CN114763485A

CN114763485A - Catalytic conversion method for preparing ethylene and propylene

Info

Publication number: CN114763485A
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: 2022-07-19
Anticipated expiration: 2041-01-11
Also published as: CN114763485B

Abstract

The present disclosure relates to a catalytic conversion process for producing ethylene and propylene, the process comprising: s1, contacting a hydrocarbon oil raw material with the olefin content of more than 50 wt% with a catalytic conversion catalyst with the temperature of more than 650 ℃, and carrying out catalytic conversion reaction in a first reactor to obtain a first reaction material flow and a first catalyst to be generated; s2, contacting the heavy raw oil with a catalytic conversion catalyst at a temperature of above 650 ℃, and carrying out a 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 separately introduced into the first reactor to continue the 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 processes, in particular to a catalytic conversion method for preparing ethylene and propylene.

Background

Ethylene is one of the most chemical products in the world and accounts for more than 75% of the whole petrochemical product yield in the world; the bulk downstream products of ethylene mainly include 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 in demand as important chemical intermediates.

The traditional route of preparing ethylene and propylene by steam cracking is adopted, and the demand for chemical light hydrocarbons such as light hydrocarbon, naphtha and the like is large. Research institutions expect that the global annual average growth rate of gasoline pool will be less than 1% from 2018 to 2026, but propylene will increase by about 4%. The high-carbon olefin in the refinery process is reasonably utilized to prepare ethylene and propylene by cracking, thereby not only meeting the goals of quality improvement and efficiency improvement of petrochemical enterprises, but also complying with the time requirement 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 the reaction temperature of 400-600 ℃ and the absolute pressure of 0.3-1.1KPa, and circulating 30-90 wt% of C4 fraction into a reactor through a separation device for cracking again. The method mainly improves the conversion rate of olefin by circulating C4 fraction, the obtained ethylene and propylene are not less than 62% of the total amount of the raw material olefin, but the ethylene/propylene ratio is relatively low, the adjustment cannot be flexibly carried out according to market demands, the reaction selectivity is low, the content of butylene in the product is high, and the problems of C4 separation energy consumption and the like exist.

Chinese patent CN101239878A discloses a method for preparing olefin with four or more carbon atomsThe olefin-rich mixture is used as a raw material, the reaction temperature is 400-680 ℃, the reaction pressure is-0.09-1.0 MPa, and the weight space velocity is 0.1-50 hours^-1The 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.

Therefore, there is a need in the art for a new method for producing ethylene and propylene in high yield to achieve efficient utilization of petroleum resources.

Disclosure of Invention

It is an object of the present disclosure to provide a catalytic conversion process for producing ethylene and propylene that further increases the ethylene and propylene yields.

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

s1, contacting a hydrocarbon oil raw material with the olefin content of more than 50 wt% with a catalytic conversion catalyst at the temperature of more than 650 ℃, and carrying out catalytic conversion reaction in a first catalytic conversion reactor to obtain a first reactant flow and a first catalyst to be generated;

s2, contacting the heavy raw oil with a catalytic conversion catalyst with the temperature of above 650 ℃, and carrying out catalytic conversion reaction in a second catalytic conversion reactor to obtain a second reaction material 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; introducing the butylene and the stream containing the olefin above C5 into the first catalytic conversion reactor respectively to continue the reaction.

Alternatively, in step S3, the butene is contacted with the catalytic conversion catalyst before the stream containing C5 or more, and the stream containing C5 or more olefins is mixed with the hydrocarbon oil feedstock and co-fed.

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

Optionally, the method further comprises: sending the first spent catalyst and the second spent catalyst to a common regenerator for scorching regeneration to obtain regenerated catalysts; returning the regenerated catalyst 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; 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 butenes are 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 butylene is (20-200): 1.

preferably, the reaction conditions under which the butenes are introduced into the first catalytic conversion reactor to continue the 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 butylene is (30-180): 1.

optionally, the method further comprises: and (3) carrying out hydrotreating on the catalytic wax oil to obtain hydrogenated catalytic wax oil, and mixing the hydrogenated catalytic wax oil with the heavy raw material oil and then jointly entering the second catalytic conversion reactor.

