CN112552957B

CN112552957B - Method for recycling and cracking hydrocarbons

Info

Publication number: CN112552957B
Application number: CN201910913743.0A
Authority: CN
Inventors: 朱金泉; 崔琰; 龚剑洪
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: 2019-09-25
Filing date: 2019-09-25
Publication date: 2022-03-11
Anticipated expiration: 2039-09-25
Also published as: CN112552957A

Abstract

A process for the recycle cracking of hydrocarbons is disclosed. The method comprises the following steps: carrying out a first catalytic reaction on a heavy hydrocarbon oil raw material to obtain a first reaction mixture, and separating to obtain a first carbon deposition catalyst and a first product; introducing the first product and the second product into a first product separation system for separation to obtain different fractions (including a second light gasoline fraction); subjecting some of the different fractions to a folding reaction to obtain a folded product; carrying out a second catalytic reaction on the superposed product and a second light gasoline fraction to obtain a second reaction mixture; and carrying out a third catalytic reaction on the second reaction mixture to obtain a third reaction mixture, and separating to obtain a second carbon deposition catalyst and the second product, wherein the distillation range of the superimposed product is 155-330 ℃, and preferably 160-253 ℃. The method disclosed by the invention can greatly reduce the reaction severity and greatly improve the selectivity of propylene cracking of the recycle stream.

Description

Method for recycling and cracking hydrocarbons

Technical Field

The present disclosure relates to a process for the recycle cracking of hydrocarbons.

Background

In the national vehicle gasoline standard GB17930-2016, the technical requirements on vehicle gasoline are that the olefin content is not more than 15% (volume fraction), the aromatic hydrocarbon content is not more than 35% (volume fraction), the benzene content is not more than 0.8% (volume fraction), and the oxygen content is not more than 2.7% (mass fraction). According to the requirements of the national institute of development and improvement, the vehicle ethanol gasoline is popularized and used nationwide in 2020, and the vehicle gasoline market is expected to be saturated according to the current popularization of E10 ethanol gasoline (containing 10% ethanol). While MTBE will be limited to motor gasoline, a new outlet needs to be found for the C4 olefin feed for MTBE.

Modern petroleum processing technology seeks to increase the yield of high value products (e.g., ethylene, propylene, C8 aromatics), and also focuses more on reducing unit feedstock processing energy consumption and reducing carbon emissions. The catalytic cracking of C4 olefin can produce more ethylene and propylene. For example, in journal of petroleum refining and chemical engineering, volume 36, 2005, 2, research on the preparation of propylene and ethylene by catalytic cracking of butene, reported that ethylene and propylene can be produced from olefin C4 by catalytic cracking.

There are also patent publications on the production of ethylene and propylene by polymerization and catalytic cracking of C4 or liquefied gas as raw material and catalysts thereof.

Patent CN102531824B discloses a process for preparing propylene and ethylene from liquefied gas containing butene: (1) the liquefied gas and the reacted oil-gas mixture exchange heat and/or are directly heated to reach the preheating temperature of 150 ℃ and 450 ℃; (2) liquefied gas enters the upper section of the reactor from the top of the reactor, olefin components in the liquefied gas are subjected to a polymerization reaction to generate macromolecular hydrocarbon under the action of a polymerization catalyst, and the temperature of oil gas is increased due to the reaction heat release; (3) the oil-gas mixture generated by the superposition reaction enters the lower section of the reactor, the cracking reaction is carried out under the action of a cracking catalyst to generate a hydrocarbon mixture containing target products of propylene and ethylene, the hydrocarbon mixture flows out from the bottom of the reactor and enters a subsequent separation system, and the propylene, the ethylene and the aromatic oil generated by the reaction are recovered. The invention adopts two or more catalyst bed layers in the process technology of preparing propylene and ethylene by liquefied gas, so that the raw material liquefied gas is contacted with two or more catalysts in sequence, and the superposition reaction and the cracking reaction occur in sequence, thereby improving the conversion rate of the raw material liquefied gas and the selectivity of propylene and ethylene. However, the upper section and the lower section of the reactor arranged in the technology are arranged in series, the reaction conditions of the upper section and the lower section can interfere with each other, the superposition reaction requires high pressure and low temperature, and the cracking reaction requires low pressure and high temperature. Meanwhile, the liquefied gas containing C4 is difficult to be converted completely when the polymerization reaction is carried out in one time when the liquefied gas passes through the upper section of the reactor.

Patent CN100537721C discloses a catalytic conversion method for increasing propylene yield, in which preheated raw oil is injected into a main riser of a double-riser reaction regeneration system, and contacts with a hot catalyst to perform catalytic cracking reaction, and a reaction product is separated, wherein a spent catalyst for coke generation is recycled after stripping and regeneration; injecting the liquefied gas product from which propylene is separated by the gas separation system into an auxiliary riser, contacting with a hot catalyst, sequentially carrying out superposition reaction, catalytic cracking and alkane dehydrogenation reaction in two reaction zones in the auxiliary riser, separating reaction products, and recycling the regenerated catalyst. The method provided by the invention is adopted to further convert the liquefied gas product after the propylene removal into propylene, and the propylene yield is obviously improved on the premise of not increasing the yield of the liquefied gas. The auxiliary riser is sequentially provided with a first reaction zone, a second reaction zone, an outlet zone and a horizontal pipe which are coaxial with each other from bottom to top, the horizontal pipe is connected with the settler, and the lower parts of the first reaction zone and the second reaction zone are respectively connected with a catalyst inlet pipe; the operating conditions were: the reaction temperature of the main riser reactor is 450-650 ℃; the weight ratio of the catalyst to the raw oil is 1-25, the reaction time is 0.5-30 seconds, and the pressure (absolute pressure) in the main riser reactor is 0.1-0.4 MPa; the temperature of the first reaction zone of the auxiliary riser reactor is 150-450 ℃, the reaction time is 0.5-2.0 seconds, the weight ratio of the catalyst to the feed gas is 1-30, the temperature of the second reaction zone is 450-650 ℃, the reaction time is 3-20 seconds, the weight ratio of the catalyst to the feed gas is 3-60, and the pressure (absolute pressure) in the auxiliary riser reactor is 0.1-0.4 MPa.

Also, patent CN100448954C discloses a catalytic conversion method for increasing propylene yield, in which preheated raw oil is injected into a main riser of a double-riser reaction regeneration system, and contacts with a hot catalyst to perform catalytic cracking reaction, and the reaction product is separated, and the catalyst to be regenerated is recycled; injecting the liquefied gas product with the separated propylene into an auxiliary lifting pipe, contacting with a hot catalyst, sequentially carrying out olefin polymerization, polymerization product cracking and alkane dehydrogenation reaction, separating reaction products, and recycling the regenerated catalyst; the catalyst is a mixture of two catalysts: the cracking catalyst comprises a first cracking catalyst containing a Y-type molecular sieve and a second cracking catalyst containing a ZSM-5 molecular sieve, a transition metal additive and a phosphorus additive, wherein the dry basis weight ratio of the first cracking catalyst to the second cracking catalyst is 10-70: 30-90. The method provided by the invention is adopted to further convert the liquefied gas product after the propylene removal into propylene, and the propylene yield is obviously improved on the premise of not increasing the yield of the liquefied gas. The technology is characterized in that olefin polymerization, polymerization product cracking and alkane dehydrogenation reaction are simultaneously completed in the auxiliary riser. The auxiliary riser is also respectively provided with a first reaction zone (150-450 ℃) for promoting the light olefin to carry out a superposition reaction and a second reaction zone (450-650 ℃) for promoting the superposition product to further crack and dehydrogenate propane to generate propylene.

These documents mainly focus on how to carry out the polymerization reaction of liquefied gas or butene containing olefin to produce gasoline fraction or diesel oil fraction, and report less on how to further catalytically crack the liquefied gas or polymerization product of C4 olefin to produce ethylene and propylene. Moreover, the polymerization product of C4 olefins is mainly C8 olefins, and the propylene yield and propylene selectivity of direct catalytic cracking are not high, because C8 olefins synthesized by C4 are easier to be further catalytically cracked into 2C 4 olefins, and the synthesized C8 olefins are easier to be subjected to aromatization reaction in the catalytic cracking process to form coke, and the coke yield is higher.

Therefore, it is necessary to develop a combined process of low-carbon olefin polymerization and selective catalytic cracking multi-production chemical materials.

Disclosure of Invention

The method for recycling and cracking the hydrocarbons overcomes the key problems of high dry gas yield and poor selectivity of product propylene by recycling cracking of C4 fraction and gasoline fraction in the existing heavy oil catalytic cracking process route, and can obtain the product gasoline with ultralow olefins.

