A method for producing low-carbon olefins from a middle distillate oil

Abstract

The present application relates to a method for producing low-carbon olefins from middle distillate oil, which first converts the middle distillate oil into a naphtha fraction rich in alkanes through hydrofining and hydrocracking, then converts the naphtha fraction rich in alkanes into a fraction rich in olefins using an alkane dehydrogenation process; and separates the products of the alkane dehydrogenation reaction into alkanes and olefins, and introduces the fraction rich in olefins obtained by the separation into an olefin cracking reactor to react with a cracking catalyst, so as to convert the olefins into low-carbon olefins such as ethylene and propylene. The present application couples processes such as hydrofining, hydrocracking, alkane dehydrogenation, alkane-olefin separation and olefin cracking to selectively convert saturated chain alkanes into olefins, and then convert the olefins into low-carbon olefins, thereby achieving selective breaking of C-C bonds, improving the utilization rate of carbon atoms and increasing the yield of low-carbon olefins.

Description

TECHNICAL FIELD

[0001] The present application belongs to the field of catalytic cracking, and particularly relates to a method for preparing low-carbon olefins from middle distillate oil. BACKGROUND

[0002] Due to the rapid development of new energy vehicles, the consumption growth of domestic finished oil is slowing down, and the structural surplus of oil refining capacity is becoming increasingly obvious. In particular, the significant slowdown in diesel demand and the continuous decline in diesel prices have prompted refineries to face a new situation of reducing the diesel-gasoline ratio. At the same time, the production capacity of low-carbon olefins such as ethylene and propylene is in short supply, and it is imperative for refining and chemical enterprises to transform from fuel-type refineries to chemical-type refineries.

[0003] At present, the main technologies for producing low-carbon olefins are steam cracking and catalytic cracking. Compared with steam cracking, catalytic cracking has the characteristics of low reaction temperature, high propylene yield, low methane yield, good raw material adaptability and flexible production scheme, and it plays an important role in the field of low-carbon olefin production as an important secondary processing technology of crude oil. However, in view of the current slowdown in demand for middle distillate oil, how to reduce the surplus of middle distillate oil while maximizing the production of low-carbon olefins in strong demand has become a hot issue for refineries to upgrade and increase efficiency and increase refinery benefits.

[0004] In this regard, CN111607425A discloses a method for catalytic cracking of straight-run diesel oil, which introduces an acidic substance into the straight-run diesel oil to remove basic nitrogen compounds, so that it can be subjected to catalytic cracking reaction with pretreated ZSM-5 type nanomolecular sieve in the presence of inert gas to produce low-carbon olefins, thereby improving the yield of low-carbon olefins. However, there is still room for optimization in terms of product distribution, especially in reducing dry gas yield and increasing low-carbon olefin yield, using this method. SUMMARY

[0005] The purpose of the present application is to further improve the yield of low-carbon olefins in the process of preparing low-carbon olefins from middle distillate oil, and to reduce the yield of dry gas.

[0006] In order to achieve the above-mentioned purpose, the present application provides a method for preparing low-carbon olefins from middle distillate oil, which comprises the following steps:

[0007] (1) contacting the middle distillate oil and the first hydrogen donor with a hydrofining catalyst to perform a hydrofining reaction, to obtain a hydrofining middle distillate oil and a spent hydrofining catalyst;

[0008] (2) contacting the hydrofining middle distillate oil and the second hydrogen donor with a hydrocracking catalyst and performing a hydrocracking reaction, to obtain a naphtha fraction rich in alkanes and unconverted hydrofining middle distillate oil;

[0009] (3) contacting the alkane-rich naphtha fraction with an alkane dehydrogenation catalyst and reacting under alkane dehydrogenation reaction conditions to obtain hydrogen gas and a mixed hydrocarbon oil product;

[0010] (4) separating olefins in the mixed hydrocarbon oil product to obtain an olefin-rich fraction, and then contacting the olefin-rich fraction with a catalytic cracking catalyst to perform an olefin cracking reaction, and separating to obtain dry gas, liquefied gas and light hydrocarbon oil;

[0011] (5) separating the liquefied gas to obtain ethylene, propylene and C4+ olefins; and returning the separated C4+ olefins to step (4) as the olefin-rich fraction to perform the olefin cracking reaction.

[0012] Optionally, in step (1), the initial boiling point of the middle distillate oil is any temperature in the range of 180-230°C, and the final boiling point is any temperature in the range of 330-380°C.