Optionally, the hydrotreating conditions include: the hydrogen partial pressure is 3.0-20.0 MPa, the reaction temperature is 300-450 ℃, the hydrogen-oil volume ratio is 300-2000, and the volume space velocity 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 90 wt% or more; more preferably, the hydrocarbon oil feedstock is a pure olefin feedstock; the heavy raw oil is selected from petroleum hydrocarbon and/or mineral oil; the petroleum hydrocarbon is at least one of vacuum gas oil, atmospheric gas oil, coking gas oil, deasphalted oil, vacuum residue, atmospheric residue and heavy aromatic raffinate oil; the mineral oil is selected from at least one of coal liquefaction oil, oil sand oil and shale oil.

Optionally, the olefins in the hydrocarbon oil feedstock are derived from a fraction above C4 produced by dehydrogenation of an alkane feedstock, a fraction above C4 produced by a catalytic cracking unit in an oil refinery, a fraction above C4 produced by a steam cracking unit in an ethylene plant, an olefin-rich fraction above C4 produced as a by-product of MTO, and an olefin-rich fraction above C4 produced as a by-product of MTP.

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

Optionally, 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 a 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% 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 method disclosed by the invention respectively carries out pyrolysis on the raw material containing the olefin and the heavy raw oil, and returns the olefin in the separated product to the device for continuous reaction. By recycling macromolecular olefin produced in the oil-gas separation process by a specific route, the methane yield is reduced, so that the aim of efficiently utilizing petroleum resources is fulfilled, the traditional scheme for producing ethylene and propylene by high-energy-consumption steam cracking can be replaced, and meanwhile, the method disclosed by the invention has high ethylene and propylene yield, selectivity and low methane yield. In addition, the method uses the common regenerator to coke and regenerate the first catalyst to be regenerated and then respectively send the coke to the first catalytic conversion reactor and the second catalytic conversion reactor, a large amount of coke with higher graphitization degree is inevitably generated when the heavy raw oil is cracked on the acid catalyst, and heat brought by the combustion of the coke can provide energy for the high-temperature cracking of olefin, thereby realizing the recycling of the catalytic conversion catalyst, further improving the utilization rate of petroleum resources and simultaneously reducing the energy consumption of the device.

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

Drawings

The accompanying drawings, which are included to provide a further understanding of the disclosure and are incorporated in and constitute a part of this specification, illustrate embodiments of the disclosure and together with the description serve to explain the disclosure without limiting the disclosure. In the drawings:

fig. 1 is a schematic flow diagram of one embodiment of the present disclosure.

Description of the reference numerals

A first reactor B second reactor

1 pipeline 2 pipeline 3 pipeline

5 line 6 stripping section 7 outlet section

8 settler 9 gas collection chamber 10 pipeline

12 inclined pipe 13 regenerator 14 pipeline

15 hydrotreating reactor 16 line 17 line

18 pipeline 19 pipe 24 pipeline

21 line 22 line 23 line

24 line 25 line 26 stripping section

27 outlet section 28 settler 29 gas collection chamber

30 line 31 product separation unit 32 line

33 line 34 line 35 line

36 line 37 line 38 line

39 line 40 olefin separation unit 41 line

Detailed Description

The following describes in detail specific embodiments of the present disclosure. It should be understood that the detailed description and specific examples, while indicating the present disclosure, are given by way of illustration and explanation only, not limitation.

The present disclosure provides a catalytic conversion method for producing ethylene and propylene in high yield, comprising the steps of:

s1, contacting a hydrocarbon oil raw material with the olefin content of more than 50 wt% with a catalytic conversion catalyst with the temperature of more than 650 ℃, and carrying out catalytic conversion reaction in a first catalytic conversion reactor to obtain a first reaction material flow and a first catalyst to be generated;

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 butene and the stream containing olefins above C5 into the first catalytic conversion reactor separately for further reaction.