The present disclosure provides a process for hydrocarbon recycle cracking, the process comprising:

contacting a heavy hydrocarbon oil raw material with a first catalytic cracking catalyst in a first catalytic reactor to perform a first catalytic reaction to obtain a first reaction mixture, and performing gas-agent separation on the first reaction mixture in a settler to obtain a first carbon deposition catalyst and a first product;

introducing the first product and the second product into a first product separation system for separation to obtain a dry gas fraction, a C3 liquefied gas fraction, a liquefied gas fraction after C3 removal, a first light gasoline fraction, a second light gasoline fraction, a third light gasoline fraction, a heavy gasoline fraction, a diesel oil fraction and a heavy oil fraction;

subjecting the liquefied gas fraction after C3 removal and/or the first light gasoline fraction to a first polymerization reaction in the presence of a first polymerization reaction catalyst to obtain a first polymerization stream; subjecting said first superimposed stream and said third light gasoline fraction to a second superimposed reaction in the presence of a second superimposed reaction catalyst to obtain a second superimposed stream;

introducing the second superposed material flow into a second product separation system for separation to obtain a superposed product;

contacting the superimposed product and a second light gasoline fraction with a second catalytic cracking catalyst in a second catalytic reactor to perform a second catalytic reaction to obtain a second reaction mixture containing the second catalytic cracking catalyst; subjecting the second reaction mixture to a third catalytic reaction in a third catalytic reactor to obtain a third reaction mixture; and carrying out gas-agent separation on the third reaction mixture in the settler to obtain a second carbon-deposited catalyst and the second product.

In one embodiment, the distillation range of the product of the polymerization is 155 to 330 ℃, preferably 160 to 253 ℃.

In one embodiment, the reaction conditions of the first polymerization reaction are 300-450 ℃, 0.5-2.0 MPa, and the mass space velocity is 1-5 h^-1(ii) a The reaction conditions of the second polymerization reaction are 180-280 ℃, the pressure of 1.0-2.0 MPa and the mass space velocity of 0.7-2.0 h^-1。

In one embodiment, the first polymerization catalyst comprises 1 to 20 mass% of NiO, 40 to 80 mass% of HZSM-5 zeolite, and 10 to 50 mass% of alumina; the second polymerization catalyst comprises 1-12 mass% of NiO, 45-82 mass% of amorphous aluminum silicate and 10-50 mass% of alumina.

In one embodiment, the first and second folding reactions are carried out in a folding reactor comprising at least two folding reaction zones, the first folding reaction being carried out in a first folding reaction zone of the folding reactor and the second folding reaction being carried out in a second folding reaction zone of the folding reactor.

In one embodiment, the weight ratio of the polymerization product to the heavy hydrocarbon oil feedstock is 0.01 to 0.6: 1, preferably 0.05 to 0.3: 1.

in one embodiment, the first light gasoline fraction has a boiling range of 9 to 70 ℃, more preferably 9 to 60 ℃; the distillation range of the second light gasoline fraction is 50-110 ℃, and further preferably 60-100 ℃; the distillation range of the third light gasoline fraction is 90-150 ℃, and the preferred distillation range is 100-130 ℃. In one embodiment, the light hydrocarbon has an olefin content of from 30 to 90 wt.%, preferably from 45 to 90 wt.%, based on the total weight of the first light gasoline fraction. In the present disclosure, the light olefins of the first light gasoline fraction are primarily olefins containing 5 carbon atoms. The light olefins of the second light gasoline fraction are predominantly olefins containing 6 or 7 carbon atoms. The light olefins of the third light gasoline fraction are predominantly olefins containing 8 carbon atoms.

In one embodiment, the operating conditions of the first catalytic reaction include: the reaction temperature is 480-600 ℃; the reaction time is 0.5-10 seconds; the weight ratio of the agent to the oil is 5-15: 1; the weight ratio of water to oil is 0.05-1: 1. optionally, the operating conditions of the first catalytic reaction include: the reaction temperature is 480-600 ℃, for example 500-560 ℃, or 510-550 ℃, or 510-530 ℃; the reaction time is 0.5-10 seconds, such as 1-5 seconds, 2-3 seconds, or 1.5-4 seconds; the weight ratio of the agent oil is (5-15): 1 or (6-12): 1 or (8-10): 1; the weight ratio of water to oil is (0.05-1): 1 is, for example, (0.08-0.5): 1 or (0.1-0.3): 1.

in one embodiment, the operating conditions of the second catalytic reaction include: the reaction temperature is 520-750 ℃; the reaction time is 0.1-3 seconds; the weight ratio of the agent to the oil is 6-40: 1; the weight ratio of water to oil is 0.1-1: 1. Optionally, the operating conditions of the second catalytic reaction include: the reaction temperature is 520-750 ℃, for example, 520-600 ℃ or 520-560 ℃; the reaction time is 0.1-3 seconds, such as 0.5-3 seconds, 1-3 seconds, or 1.3-3 seconds; the weight ratio of the agent to the oil is (6-40): 1, for example, (7 to 30): 1 or (8-25): 1 or (10-20): 1; the weight ratio of water to oil is (0.1-1): 1 is, for example, (0.08-0.5): 1 or (0.1-0.3): 1.

in one embodiment, the third catalytic reactor is a fluidized bed reactor, and the operating conditions of the third catalytic reactor are: the reaction temperature is 450-750 ℃, preferably 510-560 ℃; the weight hourly space velocity is 1-30 h^-1。

In one embodiment, the first and second catalytic cracking catalysts each contain a shape selective zeolite having an average pore diameter of less than 0.7nm, the shape selective zeolite being at least one selected from the group consisting of zeolites having an MFI structure, ferrierites, chabazites, dachiardite, erionites, a-type zeolites, epistillomites, and laumontites.

In one embodiment, the first catalytic reactor, the second catalytic reactor, and the third catalytic reactor are each one selected from the group consisting of a riser reactor, a downer reactor, a fluidized bed reactor, a riser and downer combined reactor, a riser and fluidized bed combined reactor, and a downer and fluidized bed combined reactor.

In one embodiment, the method further comprises:

stripping the first and second carbon-deposited catalysts in a stripping zone of the settler,

regenerating the stripped first and second carbon-deposited catalysts in a regenerator to obtain first and second regenerated catalysts, respectively;

feeding said first regenerated catalyst to said first catalytic reactor as said first catalytic cracking catalyst and feeding said second regenerated catalyst to said second catalytic reactor as said second catalytic cracking catalyst.

In one embodiment, the regeneration is carried out at a temperature of 600 to 800 ℃.

In one embodiment, the first regenerated catalyst has a temperature of 560-.

In one embodiment, the temperature of the second regenerated catalyst is 560-.

In one embodiment, the first carbon-deposited catalyst is stripped in a first stripping zone of the settler and the second carbon-deposited catalyst is stripped in a second stripping zone of the settler.

In one embodiment, the heavy hydrocarbon oil feedstock is at least one selected from the group consisting of petroleum hydrocarbon oils, synthetic oils, coal liquefaction oils, oil sand oils, and shale oils, preferably petroleum hydrocarbon oils, and the petroleum hydrocarbon oils are at least one selected from the group consisting of atmospheric gas oils, vacuum gas oils, coker gas oils, deasphalted oils, hydrogenated tail oils, atmospheric residues, vacuum residues, and crude oils.

In one embodiment, introducing the second superimposed stream into a second product separation system for separation also produces an unreacted C4 fraction, an unreacted first light gasoline fraction and an unreacted third light gasoline fraction, and combining the unreacted first light gasoline fraction, the unreacted third light gasoline fraction, a portion of the second light gasoline fraction, and the heavy gasoline fraction to produce a gasoline product having a low olefin content. The olefin content in the gasoline product with low olefin content can reach 5-10%.

The method for recycling and cracking the hydrocarbons overcomes the problems of high yield of dry gas, poor selectivity of propylene products and high olefin content in gasoline products in the catalytic cracking reaction of the heavy oil in the prior process route. Specifically, compared with the traditional method for producing low-carbon olefins by catalytic conversion of heavy hydrocarbon oil, the method disclosed by the invention has the following beneficial effects of any one or more, preferably all the following effects:

in the prior art, the C4 fraction and the light gasoline fraction are generally directly recycled to a catalytic cracking or catalytic cracking device for reaction, and the cracking of the C4 fraction and the light gasoline fraction requires high reaction severity, namely high reaction temperature, high catalyst-to-oil ratio, high residence time and the like, so that the yield of dry gas and coke is high, and the selectivity of propylene is poor. Meanwhile, part of components of the light gasoline, such as fractions at 100-120 ℃, are rich in C8 olefin and alkane, and are more prone to middle position homolytic cracking during cracking reaction to generate C4 hydrocarbon, and the proportion of generated C3 olefin is low.

The research of the applicant of the present disclosure finds that the performance of the catalytic cracking of the superimposed product after the superimposition of different C4 olefins (n-butene, isobutene, trans-2-butene, cis-2-butene) has large difference in the yield of ethylene and propylene; the products obtained by different-depth superposition reaction of the same C4 olefin have larger difference in the performance of producing more ethylene and propylene through catalytic cracking; different olefins in the liquefied gas are superposed to obtain products, and the products are subjected to catalytic cracking to produce ethylene and propylene in different properties. Meanwhile, the different reaction processes of catalytic cracking also have decisive influence on the performance of the low-carbon olefin polymerization product for cracking to produce more ethylene and propylene.