[0013] Optionally, in step (1), the hydrofining catalyst comprises a first carrier and a metal component supported on the first carrier; the first carrier is selected from at least one of alumina and amorphous silica-alumina; the metal component is selected from at least one of a Group VIB metal and a Group VIII metal; the hydrofining catalyst comprises the following components based on the total weight of the hydrofining catalyst: 0-11 wt% of an additive, 1-38 wt% of a Group VIII metal, 1-51 wt% of a Group VIB metal, and the balance of the carrier; the additive is selected from at least one of fluorine, phosphorus, titanium and platinum; preferably, the conditions of the hydrofining reaction include: a reaction temperature of 235-455°C, preferably 335-425°C; a reaction pressure of 2.5-25.0 MPa, preferably 3.5-18.5 MPa; a volume ratio of the first hydrogen donor to the middle distillate oil of (230-2600): 1, preferably (220-2100): 1; a liquid hourly space velocity of the middle distillate oil of 0.05-20.0 h -1 -1, preferably 0.15-10.5 h -1 .

[0014] Optionally, in step (2), the content of C5-C9 saturated hydrocarbon components in the alkane-rich naphtha fraction is 50-100 wt%, preferably 55-100 wt%; the hydrocracking catalyst comprises a second carrier and an active component supported on the second carrier; the second carrier is silica-alumina and / or amorphous silica-alumina; the active component is a hydride component of one or more metals selected from Group VIB metals and / or Group VIII metals; the content of molecular sieve in the hydrocracking catalyst is 4-31 wt%, preferably 16-28 wt%; the molecular sieve is Y molecular sieve and / or β molecular sieve.

[0015] Optionally, in step (2), the conditions for the hydrocracking reaction include a reaction temperature of 335-425°C, preferably 355-395°C, a reaction pressure of 2.5-25.0 MPa, preferably 3.5-18.5 MPa; a volume ratio of the second hydrogen donor to the hydrofinished middle distillate oil of (220-2600): 1, preferably (310-2100): 1; a liquid hourly space velocity of the hydrofinished middle distillate oil of 0.05-20.0 h -1 , preferably 0.15-10.5 h -1 .

[0016] Optionally, the alkane dehydrogenation catalyst includes a third support and a dehydrogenation active component supported on the third support; the third support is selected from one or more of alumina, a porous zeolite, kaolin, ZSM-5, and SAPO-34 molecular sieve; the dehydrogenation active component is selected from one or more of platinum, palladium, ruthenium, vanadium, titanium, cerium, molybdenum, and niobium; the content of olefins in the mixed hydrocarbon oil product is 50-100 wt%, preferably 60-100 wt%.

[0017] Optionally, the conditions for the alkane dehydrogenation reaction include a reaction pressure of 0.03-2.0 MPa, preferably 0.05-2.5 MPa; a reaction temperature of 355-650°C, preferably 400-550°C; a mass hourly space velocity of the alkane-rich naphtha fraction of 0.5-50.0 h -1 , preferably 1.5-9.5 h -1 .

[0018] Optionally, the olefins in the mixed hydrocarbon oil product are separated using an alkylene separation adsorbent, which is preferably an X-type zeolite and / or a Y-type zeolite; the conditions for separating the olefins in the mixed hydrocarbon oil product using an alkylene separation adsorbent include a temperature of 85-455°C, preferably 105-225°C; a pressure of 0.5-3.0 MPa, preferably 1.0-2.5 MPa; a mass hourly space velocity of the mixed hydrocarbon oil product of 1.5-5.5 h -1 , preferably 2.5-4.5 h -1 .

[0019] Optionally, the conditions for the olefin cracking reaction include a reaction temperature of 550-750°C, preferably 580-720°C; a weight hourly space velocity of the olefin-rich fraction of 15-95 h -1 , preferably 35-85 h -1; the weight ratio of the catalytic cracking catalyst to the olefin-rich fraction is (3-25):1, preferably (5-20):1; the reactor for performing the olefin cracking reaction is selected from one or two or more combinations of the following: a riser, an isokinetic fluidized bed, an isodiameter or variable-diameter fluidized bed, an upflowing conveying line, a downflowing conveying line.

[0020] Optionally, the catalytic cracking catalyst comprises 1-50 wt% of zeolite, 5-99 wt% of inorganic oxide and 0-70 wt% of clay, based on the total weight of the catalytic cracking catalyst; the zeolite comprises mesoporous zeolite and / or small-pore SAPO zeolite, based on the total weight of the zeolite; the mesoporous MFI zeolite has a silicon-aluminum ratio of (50-400):1, preferably (80-300):1; the conversion rate of the olefin-rich fraction is 85-100 mol%, preferably 90-100 mol%.

[0021] By the above technical solution, the present application couples intermediate distillate oil hydrofining, hydrocracking, alkane dehydrogenation, alkane-olefin separation and olefin cracking processes, selectively converts saturated chain alkanes into olefins, and converts the olefins into low-carbon olefins, realizes selective breaking of C-C bonds, improves carbon atom utilization rate, and achieves the purpose of increasing low-carbon olefin production.