The inventors of the present disclosure have found through extensive experiments that there are differences in the distribution of products formed by the reaction of alkanes and alkenes over the catalyst. When olefins are catalyzed on a high-temperature catalyst, the yields of hydrogen, methane and ethane in the products are low, the yields of ethylene and propylene are high, 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 by a specific route, effectively improves the yield of ethylene and propylene, and realizes the effective utilization of petroleum resources.

In a preferred embodiment of the present disclosure, in step S3, the butene is contacted with the catalytic conversion catalyst before the stream containing olefins above C5 is contacted with the catalytic conversion catalyst, and the stream containing olefins above C5 is co-fed in admixture with the hydrocarbon oil feedstock. The difficulty of cracking the hydrocarbon is increased along with the reduction of carbon number, and the energy required by cracking the butylene is higher, so if the butylene is preferably contacted with a high-temperature catalytic conversion catalyst firstly, and then the material flow containing the olefin above C5 is contacted with the catalytic conversion catalyst, the butylene conversion rate and the selectivity of the product ethylene and propylene can be improved, the phenomenon that the olefin is fed simultaneously to generate more byproducts is avoided, and the high-efficiency utilization of resources is realized.

According to the present disclosure, the olefin content of the C5 or greater olefin containing stream may be 50 wt% or greater.

In a preferred embodiment of the present disclosure, the method may further include: sending the first spent catalyst and the second spent catalyst to a common regenerator for scorching regeneration to obtain regenerated catalysts; returning the regenerated catalyst to the first catalytic conversion reactor and the second catalytic conversion reactor. The method uses the common regenerator to coke and regenerate the first spent catalyst and the second spent catalyst, a large amount of coke with higher graphitization degree is inevitably generated when the heavy raw oil is cracked on the acid catalyst, and the heat brought by the combustion of the coke can be used for providing energy for the high-temperature cracking of olefin, thereby further efficiently utilizing petroleum resources.

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, a fluidized bed with equal linear speed, a fluidized bed with equal diameter, an ascending conveyor line and a descending conveyor line, which are connected in series, wherein the riser is an equal-diameter riser reactor or a variable-diameter 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; 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; 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; 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; 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 butenes are 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 butylene is (20-200): 1.

further preferably, the reaction conditions under which the butenes are 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 butylene is (30-180): 1.

as a preferred embodiment of the present disclosure, the method may further include: and (3) carrying out hydrotreating on the catalytic wax oil to obtain hydrogenated catalytic wax oil, mixing the hydrogenated catalytic wax oil with the heavy raw material oil, and then jointly entering the second catalytic conversion reactor. The method disclosed by the invention has the advantages that the catalytic wax oil is subjected to hydrotreating and then continuously reacts, so that the side reaction of generating small-molecular 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 conditions of the hydrotreating may include: the hydrogen partial pressure is 3.0-20.0 MPa, the reaction temperature is 300-450 ℃, the hydrogen-oil volume ratio is 300-2000, and the volume space velocity is 0.1-3.0 hours^-1. The catalyst used for the hydrotreatment comprises a carrier and a catalyst supported on the carrierThe 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 selected from 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; based on the weight of the hydrotreating catalyst, the additive is 0-10 wt%, the group VIB metal is 12-39 wt%, and the group VIII metal is 1-9 wt%.

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 90 wt% or more; more preferably, the hydrocarbon oil feedstock is a pure olefin feedstock; the olefin in the hydrocarbon oil raw material can be from the dehydrogenation of corresponding alkane raw materials, a fraction above C4 produced by a catalytic cracking device in an oil refinery, a fraction above C4 of a steam cracking device in an ethylene plant, an olefin-rich fraction above C4 of an MTO byproduct and an olefin-rich fraction above C4 of the MTP byproduct; the heavy raw 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 liquefaction oil, oil sand oil, and shale oil. The paraffinic feedstock may be derived from at least one of naphtha, aromatic raffinate, and light hydrocarbons from other plants.