The present disclosure provides for the precise separation of gasoline fractions according to their structural characteristics: the first light gasoline fraction is rich in C5 hydrocarbon, and the molecular chain is short and difficult to crack; the second light gasoline fraction is rich in C6 hydrocarbons and a small amount of C7 hydrocarbons, and is easy to generate propylene by directly breaking from the middle of molecules; the third light gasoline fraction is rich in C8 hydrocarbons and is susceptible to cracking directly from the middle of the molecule to produce butenes instead of propylene; meanwhile, the separated products are accurately superposed into hydrocarbons such as C9, C12 and the like which are easy to crack into propylene, so that the reaction severity can be greatly reduced, and the selectivity of propylene cracking of the circulating material flow is greatly improved. The reaction severity required by the cracking reaction of the newly generated C9 and C12+ olefins is greatly reduced, and the propylene selectivity is greatly improved by the method that the first light gasoline fraction and the C4 fraction are overlapped to generate C9 olefins and C8 olefins, and the C8 olefins, the third light gasoline fraction and the C4 fraction are overlapped to generate C12 and C12 olefins.

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 a cyclic catalytic conversion process for hydrocarbons provided by the present disclosure.

Description of the reference numerals

1 first catalytic reactor

2-1 second catalytic reactor

2-2 third catalytic reactor

7 stripping zone

6 settling vessel

9 regenerator

61 first product separation System

70 superimposed reactor

70-1 polymerization reactor first polymerization reaction zone

70-2 superimposed reactor second superimposed reaction zone

62 second product separation System

11 first catalytic cracking catalyst transfer line

12 second catalytic cracking catalyst transfer line

17 spent catalyst transfer line

21 line (injecting heavy hydrocarbon oil feedstock)

20 pipeline (transporting catalytic cracking reaction oil gas)

24 pipeline (conveying dry gas)

25 line (transport C3 fraction)

26 pipeline (liquefied gas after C3 is removed in conveying)

27 line (carrying the first light gasoline fraction)

28 line (for the second light gasoline fraction)

29 line (conveying the third light gasoline fraction)

30 pipeline (conveying heavy petrol)

31 pipeline (transportation diesel)

32 pipeline (heavy oil conveying)

41 pipeline (injecting atomized steam)

43 pipeline (injecting atomized steam)

44 pipeline (injecting atomized steam)

47 line (steam stripping injection)

51 line (injection pre-lifting medium)

52 line (injection pre-lifting medium)

90 pipeline (air injection)

33 pipe (for second superimposed material flow)

34 line (for unreacted C4 fraction)

35 line (for the first unreacted light gasoline fraction)

36 line (for unreacted second light gasoline fraction)

37 pipeline (transport superimposed product)

90 pipeline (air conveying)

91 pipeline (flue gas)

100 external heat exchanger ((removing excess heat from regenerator, lowering regeneration temperature)

Detailed Description

The technical solution of the present invention is further explained below according to specific embodiments. The scope of protection of the invention is not limited to the following examples, which are set forth for illustrative purposes only and are not intended to limit the invention in any way.

contacting the superimposed product and the second light gasoline fraction with a second catalytic cracking catalyst in a second catalytic reactor to perform a second catalytic reaction to obtain a second reaction mixture containing the second catalytic cracking catalyst; subjecting the second reaction mixture to a third catalytic reaction in a third catalytic reactor to obtain a third reaction mixture; and (3) carrying out gas-agent separation on the third reaction mixture in the settler to obtain a second carbon-deposited catalyst and the second product, wherein the distillation range of the superimposed product is 155-330 ℃, and preferably 160-253 ℃.

In the existing catalytic cracking device, C4 fraction and light gasoline fraction in or outside a catalytic cracking product are directly recycled to the catalytic cracking or catalytic cracking device for reaction, and the C4 fraction and the light gasoline fraction need high reaction severity, namely high reaction temperature, high catalyst-oil ratio, high retention time and the like, due to the fact that molecular chains are short, cracking occurs, and therefore high dry gas and coke yield is caused, and propylene selectivity is poor. Meanwhile, part of components of the light gasoline, such as fractions at 100-120 ℃, are rich in C8 olefin and alkane, and are more prone to middle position homolytic cracking during cracking reaction to generate C4 hydrocarbon, and the proportion of generated C3 olefin is low. Therefore, in the prior art, although the propylene can be increased by circularly cracking the C4 fraction and the light gasoline fraction through the catalytic cracking unit, the conversion per pass of the reaction is low, a large recycle ratio is required, the selectivity of the propylene is low, the selectivity of the dry gas is high, the olefin in the product gasoline cannot be completely converted, and the olefin content of the product gasoline is high.

The research of the inventor of the present disclosure finds that the difficulty degree of gasoline cracking reaction of different fraction sections in the light gasoline fraction has large difference, and the selectivity of generating propylene and butylene has large difference. If the cracking reaction of the light gasoline fraction at 9-60 ℃ is difficult, the reaction temperature needs to be higher than 600-640 ℃, so that the dry gas yield is higher, the selectivity of propylene is poorer, and the propylene yield is lower; the cracking reaction of the light gasoline fraction at the temperature of 60-100 ℃ is easy, the yield of propylene is high, the selectivity of propylene is good, and the method is suitable for direct cracking; the cracking reaction of the light gasoline fraction at 100-130 ℃ is easy, but the yield of propylene is low, the yield of butylene is high, the selectivity of propylene is not good, and the cracking is not suitable for direct cracking. Therefore, different light gasoline fractions need to be accurately separated and then classified.

The inventor of the present disclosure finds, through laboratory research, that a catalytic cracked C4 fraction and a light gasoline fraction with a distillation range of 9-60 ℃ are subjected to a superposition reaction to generate a superposition product, and the propylene selectivity of the superposition product after the catalytic cracking reaction is greatly improved, the propylene yield is higher, and the required reaction temperature is greatly reduced; however, if the catalytic cracked C4 fraction and the light gasoline fraction with the distillation range of 60-100 ℃ are subjected to a superposition reaction to generate a superposition product, the propylene selectivity of the superposition product after the catalytic cracking reaction is greatly reduced, and the propylene yield is reduced on the contrary; if the C4 fraction subjected to catalytic cracking and the light gasoline fraction with the distillation range of 100-130 ℃ are subjected to a superposition reaction to generate a superposition product with the distillation range of 155-330 ℃, the propylene selectivity of the superposition product subjected to catalytic cracking is greatly improved, the propylene yield is higher, and the required reaction temperature is lower.

The inventor of the present disclosure finds that different reaction severity levels are required for the polymerization reaction between different hydrocarbons, different reaction conditions are required to obtain the target polymerization product, and the propylene selectivity and the propylene yield of the fraction with the distillation range of 155-330 ℃ in the polymerization product are highest through catalytic cracking.

The method provided by the present disclosure is further described below with reference to fig. 1, but the present disclosure is not limited thereto.

In fig. 1, a first catalytic reactor 1 is a riser reactor, a second catalytic reactor 2-1 is a riser reactor, and a third catalytic reactor 2-2 is a fluidized bed reactor. The fluidized bed 2-2 in the settler 6 is located above the stripping zone 7. The first catalytic cracking catalyst (hot regenerated catalyst) enters the bottom of the riser of the first catalytic reactor 1 from the regenerator 9 through the first catalytic cracking catalyst transfer line 11 and is accelerated to flow upward by the action of a pre-lift medium (which may be selected from, for example, steam, catalytic dry gas, post-de-ethylene dry gas, etc.) injected through line 51. The preheated heavy hydrocarbon oil raw material is mixed with atomized steam from a pipeline 41 through a pipeline 21 and then injected into a lifting pipe, and the weight ratio of the water steam to the hydrocarbon oil raw material is (0.05-1): 1, the outlet temperature of the riser reactor 1 is 480-600 ℃, the reaction time in the riser reactor 1 is 0.5-10 seconds, the weight ratio of the catalyst to the hydrocarbon oil raw material is 5-15, and the absolute pressure in the settler 6 is 0.1-0.40 MPa. The mixture of reaction oil gas and catalyst in the riser reactor 1 is separated by a fast-separating device at the outlet, the first carbon-deposited catalyst is introduced into a stripping zone 7, the separated reaction oil gas (first product) is sent to a subsequent first product separation system 61 for product separation through a settler 6 and a pipeline 20 at the top of the settler 6, and products such as a dry gas fraction, a C3 fraction (propylene and propane), a liquefied gas fraction (C4 fraction) after C3 is removed, a first light gasoline fraction, a second light gasoline fraction, a third light gasoline fraction, a heavy gasoline fraction, a diesel oil fraction, and a heavy oil fraction are obtained after separation (respectively led out through

pipelines

24, 25, 26, 27, 28, 29, 30, 31 and 32).

The second catalytic cracking catalyst (hot regenerated catalyst) enters the bottom of the riser of the second catalytic reactor 2-1 from the regenerator 9 through a second catalytic cracking catalyst delivery line 12 and is accelerated to flow upward by the action of a pre-lift medium (which may be selected from, for example, steam, catalytic dry gas, post-de-ethylene dry gas, etc.) injected through line 52. The second light gasoline fraction 28 from the first product separation system 61 is injected into the second catalytic reactor 2-1 mixed with the atomized steam from line 43, while the superimposed product stream 37 from the second product separation system 62 is also injected into the second catalytic reactor 2-1 mixed with the atomized steam from line 44. The weight ratio of the water vapor to the hydrocarbon oil raw material (including the second light gasoline fraction and the superimposed product) is (0.1-1): 1, the outlet temperature of a riser reactor 2-1 is 520-750 ℃, the reaction time in the riser reactor 2-1 is 0.1-3 seconds, and the weight ratio of a catalyst to a hydrocarbon oil raw material (comprising a second light gasoline fraction and a superimposed product) is 6-40: 1. here, the reaction time refers to the residence time of the oil gas in the reactor.