[0022] Other features and advantages of the present application will be described in detail in the following specific embodiments. BRIEF DESCRIPTION OF DRAWINGS

[0023] The accompanying drawings are included to provide a further understanding of the present application, and constitute a part of the specification, and are used together with the following detailed description to explain the present application, but do not constitute a limitation of the present application. In the drawings:

[0024] Figure 1 is a schematic diagram of one embodiment of the method provided by the present application.

[0025] DETAILED DESCRIPTION

[0026] 1 pipeline 2 hydrofining unit 3 hydrocracking unit 4 pipeline

[0027] 5 alkane dehydrogenation unit 6 pipeline 7 pipeline 8 alkane-olefin separation unit

[0028] 9 pipeline 10 pipeline 11 olefin cracking unit 12 pipeline

[0029] 13 pipeline 14 pipeline 15 olefin separation unit 16 pipeline

[0030] 17 pipeline DETAILED DESCRIPTION

[0031] The specific embodiments of the present application are described in detail below. It should be understood that the specific embodiments described herein are merely illustrative and explanatory in nature and are not intended to limit the present application.

[0032] The main component of the intermediate distillate oil is a chain alkane mixture 12 -C 24 The chain alkane mixture of the intermediate distillate oil, if a process of directly catalytically cracking the intermediate distillate oil into low-carbon olefins is used, due to the joint action of free radical mechanism and carbenium ion mechanism in the catalytic cracking reaction process, the C-C bond is difficult to selectively break, and there are problems such as too high dry gas yield and unreasonable utilization of carbon and hydrogen atoms. If the intermediate distillate oil is to be converted into valuable products, another method involves hydrocracking the intermediate distillate oil to produce naphtha, and the obtained naphtha can be used to produce high-value petrochemical products. However, the alkane dehydrogenation process of the naphtha fraction has been industrialized, and if the product of the hydrocracking of the intermediate distillate oil can be combined with the current situation of the market being unable to meet the demand for low-carbon olefins, the product is dehydrogenated into an olefin-rich fraction through alkane dehydrogenation, and then the olefin-rich fraction is cracked into low-carbon olefins through olefin cracking, compared with the prior art, the production of ethylene, propylene and other low-carbon olefins will be increased to a greater extent, the dry gas yield will be maximized, and the economic benefits of the refinery will be increased. Therefore, the present application provides a method for preparing low-carbon olefins from intermediate distillate oil.

[0033] The present application provides a method for preparing low-carbon olefins from intermediate distillate oil, which comprises the following steps:

[0034] (1) contacting the intermediate distillate oil and a first hydrogen donor with a hydrofining catalyst to perform a hydrofining reaction, to obtain a hydrofined intermediate distillate oil and a spent hydrofining catalyst;

[0035] (2) contacting the hydrofined intermediate distillate oil and a second hydrogen donor with a hydrocracking catalyst to perform a hydrocracking reaction, to obtain an alkane-rich naphtha fraction and unconverted hydrofined intermediate distillate oil;

[0036] (3) contacting the alkane-rich naphtha fraction with an alkane dehydrogenation catalyst and reacting under alkane dehydrogenation reaction conditions, to obtain hydrogen and a mixed hydrocarbon oil product;

[0037] (4) separating olefins in the mixed hydrocarbon oil product to obtain an olefin-rich fraction, and then contacting the olefin-rich fraction with a catalytic cracking catalyst to perform an olefin cracking reaction, to separate dry gas, liquefied gas and light hydrocarbon oil;

[0038] (5) separating the liquefied gas to obtain ethylene, propylene and C4+ olefins; and returning the separated C4+ olefins to step (4) to perform an olefin cracking reaction.

[0039] The present application realizes the method of high value-added utilization of middle distillate oil by coupling processes such as middle distillate oil hydrofining, hydrocracking, alkane dehydrogenation, alkane-alkene separation and olefin cracking, and high-selectively converting saturated chain alkanes into olefins, and then converting the olefins into low-carbon olefins, realizing the selective breaking of C-C bond, improving the carbon atom utilization rate, and achieving the purpose of increasing low-carbon olefin production.

[0040] The method of the present application realizes the maximum conversion of excess middle distillate oil into low-carbon olefins, while significantly reducing the dry gas yield, increasing the economic benefit of the refining and chemical enterprises while reducing the market share of the middle distillate oil.

[0041] It should be noted that the "C4 and above olefins" in the present application refer to olefins with a carbon atom number of 4 and above, which can be, for example, butene, 1-pentene, 1-hexene, 1-heptene, 1-octene, or a mixture of various olefins.