In a further embodiment, the method for preparing olefin by alkane dehydrogenation comprises the step of carrying out contact reaction on alkane and a dehydrogenation catalyst, wherein the inlet temperature of a reactor is 400-700 ℃, and the volume space velocity of alkane is 200-5000h^-1The pressure of the contact reaction is 0-1.0 MPa. The dehydrogenation catalyst consists of a carrier, and an active component and an auxiliary agent which are loaded on the carrier; based on the total weight of the catalyst being 100%, 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%; 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 above-mentionedThe active component is platinum or chromium; the auxiliary agent is bismuth and an alkali metal component or bismuth and an alkaline earth metal component; 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 components 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, based on the weight of the catalytic conversion catalyst, from 1 to 50 wt% of a molecular sieve, from 5 to 99 wt% of an inorganic oxide, and from 0 to 70 wt% of a clay; 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, and further, the ZSM molecular sieve may be one or more selected from ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, and ZSM-48.

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

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

The pre-lifting medium enters from the bottom of the second catalytic conversion reactor B through a pipeline 21, the thermal regeneration catalytic conversion catalyst from the pipeline 18 moves upwards in an accelerated manner along the second catalytic conversion reactor B under the lifting action of the pre-lifting medium, the heavy raw oil is injected into the bottom of the second catalytic conversion reactor B through a pipeline 23 together with the atomized steam from a pipeline 22, and the heavy raw oil reacts on the thermal catalytic conversion catalyst and moves upwards in an accelerated manner. 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 the 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 fine powder of the second spent catalyst returns to the settler through a dipleg. The second spent catalyst in the settler flows to stripping section 26 and contacts stripping steam from line 25. The oil gas stripped from the second spent catalyst enters a gas collecting chamber 29 after passing through a cyclone separator, and the stripped second spent catalyst enters the regenerator 13 through the inclined tube 19.

In the regenerator 13, the main air enters the regenerator through a pipeline 14, and coke on the first spent catalyst and the second spent catalyst is burned off, so that the inactivated first spent catalyst and the inactivated second spent catalyst are regenerated. The smoke enters the cigarette machine through line 16. The regenerated catalyst enters the riser of the first catalytic conversion reactor A and the riser of the second catalytic conversion reactor B through the

lines

17 and 18, respectively.

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

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

The preparation of the catalytic conversion catalyst A used in the examples is briefly as follows:

using 4300 g of decationized water to pulp 969 g of halloysite (a product of China Kaolin company, with the solid content of 73 percent), adding 781 g of pseudoboehmite (a product of Shandong Zibostone factory, with the solid content of 64 percent) and 144 ml of hydrochloric acid (with the concentration of 30 percent and the specific gravity of 1.56) to be uniformly stirred, standing and aging for 1 hour at 60 ℃, keeping the pH value at 2-4, cooling to the normal temperature, and adding 5000 g of prepared slurry, wherein 1600g of a medium-pore ZSM-5 molecular sieve and a macroporous Y-type molecular sieve (produced by China petrochemical catalyst Qilu division) are added, and the weight ratio of the medium-pore ZSM-5 molecular sieve to the macroporous Y-type molecular sieve is 9: 1. stirring, spray drying, and washing off free Na⁺And obtaining the catalyst. The resulting catalyst was aged at 800 ℃ and 100% steam, the aged catalyst was designated as catalytic conversion catalyst a, the properties of which are shown in table 3.

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

Used in the examplesThe preparation method of the hydrotreating catalyst D is briefly as follows: ammonium metatungstate ((NH) was weighed₄)₂W₄O₁₃·18H₂O, chemically pure) and nickel nitrate (Ni (NO)₃)₂·18H₂O, chemically pure) was made up with 200 ml of water. The solution was added to 50 g of alumina support, immersed at room temperature for 3 hours, the immersion liquid was treated with ultrasonic waves for 30 minutes during the immersion, cooled, filtered, and dried in a microwave oven for about 15 minutes. The catalyst comprises the following components: 30.0% by weight of WO₃3.l wt% NiO and balance alumina, designated hydrotreating catalyst D.