Reaction oil of riser reactor 2-1The mixture of gas and catalyst is further introduced into a third catalytic reactor 2-2 (fluidized bed reactor) through a riser outlet to continue reacting, the reaction temperature of the fluidized bed 2-2 is 450-750 ℃, and the weight hourly space velocity is 1-30 h^-1The reacted oil gas and a part of the carbon-deposited spent catalyst enter a settler 6 for separation through a fluidized bed reactor 2-2, a second carbon-deposited catalyst is separated and enters a stripping zone 7, and the separated reaction oil gas (a second product) is also sent to a subsequent first product separation system 61 for product separation through the settler 6 and a pipeline 20 at the top of the settler.

The liquefied gas fraction 26(C4 fraction) after the C3 is removed and the first light gasoline fraction 27 enter a first polymerization reaction zone 70-1 of a polymerization reactor 70 to react to obtain a first polymerization stream. The first superimposed stream and the third light gasoline fraction 29 are passed to a second superimposed reaction zone 70-2 of superimposed reactor 70 for reaction to produce a second superimposed stream 33.

This second superimposed stream 33 is introduced into a second product separation system 33 for separation to yield an unreacted C4 fraction, an unreacted first light gasoline fraction, an unreacted third light gasoline fraction, and a superimposed product (withdrawn via lines 34, 35, 36, and 37, respectively). The superimposed product may be introduced into the second catalytic reactor 2-1 for circulating reaction.

The stripping steam is injected into the stripping zone 7 through a pipeline 47 and contacts with the coked spent catalyst in a countercurrent manner, and the reaction oil gas carried by the spent catalyst is stripped as completely as possible. Air is injected into the regenerator 9 through a pipeline 90, the stripped first carbon-deposited catalyst and the stripped second carbon-deposited catalyst are sent into the regenerator 9 through a spent agent conveying pipeline 17, and are contacted with the heated air and regenerated at the temperature of 600-800 ℃, so that a first regenerated catalyst and a second regenerated catalyst are obtained and are used as a first catalytic cracking catalyst and a second catalytic cracking catalyst for recycling. The regeneration flue gas is led out through a line 91. In fig. 1, 100 is an external heat remover, which is used to remove heat from the regenerator by heat exchange if necessary, and to lower the regeneration temperature.

As described above, the heavy hydrocarbon oil feedstock is contacted with the first catalytic cracking catalyst in the first catalytic reactor to perform the first catalytic reaction to obtain the first reaction mixture, and the first reaction mixture is subjected to gas-agent separation in the settler to obtain the first carbon deposition catalyst and the first product.

According to the present disclosure, a heavy hydrocarbon oil feedstock is contacted with a first catalytic cracking catalyst in a fluidized state in a first catalytic reactor to perform a first catalytic reaction. The operating conditions of the first catalytic reaction may include: the reaction temperature is 480-600 ℃, for example 500-560 ℃, or 510-550 ℃, or 510-530 ℃; the reaction time is 0.5-10 seconds; the weight ratio of the catalyst to the oil (namely the weight ratio of the first catalytic cracking catalyst to the heavy hydrocarbon oil raw material) is (5-15): 1 or (6-12): 1 or (8-10): 1; the weight ratio of water to oil (namely the weight ratio of water vapor to heavy hydrocarbon oil raw material) is (0.05-1): 1 is, for example, (0.08-0.5): 1 or (0.1-0.3): 1. here, the reaction time refers to the residence time of the reaction oil gas in the first catalytic reactor.

Introducing the first product into a first product separation system for separation, wherein separated products comprising a dry gas fraction, a liquefied gas product, a gasoline product, a diesel oil fraction and a heavy oil fraction can be obtained according to different distillation ranges (boiling point ranges); and the liquefied gas product can be further separated into a C3 liquefied gas fraction (propylene and propane) and a liquefied gas fraction (C4 fraction) after C3 is removed, and the gasoline product can be further separated into a first light gasoline fraction, a second light gasoline fraction, a third light gasoline fraction and a heavy gasoline fraction.

Methods for separating the first product and the second product in the first product separation system are known, and for example, various fractions can be obtained by separation according to a set distillation range in the form of a fractionating column, a rectifying column, or the like: separation products including a dry gas fraction, a liquefied gas product, a gasoline product, a diesel fraction, and a heavy oil fraction; and the liquefied gas product can be further separated into a C3 liquefied gas fraction (propylene and propane) and a liquefied gas fraction (C4 fraction) after C3 is removed, and the gasoline product can be further separated into a first light gasoline fraction, a second light gasoline fraction, a third light gasoline fraction and a heavy gasoline fraction. The first product separation system may include one or more fractionation or rectification columns.

Optionally, the distillation range of the first light gasoline fraction is 9-70 ℃, and more preferably 9-60 ℃; the distillation range of the second light gasoline fraction is 50-110 ℃, and further preferably 60-100 ℃; the distillation range of the third light gasoline fraction is 90-150 ℃, and the preferred distillation range is 100-130 ℃. The light hydrocarbon has an olefin content of 30 to 90 wt%, preferably 45 to 90 wt%, based on the total weight of the first light gasoline fraction. The light olefins of the first light gasoline fraction are predominantly olefins containing 5 carbon atoms. The light olefins of the second light gasoline fraction are predominantly olefins containing 6 or 7 carbon atoms. The light olefins of the third light gasoline fraction are predominantly olefins containing 8 carbon atoms.

In one embodiment, the dry gas fraction is primarily hydrogen, methane, ethylene and ethane, the C3 liquefied gas fraction is propylene and propane, the liquefied gas fraction after C3 removal is C4 fraction, the distillation range of the heavy gasoline fraction is 150-.

Then, the second light gasoline fraction and the superimposed product are contacted with a second catalytic cracking catalyst in a second catalytic reactor to carry out a second catalytic reaction to obtain a second reaction mixture containing the second catalytic cracking catalyst; subjecting the second reaction mixture (together with a second catalytic cracking catalyst) to a third catalytic reaction in a third catalytic reactor to obtain a third reaction mixture; and carrying out gas-agent separation on the third reaction mixture in the settler to obtain a second carbon-deposited catalyst and a second product. The second product may be introduced into the first product separation system with the first product for separation.

In one embodiment, the weight ratio of the superimposed product to the heavy hydrocarbon oil feedstock is (0.01-0.6): 1, preferably (0.05-0.3): 1. the weight ratio of the second light gasoline fraction to the heavy hydrocarbon oil raw material is 0.01-0.6: 1, preferably 0.05 to 0.03: 1.

according to the present disclosure, the superimposed product and the second light gasoline fraction are contacted with a second catalytic cracking catalyst in a fluidized state in a second catalytic reactor for a second catalytic reaction. The operating conditions of the second catalytic reaction may include: the reaction temperature is 520-750 ℃, for example, 520-600 ℃ or 520-560 ℃; the reaction time is 0.1-3 seconds, such as 0.5-3 seconds, or 1-3 seconds, or 1.3-3 seconds; the catalyst-oil weight ratio (namely the weight ratio of the second catalytic cracking catalyst to the total amount of the superimposed product and the second light gasoline fraction) is (6-40): 1, for example, (7 to 30): 1, or (8-25): 1 or (10-20); the weight ratio of water to oil (namely the weight ratio of the water vapor to the total amount of the superimposed product and the second light gasoline fraction) is (0.1-1): 1 is, for example, (0.08-0.5): 1 or (0.1-0.3): 1. here, the reaction time refers to the residence time of the oil gas in the second catalytic reactor.

The second light gasoline fraction is rich in C6 olefins and a small amount of C7 olefins, wherein each C6 olefin molecule is susceptible to breaking directly from the middle of the molecule to produce 2 propylene and each C7 olefin molecule is susceptible to breaking directly from the middle of the molecule to produce 1 propylene and 1 butene molecule in the second catalytic reactor. The superposed product is rich in C9-C13 olefin, the reaction severity required by hydrocarbon cracking reaction is greatly reduced as the molecular chain is longer, for example, C12 olefin molecules are easy to directly generate 2C 6 olefin molecules by intermediate chain scission, each C6 olefin molecule is further subjected to intermediate scission to generate 2 propylene, and the propylene selectivity is greatly improved.

According to the present disclosure, the heavy hydrocarbon oil feedstock may be at least one selected from the group consisting of petroleum hydrocarbon oils, synthetic oils, coal liquefaction oils, oil sand oils, and shale oils. The synthetic oil can be distillate oil obtained by Fischer-Tropsch (F-T) synthesis of coal and natural gas. Preferably, the heavy hydrocarbon oil feedstock is a petroleum hydrocarbon oil, such as at least one selected from the group consisting of atmospheric gas oil, vacuum gas oil, coker gas oil, deasphalted oil, hydrogenated tail oil, atmospheric residue, vacuum residue, and crude oil.

According to the present disclosure, the first catalytic cracking catalyst and the second catalytic cracking catalyst may be the same catalytic cracking catalyst, or may be different catalytic cracking catalysts, and are preferably the same catalytic cracking catalyst, which are conventionally used in the field of catalytic cracking reactions.

