefin-rich stream was not returned to the first catalytic conversion reactor for continued reaction. The heavy raw oil a and the catalytic conversion catalyst A are contacted at the bottom of the second catalytic conversion reactor, the reaction pressure is 0.1MPa at the reaction temperature of 530 ℃, 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. Catalytic cracking wax oil and hydrogenation catalyst D at a temperature of 350 ℃, a hydrogen partial pressure of 18MPa and a volume space velocity of 1.5 hours ^-1 And (3) reacting under the condition of hydrogen oil volume ratio of 1500 to obtain the hydrogenation catalytic cracking wax oil. Introducing the obtained butene into the bottom of a first catalytic conversion reactor for recracking, wherein the reaction temperature is 740 ℃, and the weight ratio of the catalytic conversion catalyst A to the butene is 100:1, the reaction time is 0.2s; the hydrocracked wax oil is mixed with heavy raw oil and then returned to the second catalytic conversion reactor for reaction. The reaction conditions and product distribution are shown in Table 4.

Example 2

The process flow of this example is shown in FIG. 1 and is tested on a medium-sized apparatus in a riser reactor. The heavy raw oil a and the catalytic conversion catalyst A are contacted at the bottom of the second catalytic conversion reactor, the reaction pressure is 0.1MPa at the reaction temperature of 530 ℃, 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 for burning regeneration, and reactingThe reaction products are introduced together into a combined separation system to produce a product comprising ethylene, propylene, butene, an olefin-rich stream, and catalytically cracked wax, and the like. Catalytic cracking wax oil and hydrogenation catalyst D at 350 ℃, hydrogen partial pressure of 18MPa and volume space velocity of 1.5 hours ^-1 And (3) reacting under the condition of hydrogen oil volume ratio of 1500 to obtain hydrogenated wax oil. Introducing the obtained butene into the bottom of a first catalytic conversion reactor 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 stream into the bottom of the first catalytic conversion reactor for recracked, wherein the reaction temperature is 700 ℃, and the reaction time is 5s; the hydrocracked wax oil is mixed with heavy raw oil and then returned to the second catalytic conversion reactor 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 produces ethylene and propylene as in example 2, except that: raw material b was used as raw material. The catalytic cracking wax oil of this example was not subjected to deep hydrotreatment, contacted with hydrodesulfurization catalyst E in a hydrodesulfurization reactor at a reaction pressure of 6.0MPa, a reaction temperature of 350℃and a hydrogen-to-oil volume ratio of 350, a volume space velocity of 2.0 hours ^-1 And (3) carrying out reaction to obtain low-sulfur hydrogenated distillate oil serving 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 in a riser reactor, with feedstock b and catalytic conversion catalyst C in contact at the bottom of the riser, at a reaction temperature of 530 ℃, the weight ratio of catalyst to feedstock being 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 produces ethylene and propylene as in example 1, except that: the first catalytic conversion reactor is fed with light fraction C of catalytic cracking gasoline. The reaction conditions and product distribution are shown in Table 4.

Example 5

The method and apparatus of example 1 were used, with the only differences: the reaction conditions in each reaction zone were varied, and the specific 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, the higher yields of ethylene and propylene were obtained in the high temperature cracking of the olefins in examples 1-4, the higher the ethylene and propylene contents in the resulting reactant streams were 60% or more, and the more effective the feed olefin content was, the more effective the ethylene content was 26.96% and the propylene content was 35.57% when 1-pentene having an olefin content of 100% was used as the feed, and the sum of both could be as high as 62.53%. In addition, benzene, toluene and xylene yields are significantly increased.

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 ethylene and propylene, the process comprising the steps of:

s3, separating the first reactant flow and the second reactant flow to obtain ethylene, propylene, butylene, a flow containing olefins with more than C5 and catalytic wax oil; introducing the butenes and the stream containing olefins having more than 5 carbon atoms into the first catalytic conversion reactor respectively to continue the reaction; wherein the butenes are contacted with the catalytic conversion catalyst prior to the C5 or greater olefin-containing stream, the C5 or greater olefin-containing stream being co-fed in admixture with the hydrocarbon oil feedstock;

the content of olefin in the stream containing C5 or higher olefin is more than 50 wt%;

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 reaction conditions under which the butene is introduced into the first catalytic conversion reactor to continue the reaction 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 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 olefin in the hydrocarbon oil raw material is selected 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 raw material is at least one of naphtha, aromatic raffinate oil and light hydrocarbon;

the catalytic conversion catalyst comprises 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 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.

2. The method of claim 1, wherein the method further comprises: the first spent catalyst and the second spent catalyst are sent to a shared regenerator for burning regeneration, and a regenerated catalyst is obtained;

the regenerated catalyst is returned to the first catalytic conversion reactor and the second catalytic conversion reactor.

3. The method of claim 1, wherein the reaction conditions under which the butene is introduced into the first catalytic conversion reactor to continue the reaction 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) hydrotreating the catalytic wax oil to obtain hydrogenation catalytic wax oil, and mixing the hydrogenation catalytic wax oil with the heavy raw oil and then jointly entering the second catalytic conversion reactor.

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 olefin content in the hydrocarbon oil feedstock is 80 wt% or more.

7. The method of claim 6, wherein the olefin content in the hydrocarbon oil feedstock is above 90 wt%.

8. The method of claim 7, wherein the hydrocarbon oil feedstock is a pure olefin feedstock.