The preparation of the hydrodesulfurization catalyst E used in the examples is as follows: weighing 1000 g of pseudo-boehmite produced by China petrochemical catalyst ChangLing division, adding 1000 ml of aqueous solution containing 10 ml of nitric acid (chemical purity), extruding and molding on a double-screw extruder, drying at 120 ℃ for 4 hours, and roasting at 800 ℃ for 4 hours to obtain the catalyst carrier. Dipping for 2 hours by 900 ml of aqueous solution containing 120 g of ammonium fluoride, drying for 3 hours at 120 ℃, and roasting for 3 hours at 600 ℃; cooling to room temperature, soaking in 950 ml of water solution containing 133 g of ammonium metatolybdate for 3 hours, drying at 120 ℃ for 3 hours, roasting at 600 ℃ for 3 hours, cooling to room temperature, soaking in 900 ml of water solution containing 180 g of nickel nitrate and 320 g of ammonium metatungstate for 4 hours, soaking the fluoridated alumina carrier in 0.1 wt% of ammonium metatolybdate (chemical purity) and 0.1 wt% of nickel nitrate (chemical purity) in the mixed water solution of the catalyst carrier for 4 hours, drying at 120 ℃ for 3 hours, and roasting at 600 ℃ for 4 hours to obtain the hydrodesulfurization catalyst E.

TABLE 1-1

Nature of the feed	a
		Density (20 ℃ C.)/(kg/m)³)	859.7
Kang's carbon residue, wt.%	0.07
		C, weight%	85.63
H, weight%	13.45
		S, wt.%	0.077
N, weight%	0.058
		Fe in microgram/g	2.3
Na in microgram/g	0.6
		Ni in microgram/gram	4.9
V, microgram/gram	0.4
		Group composition, weight%
Saturated hydrocarbons	58.1
		Aromatic hydrocarbons	26.3
Glue	15.3
		Asphaltenes	0.3

Tables 1 to 2

TABLE 2

Carbon number	Alkane weight%	Olefin weight%	Weight% of cycloalkanes	Weight% of aromatic hydrocarbons
					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

Catalyst and process for preparing same	A	B	C
				Chemical composition/weight%
Al₂O₃	49.2	46.3	26.5
				Na₂O	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
				Abrasion index/(%. h)^-1)	1.1	1	4.2
Screening composition/weight%
				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 was as shown in FIG. 1 and was tested on a pilot plant of a riser reactor. 1-pentene and a high-temperature catalytic conversion catalyst A are contacted at the bottom of a first catalytic conversion reactor, the reaction temperature is 700 ℃, the reaction pressure is 0.1MPa, the reaction time is 5s, and the weight ratio of the catalytic conversion catalyst A to raw materials is 45: 1, catalytic conversion reaction is carried out, the heavy raw oil A and the catalytic conversion catalyst A are contacted at the bottom of a second catalytic conversion reactor, and catalytic conversion is carried out at the reaction temperature of 530 ℃, the reaction pressure of 0.1MPa and the reaction time of 6sThe weight ratio of the catalyst A to the raw materials is 5: 1, catalytic conversion reaction is carried out, and the ratio of 1-pentene to heavy raw materials is 1: 9. separating reaction products and spent catalysts of the reaction, introducing the obtained spent catalysts into a regenerator for coke burning regeneration, and introducing the obtained reaction products into a combined separation system to obtain products comprising ethylene, propylene, butylene, a material flow rich in olefin (the final boiling point is 250 ℃), catalytic cracking wax oil (the initial boiling point is 250 ℃), and the like. The catalytic cracking wax oil and the hydrogenation catalyst D are carried out at the temperature of 350 ℃, the hydrogen partial pressure of 18MPa and the volume space velocity of 1.5 hours^-1And reacting under the condition that the volume ratio of hydrogen to oil is 1500 to obtain the hydrocatalytic cracking wax oil. Introducing the obtained butylene into the bottom of a first catalytic conversion reactor for cracking, wherein the reaction temperature is 740 ℃, and the weight ratio of a catalytic conversion catalyst to the butylene is 100: 1, the reaction time is 0.2 s; introducing the olefin-rich material flow into the bottom of a first catalytic conversion reactor for cracking, wherein the reaction temperature is 700 ℃, and the reaction time is 5 s; and the hydrocatalytic cracking wax oil is mixed with the heavy raw oil and then returns to the second catalytic conversion reactor for reaction. The reaction conditions and product distribution are listed in table 4.