The present disclosure is not particularly limited with respect to the specific kinds of the first catalytic cracking catalyst and the second catalytic cracking catalyst. Preferably, the first catalytic cracking catalyst and the second catalytic cracking catalyst each contain a shape-selective zeolite having an average pore diameter of less than 0.7nm, and the shape-selective zeolite may be at least one selected from the group consisting of zeolite having an MFI structure, ferrierite, chabazite, dachiardite, erionite, a-type zeolite, epistilbite, and turbid zeolite. Wherein the MFI structure zeolite may be one or more of ZSM-5 and ZRP series zeolites, and may be one or more of ZSM-5 and ZRP series zeolites modified with at least one element of RE, P, Fe, Co, Ni, Cu, Zn, Mo, Mn, Ga and Sn. In an alternative embodiment of the present disclosure, the catalytic cracking catalyst comprises, based on the dry weight (weight calcined at 800 ℃ for 1 hour) of the catalytic cracking catalyst, 15 to 50 wt% of clay on a dry basis, 15 to 50 wt% of molecular sieve on a dry basis, and 10 to 35 wt% of binder on a dry basis, wherein the molecular sieve is a zeolite of MFI structure or consists of 25 to 100 wt% of zeolite of MFI structure and 0 to 75 wt% of other zeolites except for zeolite of MFI structure; the MFI structure zeolite is preferably a ZSM-5 molecular sieve and/or an HZSM-5 molecular sieve modified with phosphorus and at least one element selected from RE, P, Fe, Co, Ni, Cu, Zn, Mo, Mn, Ga and Sn. The clay is preferably, for example, one or more selected from kaolin, halloysite, montmorillonite, diatomaceous earth, halloysite, pseudohalloysite, saponite, rectorite, sepiolite, attapulgite, hydrotalcite and bentonite. The binder is one or more of acidified pseudo-boehmite, aluminum sol, silica sol, magnesium aluminum sol, zirconium sol and titanium sol, preferably acidified pseudo-boehmite, aluminum sol and the like.

In accordance with the present disclosure, the first catalytic reactor and the second catalytic reactor may be catalytic conversion reactors well known to those skilled in the art, for example, the first catalytic reactor and the second catalytic reactor are each one selected from a riser reactor, a downer reactor, a fluidized bed reactor, a riser and downer combined reactor, a riser and fluidized bed combined reactor, and a downer and fluidized bed combined reactor. The fluidized bed reactor may be one selected from the group consisting of a fixed fluidized bed reactor, a bulk fluidized bed reactor, a bubbling bed reactor, a turbulent bed reactor, a fast bed reactor, a transport bed reactor, and a dense phase fluidized bed. The riser reactor, the downer reactor and the fluidized bed reactor can be equal-diameter riser reactors, downer reactors and fluidized bed reactors, and can also be variable-diameter riser reactors, downer reactors and fluidized bed reactors. In one embodiment, the first catalytic reactor and the second catalytic reactor are both riser reactors.

In accordance with the present disclosure, the third catalytic reactor is preferably a fluidized bed reactor, which may be one selected from the group consisting of a fixed fluidized bed reactor, a bulk fluidized bed reactor, a bubbling bed reactor, a turbulent bed reactor, a fast bed reactor, a transport bed reactor, and a dense phase fluidized bed. The fluidized bed reactor can be in a constant-diameter fluidized bed structure or a variable-diameter fluidized bed structure. The operating conditions of the third catalytic reactor may be: the reaction temperature is 450-750 ℃, for example 480-600 ℃, or 500-580 ℃, or 510-560 ℃, or 520-550 ℃, preferably 510-560 ℃; the weight hourly space velocity is 1-30 h^-1For example, 3 to 28h^-1Or 5 to 25 hours^-1Or 6 to 20 hours^-1. According to the present disclosure, the third catalytic reactor is located inside a settler, which has an absolute pressure of 0.15-0.40 MPa. In one embodiment, the second catalytic reactor is a riser reactor, the third catalytic reactor is a fluidized bed reactor, and the second catalytic reactor and the third catalytic reactor are connected in series.

The polymerization product may be produced by the process of the present disclosure, as well as from other devices and processes, in accordance with the present disclosure. In order to further improve the yield of high-quality propylene, the distillation range of the superimposed product is 155-330 ℃, and the distillation range is more preferably 160-253 ℃. The olefin content of the polymerization product may be 60 to 100 wt%, preferably 80 to 100 wt%, for example 80 to 90 wt%, or 85 to 95 wt%, or 90 to 100 wt%, based on the total weight of the polymerization product.

According to the present disclosure, the superimposed product may be obtained by subjecting the liquefied gas fraction after the removal of C3 (C4 fraction), the first light gasoline fraction, and the third light gasoline fraction to a subsequent superimposing reaction, thereby providing the superimposed product by the method of the present disclosure. Subjecting the liquefied gas fraction after C3 removal and/or the first light gasoline fraction to a first polymerization reaction in the presence of a first polymerization reaction catalyst to obtain a first polymerization stream; and carrying out a second superposition reaction on the first superposition material flow and the third light gasoline fraction in the presence of a second superposition reaction catalyst to obtain a second superposition material flow.

The liquefied gas fraction after the C3 removal is rich in C4 olefins, 2C 4 olefins can carry out a polymerization reaction to generate C8 olefins under the action of an acid catalyst at a lower reaction temperature and a higher reaction pressure of the first polymerization reactor, and simultaneously the first light gasoline is rich in C5 olefins, and C5 olefins can also carry out a polymerization reaction with C4 olefins to generate C9 olefins.

The C8 olefin, C9 olefin and C8 olefin from the third polymerization reactor can be further reacted with unreacted C4 olefin and C5 olefin under the action of an acid catalyst at a lower reaction temperature and a higher reaction pressure in the second polymerization reactor, the C8 olefin is further subjected to polymerization reaction with C4 olefin to form C12 olefin, and the C9 olefin is further subjected to polymerization reaction with C4 olefin to form C13 olefin to form a polymerization product.

According to the disclosure, the reaction conditions of the first polymerization reaction are 300-450 ℃, the pressure of 0.5-2.0 MPa and the mass space velocity of 1-5 h^-1. According to the disclosure, the reaction conditions of the second polymerization reaction are 180-280 ℃ of temperature, 1.0-2.0 MPa of pressure and 0.7-2.0 h of mass space velocity^-1。

According to the present disclosure, the first polymerization catalyst comprises 1-20 mass% of NiO, 40-80 mass% of HZSM-5 zeolite, and 10-50 mass% of alumina. According to the present disclosure, the second polymerization catalyst includes 1 to 12 mass% of NiO, 45 to 82 mass% of amorphous aluminum silicate, and 10 to 50 mass% of alumina.

According to the present disclosure, the first and second stacking reactions may be carried out in a stacking reactor comprising at least two stacking reaction zones, the first stacking reaction being carried out in a first stacking reaction zone of the stacking reactor and the second stacking reaction being carried out in a second stacking reaction zone of the stacking reactor. The conditions and catalyst in the first polymerization zone may be the reaction conditions and first polymerization catalyst of the first polymerization reaction as described above, and the conditions and catalyst in the second polymerization zone may be the reaction conditions and second polymerization catalyst of the second polymerization reaction as described above.

As mentioned above, the method of the present disclosure further comprises introducing the second superimposed stream into a second product separation system for separation to obtain the superimposed product, wherein the distillation range of the superimposed product is 155 to 330 ℃, and more preferably 160 to 253 ℃. The product of the superposition may be recycled to the second catalytic reactor for reaction. In addition to the superimposed product, an unreacted C4 fraction, an unreacted first light gasoline fraction, and an unreacted third light gasoline fraction may be obtained by the second product separation system. These unreacted C4 fraction, the unreacted first light gasoline fraction, and the unreacted third light gasoline fraction may be recycled to the polymerization reaction to increase the amount of polymerization product until the olefin content in the above fractions is less than 5% by weight, which may not be recycled based on energy consumption considerations. These unreacted first light gasoline fraction, as well as the unreacted third light gasoline fraction, may also be combined with the heavy gasoline fraction from the first product separation system and a portion of the second light gasoline fraction to produce a low olefin gasoline product. The olefin content in the low olefin gasoline product may be 5-10%.

The separation of the second superimposed stream in the second product separation system is primarily by carbon number or distillation range. The unreacted C4 fraction separated was predominantly C4 alkanes, the unreacted first light gasoline fraction was predominantly C5 alkanes, and the unreacted third light gasoline fraction was predominantly C8 alkanes and aromatics.

According to the present disclosure, the method further comprises:

feeding said first regenerated catalyst as said first catalytic cracking catalyst into said first catalytic reactor and feeding said second regenerated catalyst as said second catalytic cracking catalyst into said second catalytic reactor.

According to the present disclosure, the stripping zone in the settler may be divided into a first stripping zone that strips the first carbon-deposited catalyst and a second stripping zone that strips the second carbon-deposited catalyst, which can prevent the two carbon-deposited catalysts from mixing in the stripping zone. The manner of dividing the first stripping zone and the second stripping zone is not particularly limited as long as the above-mentioned object can be satisfied, and may be divided by a baffle, for example. The first stripping zone and the second stripping zone may also be the same stripping zone. By stripping, the reaction oil gas carried by the spent catalyst can be stripped as clean as possible. The absolute pressure in the settler may be 0.1-0.40 MPa.

According to the present disclosure, the stripped first and second carbon-deposited catalysts are introduced into a regenerator for regeneration to obtain first and second regenerated catalysts, the first regenerated catalyst is sent into a first catalytic reactor as the first catalytic cracking catalyst, and the second regenerated catalyst is sent into a second catalytic reactor as the second catalytic cracking catalyst. Thereby, the first catalytic cracking catalyst and the second catalytic cracking catalyst can be recycled and reused. In the regenerator, the stripped first and second carbon-deposited catalysts are contacted with heated air and regenerated at 600-800 ℃. The regenerator can also be divided into two zones, a first regeneration zone and a second regeneration zone, wherein the first regeneration zone is used for regenerating the first carbon-deposited catalyst, and the second regeneration zone is used for regenerating the second carbon-deposited catalyst. The manner of dividing the first regeneration zone and the second regeneration zone is not particularly limited as long as the above-described object can be satisfied, and may be divided by a baffle, for example.

According to the present disclosure, the temperature of the first regenerated catalyst is 560-. In another embodiment, the temperature of the second regenerated catalyst is 560-.

The methods provided by the present disclosure are further illustrated below by examples, but the present disclosure is not limited thereto.

The first and second catalytic cracking catalysts used in the following examples and comparative examples are cracking catalysts manufactured by the Chinese petrochemical catalyst, Qilu division, having a trade mark of OMT, and having specific properties as shown in Table 1-1, and comprise shape selective zeolite having an average pore diameter of less than 0.7 nm.

Example 1

Example i illustrates the effect of the process provided by the present disclosure on high selectivity propylene production during a hydrocarbon recycle catalytic conversion process.

The experiment was carried out using a medium-sized apparatus for continuous reaction-regeneration operation having three reactors, in which the first catalytic reactor was a riser, and the riser reactor had an inner diameter of 16 mm and a height of 3800 mm. The second catalytic reactor was a riser reactor with an internal diameter of 16 mm and a height of 3200 mm. The outlet of the second riser reactor was introduced into a fluidized bed reactor, which had an internal diameter of 64 mm and a height of 300 mm, as the third catalytic reactor.

The first catalytic cracking catalyst is a first regenerated catalyst with the temperature of 680 ℃, enters the bottom of a riser reactor of the first catalytic reactor through a first catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. The heavy hydrocarbon oil raw material is atmospheric residue (the main properties are shown in table 2) which is heated to 350 ℃ by a preheating furnace and then mixed with atomized water vapor, and the mixture is sprayed into a first catalytic reactor through a feeding nozzle and contacts with a hot first regenerated catalyst to carry out catalytic conversion reaction. The reaction oil gas (first product) and the first carbon-deposited catalyst enter a settler from the outlet of a riser of the first catalytic reactor for rapid separation, and the first carbon-deposited catalyst enters a stripping zone for stripping. The second catalytic cracking catalyst is a second regenerated catalyst with the temperature of 680 ℃, enters the bottom of a riser reactor of the second catalytic reactor through a second catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. And the second light gasoline fraction (9-60 ℃) and the superimposed product (155-330 ℃) sequentially enter a second catalytic reactor to contact with a hot second regenerated catalyst for catalytic reaction. The weight ratio of the second light gasoline fraction to the heavy hydrocarbon oil feedstock is 0.05: 1, the weight ratio of the superimposed product to the heavy hydrocarbon oil raw material is 0.35: 1. introducing the reaction oil mixture from the second catalytic reactor into a fluidized bed of a third catalytic reactor at the outlet of a riser for further reaction, separating oil gas (a second product) and a second carbon-deposited catalyst after the reaction in a settler through the fluidized bed reactor, and allowing the separated second carbon-deposited catalyst to enter a stripping zone for stripping under the action of gravity. And the first carbon-deposited catalyst and the second carbon-deposited catalyst after steam stripping enter a regenerator through a spent agent conveying pipe, are contacted with heated air and are regenerated at 700 ℃, and the obtained hot first regenerated catalyst and the second regenerated catalyst are respectively returned to the first catalytic reactor and the second catalytic reactor for recycling. Reaction oil gas from the first catalytic reactor and the third catalytic reactor is led out from the settler together, introduced into a product separation system for product separation to obtain a gas product and various liquid products, and simultaneously, part of the reaction oil gas is separated to obtain a C4 fraction, a first light gasoline fraction, a second light gasoline fraction, a third light gasoline fraction and a heavy gasoline fraction (the distillation ranges are shown in Table 3).

Introducing the C4 fraction and a first light gasoline fraction into a first polymerization reaction zone of a polymerization reactor, and carrying out a first polymerization reaction in the presence of a first polymerization reaction catalyst to obtain a first polymerization material flow; and further introducing the first superposed material flow and the third light gasoline fraction into a second superposed reaction zone of the superposed reactor, and carrying out a second superposed reaction in the presence of a second superposed reaction catalyst to obtain a second superposed material flow. Second oneThe superimposed stream is introduced into a second product separation system to be separated into an unreacted C4 fraction, an unreacted first light gasoline fraction, an unreacted third light gasoline fraction and a superimposed product. Wherein the first polymerization catalyst is an SXC-6 catalyst produced by the petrochemical engineering scientific research institute of China petrochemical engineering Ltd, and the second polymerization catalyst is an LXC-10 catalyst produced by the petrochemical engineering scientific research institute of China petrochemical engineering Ltd; the polymerization reactor is a fixed bed reactor; the first polymerization reaction conditions are as follows: the reaction temperature is 350 ℃, the reaction pressure is 1.0Mpa, and the weight hourly space velocity is 3.0h^-1(ii) a The second polymerization reaction conditions are as follows: the reaction temperature is 250 ℃, the reaction pressure is 2.0Mpa, and the weight hourly space velocity is 1.0h^-1. The SXC-6 catalyst and LXC-10 catalyst were prepared as follows.

The unreacted first light gasoline fraction, the unreacted third light gasoline fraction and the heavy gasoline fraction are blended with a portion of the second light gasoline fraction to form a low olefin gasoline product.

The main operating conditions and results are listed in table 3.

The SXC-6 catalyst is prepared by the following steps:

mixing 13.4g of HZSM-5 zeolite with 8.5g of alumina powder, uniformly mixing 15g of deionized water, adding a proper amount of dilute nitric acid solution, kneading and extruding into a strip with the diameter of 1.5 mm, airing at room temperature, drying at 120 ℃ for 4 hours, roasting at 540 ℃ for 3 hours, crushing and sieving into 0.6-0.9 mm particles to obtain HZSM-5-Al₂O₃And (3) a composite carrier. SiO of HZSM-5 zeolite used₂/Al₂O₃The molar ratio was 200.

The obtained composite carrier (11 g) was mixed with Ni (NO) (5.7 g)₃)₂﹒6H₂Soaking the solution prepared by O for 6h by conventional method, filtering, drying at 100 deg.C, and adding N₂Activating for 6 hours at 450 ℃ in atmosphere to obtain the catalyst of the invention, the specific surface of which is 320m²Pore volume was 0.26 ml/g.

The preparation process of the LXC-10 catalyst is as follows:

mixing amorphous aluminum silicate powder 12.3g and alumina powder 7.3g, adding deionized water 15g, mixing, and addingKneading and extruding a certain amount of dilute nitric acid solution into a strip-shaped object with the diameter of 1.5 mm, airing at room temperature, drying at 120 ℃ for 4 hours, roasting at 540 ℃ for 3 hours, crushing and sieving to obtain particles with the diameter of 0.6-0.9 mm to obtain amorphous aluminum silicate-Al₂O₃And (3) a composite carrier. SiO of amorphous aluminum silicate powder₂/Al₂O₃The molar ratio was 10.

The obtained composite carrier (10 g) was mixed with Ni (NO) (2.2 g)₃)₂﹒6H₂Soaking the solution prepared by O for 6h by conventional method, filtering, drying at 100 deg.C, and adding N₂Activating for 6 hours at 450 ℃ in atmosphere to obtain the catalyst of the invention, the specific surface of which is 310m²Pore volume was 0.31 ml/g.

The specific chemical composition properties of the SXC-6 catalyst and the LXC-10 catalyst are shown in tables 1-2.

Comparative example 1-1

Comparative examples 1-l illustrate the effect of cyclic cracking of a whole light gasoline fraction and a C4 fraction to increase propylene yield during a hydrocarbon recycle catalytic conversion process.