Comparative example 1

This comparative example was run on a pilot plant of a riser reactor using a similar process flow as in example 1, except that the olefin rich stream was not returned to the first catalytic conversion reactor for continued reaction. Contacting the heavy raw oil a with a catalytic conversion catalyst A at the bottom of a second catalytic conversion reactor, wherein 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 material is 5: catalytic conversion reaction takes place at 1. Separating reaction products and spent catalyst of the reaction, introducing the spent catalyst into a regenerator for coke burning regeneration, and introducing the obtained reaction products into a combined separation system to obtain products such as ethylene, propylene, butylene, a material flow rich in olefin, catalytic cracking wax oil and the like. The catalytic cracking wax oil and the hydrogenation catalyst D are carried out at the temperature of 350 ℃, the hydrogen partial pressure of 18MPa and the volume space velocity of 1.5 hours^-1And reacting under the condition that the volume ratio of hydrogen to oil is 1500 to obtain the hydrocatalytic cracking wax oil. Introducing the obtained butene into the bottom of a first catalytic conversion reactor for cracking at 740 DEG CThe weight ratio of the catalytic conversion catalyst A to the butylene is 100: 1, the reaction time is 0.2 s; and the hydrocatalytic cracking wax oil is mixed with the heavy raw oil and then returns to the second catalytic conversion reactor for reaction. The reaction conditions and product distribution are listed in Table 4.

Example 2

The process flow of this example was as shown in FIG. 1 and was tested on a pilot plant of a riser reactor. Contacting the heavy raw oil a with a catalytic conversion catalyst A at the bottom of a second catalytic conversion reactor, wherein 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 material is 5: 1, carrying out catalytic conversion reaction, separating reaction products and spent catalyst of the reaction, introducing the obtained spent catalyst into a regenerator together for scorching regeneration, introducing the obtained reaction products into a combined separation system together, and obtaining products comprising ethylene, propylene, butylene, olefin-rich material flow, catalytic cracking wax oil and the like. The catalytic cracking wax oil and the hydrogenation catalyst D have the hydrogen partial pressure of 18MPa and the volume space velocity of 1.5 hours at the temperature of 350 DEG C^-1And reacting under the condition that the volume ratio of hydrogen to oil is 1500 to obtain hydrogenated wax oil. Introducing the obtained butene into the bottom of a first catalytic conversion reactor for cracking, wherein the reaction temperature is 740 ℃, and the weight ratio of the catalyst to the raw materials is 100: 1, the reaction time is 0.2 s; introducing the olefin-rich material flow into the bottom of a first catalytic conversion reactor for cracking, wherein the reaction temperature is 700 ℃, and the reaction time is 5 s; and the hydrocatalytic cracking wax oil is mixed with the heavy raw oil and then returns to the second catalytic conversion reactor for reaction. The reaction conditions and product distribution are listed in Table 4.

Comparative example 2

The test is carried out on a medium-sized device of a riser reactor, heavy raw oil a and a catalytic conversion catalyst B are contacted at the bottom of the riser, the reaction temperature is 610 ℃, and the weight ratio of the catalyst to the raw material is 16.9: 1, the reaction pressure is 0.1MPa, the catalytic conversion reaction is carried out within 6s of reaction time, and the product is not subjected to hydrotreating and returns to the device for continuous reaction. The reaction conditions and product distribution are listed in Table 4.

Example 3

This example produced ethylene and propylene according to the method of example 2Alkene, except that: the raw material b was used as a raw material. The catalytic cracking wax oil of the embodiment is not subjected to deep hydrogenation treatment, and is contacted with a hydrodesulfurization catalyst E in a hydrodesulfurization reactor at the reaction pressure of 6.0MPa, the reaction temperature of 350 ℃, the hydrogen-oil volume ratio of 350 and the volume airspeed of 2.0 hours^-1Then low-sulfur hydrogenated distillate oil is obtained and is used as a light oil component. The reaction conditions and product distribution are listed in Table 4.