The first catalytic cracking catalyst is a first regenerated catalyst with the temperature of 680 ℃, enters the bottom of a riser reactor of the first catalytic reactor through a first catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. The heavy hydrocarbon oil raw material is atmospheric residue (the main properties are shown in table 2) which is heated to 350 ℃ by a preheating furnace and then mixed with atomized water vapor, and the mixture is sprayed into a first catalytic reactor through a feeding nozzle and contacts with a hot first regenerated catalyst to carry out catalytic conversion reaction. The reaction oil gas (first product) and the first carbon-deposited catalyst enter a settler from the outlet of a riser of the first catalytic reactor for rapid separation, and the first carbon-deposited catalyst enters a stripping zone for stripping. The second catalytic cracking catalyst is a second regenerated catalyst with the temperature of 680 ℃, enters the bottom of a riser reactor of the second catalytic reactor through a second catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. And C4 fraction and full-fraction light gasoline (9-130 ℃) sequentially enter a second catalytic reactor to contact with a hot second regenerated catalyst for catalytic reaction. The weight ratio of the C4 fraction to the heavy hydrocarbon oil feedstock is 0.2: 1, the weight ratio of the full-fraction light gasoline to the heavy hydrocarbon oil raw material is 0.2: 1. introducing the reaction oil mixture from the second catalytic reactor into a fluidized bed of a third catalytic reactor at the outlet of a riser for further reaction, separating oil gas (a second product) and a second carbon-deposited catalyst after the reaction in a settler through the fluidized bed reactor, and allowing the separated second carbon-deposited catalyst to enter a stripping zone for stripping under the action of gravity. And the first carbon-deposited catalyst and the second carbon-deposited catalyst after steam stripping enter a regenerator through a spent agent conveying pipe, are contacted with heated air and are regenerated at 700 ℃, and the obtained hot first regenerated catalyst and the second regenerated catalyst are respectively returned to the first catalytic reactor and the second catalytic reactor for recycling. Reaction oil gas from the first catalytic reactor and the third catalytic reactor is led out from the settler together, introduced into a product separation system for product separation to obtain gas products and various liquid products, and simultaneously, part of the gas products is separated to obtain C4 fraction, full fraction light gasoline and heavy gasoline fraction.

Part of the whole light gasoline fraction and the heavy gasoline fraction are mixed into a gasoline product with low olefin.

Wherein the reaction conditions of the second catalytic reactor and the third catalytic reactor were more severe than in example 1, the main operating conditions and results are shown in table 3. This comparative example 1-1 did not carry out a polymerization reaction.

Comparative examples 1 to 2

Comparative examples 1-2 illustrate the effect of cyclic cracking of a whole light gasoline fraction and a C4 fraction to increase propylene yield during a hydrocarbon recycle catalytic conversion process.

The reaction apparatus used was the same as in comparative example 1-1. The raw materials and the main experimental steps are the same as those of the comparative example 1-1, except that the reaction severity of the second catalytic reactor and the third catalytic reactor is greatly reduced, namely the same reaction conditions as those of the example 1 are adopted. The main operating conditions and results are listed in table 3. The comparative examples 1-2 did not undergo a polymerization reaction.

Comparative example 2

Comparative example 2 illustrates the effect of cyclic cracking of a full range light gasoline and a reformate (98-135 ℃) to increase propylene production during a hydrocarbon cyclic catalytic conversion process.

The reaction apparatus used was the same as in example 1.

The first catalytic cracking catalyst is a first regenerated catalyst with the temperature of 680 ℃, enters the bottom of a riser reactor of the first catalytic reactor through a first catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. The heavy hydrocarbon oil raw material is atmospheric residue (the main properties are shown in table 2) which is heated to 350 ℃ by a preheating furnace and then mixed with atomized water vapor, and the mixture is sprayed into a first catalytic reactor through a feeding nozzle and contacts with a hot first regenerated catalyst to carry out catalytic conversion reaction. The reaction oil gas (first product) and the first carbon-deposited catalyst enter a settler from the outlet of a riser of the first catalytic reactor for rapid separation, and the first carbon-deposited catalyst enters a stripping zone for stripping. The second catalytic cracking catalyst is a second regenerated catalyst with the temperature of 680 ℃, enters the bottom of a riser reactor of the second catalytic reactor through a second catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. And (3) sequentially feeding the full-fraction light gasoline (with the distillation range of 9-130 ℃) and the superimposed product (with the distillation range of 98-135 ℃) into a second catalytic reactor to contact with a hot second regenerated catalyst for catalytic reaction. The weight ratio of the full fraction light gasoline to the heavy hydrocarbon oil raw material is 0.2: 1, the weight ratio of the superimposed product to the heavy hydrocarbon oil raw material is 0.2: 1. introducing the reaction oil mixture from the second catalytic reactor into a fluidized bed of a third catalytic reactor at the outlet of a riser for further reaction, separating oil gas (a second product) and a second carbon-deposited catalyst after the reaction in a settler through the fluidized bed reactor, and allowing the separated second carbon-deposited catalyst to enter a stripping zone for stripping under the action of gravity. And the first carbon-deposited catalyst and the second carbon-deposited catalyst after steam stripping enter a regenerator through a spent agent conveying pipe, are contacted with heated air and are regenerated at 700 ℃, and the obtained hot first regenerated catalyst and the second regenerated catalyst are respectively returned to the first catalytic reactor and the second catalytic reactor for recycling. Reaction oil gas from the first catalytic reactor and the third catalytic reactor is led out from the settler together, introduced into a product separation system for product separation to obtain gas products and various liquid products, and simultaneously, part of the gas products is separated to obtain C4 fraction, full fraction light gasoline and heavy gasoline fraction.

And introducing the C4 fraction into a first superposition reaction zone of a superposition reactor to carry out superposition reaction to obtain a first superposition stream, and introducing the first superposition stream into a second product separation system to separate the first superposition stream into an unreacted C4 fraction and a superposition product. Wherein the first polymerization catalyst used is the above SXC-6 catalyst; the polymerization reactor is a fixed bed reactor; the first polymerization reaction conditions are as follows: the reaction temperature is 350 ℃, the reaction pressure is 1.0Mpa, and the weight hourly space velocity is 3.0h^-1。

The heavy gasoline fraction and part of the whole fraction light gasoline are mixed into a low-olefin gasoline product.

The main operating conditions and results are listed in table 3.

Example 2

Example 2 illustrates the effect of the process provided by the present disclosure on high selectivity propylene production during a hydrocarbon recycle catalytic conversion process.

The reaction apparatus used was the same as in example i. The main experimental steps are the same as example l, except that the heavy hydrocarbon oil raw material is hydrogenated wax oil. The main operating conditions and results are listed in table 3.

TABLE 1-1

Name of catalyst	OMT
		ChemistryProperty by weight%
Al₂O₃	54.6
		P₂O₅	2.31
RE₂O₃	0.75
		Physical Properties
Total pore volume, ml/g	0.19
		Micropore volume, ml/g	0.018
Specific surface area, m²/g	138
		Area of micropores, m²/g	103
Specific surface of substrate, m²/g	37
		Bulk density, g/ml	0.72
Particle size distribution,% by weight
		0～20μm	1.5
0～40μm	15.1
		0～80μm	58.2
0～110μm	76.3
		0～149μm	92.7
Cracking Activity, wt%	67

Tables 1 to 2

Name of catalyst	SXC-6	LXC-10
			Chemical property, weight%
Al₂O₃	30.4	36.2
			NiO	14.2	2.8
HZSM-5	55.4
			Amorphous aluminum silicate		61.0

TABLE 2

Raw oil name	Atmospheric residuum	Hydrogenated wax oil
			Density (20 ℃), kg/m³	891.6	889.7
The composition of elements%
			C	86.20	86.15
H	13.06	13.36
			S	0.28	0.11
N	0.29	0.13
			Basic nitrogen	922
Group composition of%
			Saturated hydrocarbons	59.0	75.4
Aromatic hydrocarbons	22.3	12.8
			Glue	18.3	11.7
Asphaltenes	0.4	0.1
			Residual carbon value,%)	5.44	1.36
Freezing point, DEG C	>50
			Refractive index, 70 deg.C	1.4848
Total acid value of mgKOH/g	0.44
			Average relative molecular mass	528
Metal content, mg/kg
			Fe	4.2	1.1
Ni	17.9	5.0
			Cu	<0.1	<0.1
V	0.2	0.4
			Na	0.3	0.4
Ca	0.7	0.3
			Zn	0.9
Reduced pressure volumetric distillation range, deg.C
			IBP	258.0
5％	365.9
			10％	388.7
30％	435.7
			50％	489.0
66.5％	569.4

TABLE 3

Examples of the invention	Example 1	Comparative examples 1 to 1	Comparative examples 1 to 2	Comparative example 2	Example 2
						Intermediate product Properties
First light gasoline cut distillation range, deg.C	9-60				9-60
						Second light gasoline cut distillation range, deg.C	60-100				60-100
Distillation range of the third light gasoline fraction, DEG C	100-130				100-130
						Distillation range of full-fraction light gasoline, DEG C	9-130	9-130	9-130	9-130	9-130
Distillation range of the superimposed product, DEG C	155-330			98-135	160-253
						Olefin content of the polymerization product,% by weight	98.1			98.3	87.8
Feed recycle ratio
						Weight ratio of full fraction light gasoline to heavy hydrocarbon oil		0.2:1	0.2:1	0.2:1
Weight ratio of C4 to heavy hydrocarbon oil		0.2:1	0.2:1
						Weight ratio of the second light gasoline fraction to the heavy hydrocarbon oil	0.05:1				0.05:1
Weight ratio of the polymerization product to the heavy hydrocarbon oil	0.35:1			0.2:1	0.35:1
						Reaction conditions of the first catalytic reactor:
riser outlet temperature,. deg.C	540.0	540.0	540.0	540.0	570
						Total riser reaction time in seconds	2.2	2.2	2.2	2.2	2.5
Catalyst to feedstock weight ratio	9.0	9.0	9.0	9.0	11
						Proportion of atomized water vapor, wt%	15.0	15.0	15.0	15.0	26
Temperature of the first catalytic cracking catalyst,. degree.C	680	680	680	680	680
						Reaction conditions of the second catalytic reactor:
riser outlet temperature,. deg.C	560.0	620.0	560.0	560.0	575.0
						Total riser reaction time in seconds	1.46	1.46	1.46	1.46	1.46
Catalyst to feedstock weight ratio	10.0	10.0	10.0	10.0	10.0
						Proportion of atomized water vapor, wt%	25.0	25.0	25.0	25.0	25.0
Temperature of the second catalytic cracking catalyst,. degree.C	680	680	680	680	680
						Reaction conditions of the third catalytic reactor:
reaction temperature of	555.0	615.0	555.0	555.0	570.0
						Mass weight hourly space velocity of the fluidized bed h^-1	10.0	10.0	10.0	10.0	12.0
Settler pressure, MPa (absolute pressure)	0.21	0.21	0.21	0.21	0.21
						The materials are balanced and heavy
Dry gas	5.52	8.20	5.62	6.67	7.23
						Liquefied gas	35.81	34.14	34.49	35.27	41.10
C5 gasoline (C5 ~ 221 degree C, TBP)	36.03	33.97	37.26	34.92	32.37
						Diesel oil (221 to 330 ℃ and TBP)	10.41	10.55	10.45	10.41	8.41
Heavy oil (>330℃,TBP)	3.59	3.57	3.55	3.54	2.48
						Coke	8.64	9.57	8.63	9.19	8.41
Total of	100.00	100.00	100.00	100.00	100.00
						Propylene yield, wt.%	26.54	20.14	16.70	20.39	30.04
C5 gasoline olefin content, weight%	9.01	18.12	26.35	25.81	8.19