Comparative example 3

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

Example 4

This example produced ethylene and propylene following the procedure of example 1, except that: the first catalytic conversion reactor feed is catalytic cracked gasoline light fraction C. The reaction conditions and product distribution are listed in Table 4.

Example 5

The method and apparatus of example 1 were used, differing only in that: the reaction conditions in each reaction zone were varied and the specific reaction conditions and product distribution are shown in Table 4.

TABLE 4

As can be seen from Table 4, the olefins in examples 1-4 were cracked at high temperature to obtain higher yields of ethylene and propylene, and the resulting reaction streams had ethylene and propylene contents of 60% or more, and the higher the olefin content of the feedstock, the better the results, and when 1-pentene having an olefin content of 100% was used as the feedstock, the ethylene content was 26.96%, the propylene content was 35.57%, and the sum of both contents could be as high as 62.53%. In addition, the benzene, toluene and xylene yields are significantly increased.

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

It should be noted that the various features described in the above embodiments may be combined in any suitable manner without departing from the scope of the invention. In order to avoid unnecessary repetition, various possible combinations will not be separately described in this disclosure.

In addition, any combination of various embodiments of the present disclosure may be made, and the same should be considered as the disclosure of the present disclosure, as long as it does not depart from the spirit 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 olefin with more than C5 and catalytic wax oil; introducing the butene and the stream containing olefins above C5 into the first catalytic conversion reactor separately for further reaction.

2. The process of claim 1, wherein in step S3, the butene is contacted with the catalytic conversion catalyst before the stream containing C5 or more olefins is contacted, and the stream containing C5 or more olefins is co-fed in admixture with the hydrocarbon oil feedstock.

3. The process of claim 1 wherein the olefin content of the C5 or greater olefin containing stream is 50 wt.% or greater.

4. The method of claim 1, wherein the method further comprises: sending the first spent catalyst and the second spent catalyst to a common regenerator for scorching regeneration to obtain a regenerated catalyst;

returning the regenerated catalyst to the first catalytic conversion reactor and the second catalytic conversion reactor.

5. The method of claim 1, wherein,

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;

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.

6. the method of claim 1, wherein,

the reaction conditions under which the butenes are 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 butylene is (20-200): 1;

7. the method of claim 1, wherein the method further comprises: and (3) carrying out hydrotreating on the catalytic wax oil to obtain hydrogenated catalytic wax oil, mixing the hydrogenated catalytic wax oil with the heavy raw material oil, and then jointly entering the second catalytic conversion reactor.

8. The method of claim 1, wherein,

the hydrotreating conditions include: the hydrogen partial pressure is 3.0-20.0 MPa, the reaction temperature is 300-450 ℃, the hydrogen-oil volume ratio is 300-2000, and the volume space velocity is 0.1-3.0 hours^-1。

9. The method according to claim 1, wherein the content of olefins in the hydrocarbon oil feedstock is 80% by weight 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 raw oil is selected from petroleum hydrocarbon and/or mineral oil; the petroleum hydrocarbon is at least one of vacuum gas oil, atmospheric gas oil, coking gas oil, deasphalted oil, vacuum residue, atmospheric residue and heavy aromatic raffinate oil; the mineral oil is selected from at least one of coal liquefaction oil, oil sand oil and shale oil.

10. The process according to claim 9, wherein the olefins in the hydrocarbon oil feedstock are selected from a fraction of greater than or equal to C4 produced by dehydrogenation of an alkane feedstock, a fraction of greater than or equal to C4 produced by a catalytic cracking unit in an oil refinery, a fraction of greater than or equal to C4 produced by a steam cracking unit in an ethylene plant, an olefin-rich fraction of greater than or equal to C4 produced as an MTO by-product, and an olefin-rich fraction of greater than or equal to C4 produced as an MTP by-product;

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

11. The process of claim 1, wherein 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 a 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.