As can be seen from table 3, example 1, using the process provided by the present disclosure, the yield of propylene produced by the hydrocarbon recycle catalytic conversion process with the heavy hydrocarbon oil feedstock as feed was 26.54 wt%, which is significantly higher than that of comparative examples 1-1 and 1-2; the yield of propylene produced by the cyclic catalytic conversion process of C4 fraction and full range light gasoline using the heavy hydrocarbon oil feedstock as feed in comparative examples 1-2 was 16.70 wt% using the same reaction conditions.

In comparative example 1-1, the yield of propylene prepared by the cyclic catalytic conversion method of the C4 fraction and the full-range light gasoline using the heavy hydrocarbon oil as the feed was 20.14 wt%, but the yield of dry gas and coke was greatly increased after the reaction severity was increased. Example 1 compared to comparative example 1-1, the dry gas and coke yields decreased by 2.68 wt% and 0.93 wt%, respectively, and the propylene yield increased by 6.40 wt%.

The yield of propylene prepared by the catalytic conversion method of the polymerization product and the whole fraction light gasoline by using the heavy hydrocarbon oil raw material as the feed in the comparative example 2 is 20.39 wt%, and compared with the comparative example 2, the yield of propylene is increased by 6.16 wt% in the example 1 by adopting the method provided by the disclosure, and the yields of dry gas and coke are respectively reduced by 1.15 wt% and 0.55 wt%. And the olefin content of the gasoline is only 9.01 percent by weight.

Example 2 using the process provided by the present disclosure, propylene was produced in a 30.04 wt% yield from a hydrocarbon recycle catalytic conversion process using a heavy hydrocarbon oil feedstock as feed. And the olefin content of the gasoline is only 8.19 weight percent.

The preferred embodiments of the present disclosure are described in detail with reference to the accompanying drawings, 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, in the foregoing embodiments, various features described in the above embodiments may be combined in any suitable manner, and in order to avoid unnecessary repetition, various combinations that are possible in the present disclosure are not described again.

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 process for the recycle cracking of hydrocarbons, the process comprising:

contacting the superimposed product and the second light gasoline fraction with a second catalytic cracking catalyst in a second catalytic reactor to perform a second catalytic reaction to obtain a second reaction mixture containing the second catalytic cracking catalyst; subjecting the second reaction mixture to a third catalytic reaction in a third catalytic reactor to obtain a third reaction mixture; and carrying out gas-agent separation on the third reaction mixture in the settler to obtain a second carbon-deposited catalyst and the second product.

2. The process according to claim 1, wherein the distillation range of the product of the superposition is from 155 to 330 ℃.

3. The process according to claim 2, wherein the distillation range of the superimposed product is 160-253 ℃.

4. The method of claim 2, wherein the reaction conditions of the first polymerization reaction are a temperature of 300-450 ℃, a pressure of 0.5-2.0 MPa, and a mass space velocity of 1-5 h^-1；

The reaction conditions of the second polymerization reaction are 180-280 ℃, the pressure of 1.0-2.0 MPa and the mass space velocity of 0.7-2.0 h^-1。

5. The method of claim 2, wherein the first polymerization catalyst comprises 1 to 20 mass% NiO, 40 to 80 mass% HZSM-5 zeolite, and 10 to 50 mass% alumina;

the second polymerization catalyst comprises 1-12 mass% of NiO, 45-82 mass% of amorphous aluminum silicate and 10-50 mass% of alumina.

6. The process of claim 2, wherein the first and second folding reactions are carried out in a folding reactor comprising at least two folding reaction zones, the first folding reaction being carried out in a first folding reaction zone of the folding reactor and the second folding reaction being carried out in a second folding reaction zone of the folding reactor.

7. The method according to claim 1, wherein the weight ratio of the superimposed product to the heavy hydrocarbon oil feedstock is from 0.01 to 0.6: 1.

8. the method according to claim 7, wherein the weight ratio of the superimposed product to the heavy hydrocarbon oil feedstock is from 0.05 to 0.3: 1.

9. the method of claim 1, wherein,

the distillation range of the first light gasoline fraction is 9-70 ℃;

the distillation range of the second light gasoline fraction is 50-110 ℃;

the distillation range of the third light gasoline fraction is 90-150 ℃.

10. The process according to claim 9, wherein the first light gasoline fraction has a distillation range of 9-60 ℃.

11. The process according to claim 9, wherein the second light gasoline fraction has a distillation range of 60 to 100 ℃.

12. The process according to claim 9, wherein the distillation range of the third light gasoline fraction is from 100 to 130 ℃.

13. The method of claim 1, wherein the operating conditions of the first catalytic reaction comprise: the reaction temperature is 480-600 ℃; the reaction time is 0.5-10 seconds; the weight ratio of the agent to the oil is 5-15: 1; the weight ratio of water to oil is 0.05-1: 1.

14. the method of claim 1, wherein the operating conditions of the second catalytic reaction comprise: the reaction temperature is 520-750 ℃; the reaction time is 0.1-3 seconds; the weight ratio of the agent to the oil is 6-40: 1; the weight ratio of water to oil is 0.1-1: 1.

15. the process of claim 1, wherein the third catalytic reactor is a fluidized bed reactor and the operating conditions of the third catalytic reactor are: the reaction temperature is 450-750 ℃; the weight hourly space velocity is 1-30 h^-1。

16. The method of claim 15, wherein the reaction temperature of the third catalytic reactor is 510-560 ℃.

17. The process of claim 1 wherein the first and second catalytic cracking catalysts each contain a shape selective zeolite having an average pore diameter of less than 0.7nm, the shape selective zeolite being at least one selected from the group consisting of zeolites having the MFI structure, ferrierites, chabazites, cyclospar, erionites, a-zeolites, epistillomites, and laumontites.

18. The method of claim 1, wherein the first catalytic reactor, the second catalytic reactor, and the third catalytic reactor are each one selected from a riser reactor, a downer reactor, a fluidized bed reactor, a riser and downer combined reactor, a riser and fluidized bed combined reactor, and a downer and fluidized bed combined reactor.

19. The method of claim 1, further comprising:

20. The method of claim 19, wherein the regeneration is performed at a temperature of 600-800 ℃.

21. The process as set forth in claim 19 wherein the first regenerated catalyst has a temperature of 560 ℃ and a carbon deposit content of 0.01 to 0.1% by weight, based on the dry weight of the first regenerated catalyst.

22. The process as set forth in claim 19 wherein the temperature of the second regenerated catalyst is 560 ℃ and 800 ℃, the carbon deposit content of the second regenerated catalyst is from 0.01 to 0.5% by weight, based on the dry weight of the second regenerated catalyst.

23. The process of claim 22, wherein the second regenerated catalyst has a carbon deposit content of 0.1 to 0.2 wt.%.

24. The process of claim 19, wherein the first coked catalyst is stripped in a first stripping zone of the settler and the second coked catalyst is stripped in a second stripping zone of the settler.

25. The process according to claim 1, wherein the heavy hydrocarbon oil feedstock is at least one selected from the group consisting of petroleum hydrocarbon oils, synthetic oils, coal liquefaction oils, oil sand oils, and shale oils.

26. The process of claim 25, wherein the heavy hydrocarbon oil feedstock is a petroleum hydrocarbon oil, which is at least one selected from the group consisting of atmospheric gas oil, vacuum gas oil, coker gas oil, deasphalted oil, hydrogenated tail oil, atmospheric residue, vacuum residue, and crude oil.

27. The process of claim 1 wherein introducing said second superimposed stream into a second product separation system for separation also produces an unreacted C4 fraction, an unreacted first light gasoline fraction and an unreacted third light gasoline fraction, and combining said unreacted first light gasoline fraction, said unreacted third light gasoline fraction, a portion of said second light gasoline fraction, and said heavy gasoline fraction to produce a gasoline product having a low olefin content.