[0042] In step (1), the middle distillate oil is a diesel oil fraction obtained by conventional distillation of crude oil, and specifically, the initial boiling point of the middle distillate oil is any temperature in the range of 180-230 DEG C, and the final boiling point is any temperature in the range of 330-380 DEG C.

[0043] In step (1), the hydrofining catalyst comprises a first carrier and a metal component supported on the first carrier; the first carrier is selected from at least one of alumina and amorphous silica-alumina; the metal component is selected from at least one of Group VIB metal and Group VIII metal; in the specific embodiment of the present application, based on the total weight of the hydrofining catalyst, the hydrofining catalyst comprises the following components: 0-11 wt% of an additive, 1-38 wt% of Group VIII metal, 1-51 wt% of Group VIB metal, and the balance of the carrier; the additive is selected from at least one of fluorine, phosphorus, titanium and platinum.

[0044] Preferably, the conditions of the hydrofining reaction include: the reaction temperature is 235-455 DEG C, preferably 335-425 DEG C; the reaction pressure is 2.5-25.0 MPa, preferably 3.5-18.5 MPa; the volume ratio of the first hydrogen donor to the middle distillate oil is (230-2600): 1, preferably (220-2100): 1; the liquid hourly space velocity of the middle distillate oil is 0.05-20.0 h -1 -1, preferably 0.15-10.5 h -1 By hydrofining the middle distillate oil, the sulfur and nitrogen-containing substances in the middle distillate oil are removed, providing high-quality raw materials for subsequent processes.

[0045] In step (2), the content of C5-C9 saturated hydrocarbon components in the alkane-rich naphtha fraction is 50-100 wt%, preferably 55-100 wt%. Preferably, the unconverted hydrofinished middle distillate is recycled to the hydrocracking.

[0046] In the specific embodiment of the present application, the hydrocracking catalyst comprises a second support and an active component supported on the second support; the second support is silica-alumina and / or amorphous silica-alumina; the active component is a hydride component of one or more metals selected from the group consisting of Group VIB metals and / or Group VIII metals; the content of the molecular sieve in the hydrocracking catalyst is 4-31 wt%, preferably 16-28 wt%; the molecular sieve is Y molecular sieve and / or β molecular sieve. In addition, the molecular sieve can also be a molecular sieve modified according to the conventional method in the art.

[0047] In step (2), the conditions of the hydrocracking reaction include: the reaction temperature is 335-425℃, preferably 355-395℃, the reaction pressure is 2.5-25.0 MPa, preferably 3.5-18.5 MPa; the volume ratio of the second hydrogen donor to the hydrofinished middle distillate is (220-2600):1, preferably (310-2100):1; the liquid hourly space velocity of the hydrofinished middle distillate is 0.05-20.0 h -1 , preferably 0.15-10.5 h -1 .

[0048] In step (3), the alkane dehydrogenation catalyst comprises a third support and a dehydrogenation active component supported on the third support. The third support is preferably a support with regular or irregular pore structure, for example, the third support is selected from one or more of alumina, porous zeolite, kaolin, ZSM-5 and SAPO-34 molecular sieve; the dehydrogenation active component is selected from one or more of platinum, palladium, ruthenium, vanadium, titanium, cerium, molybdenum and niobium; the content of olefins in the mixed hydrocarbon oil product is 50-100 wt%, preferably 60-100 wt%, and the unconverted alkane is recycled to the dehydrogenation.

[0049] In step (3), the conditions of the alkane dehydrogenation reaction include: the reaction pressure is 0.03-2.0 MPa, preferably 0.05-2.5 MPa; the reaction temperature is 355-650℃, preferably 400-550℃; the mass hourly space velocity of the alkane-rich naphtha fraction is 0.5-50.0 h -1 , preferably 1.5-9.5 h -1 .

[0050] wherein, in step (4), the alkene in the mixed hydrocarbon oil product is separated by using an alkene separation adsorbent, and the alkene separation adsorbent is preferably X-type zeolite and / or Y-type zeolite; the conditions for separating the alkene in the mixed hydrocarbon oil product by using the alkene separation adsorbent include: a temperature of 85-455°C, preferably 105-225°C; a pressure of 0.5-3.0 MPa, preferably 1.0-2.5 MPa; a mass space velocity of the mixed hydrocarbon oil product of 1.5-5.5 h -1 , preferably 2.5-4.5 h -1 .

[0051] In the preferred embodiment of the present application, the alkene separation adsorbent is selected from the adsorbents commonly used in the art, such as X-type zeolite and / or Y-type zeolite, and the mixed hydrocarbon oil product is separated by using the alkene separation adsorbent, and the reaction stream obtained from the alkane dehydrogenation step is separated, and the stream rich in alkane is returned to the alkane dehydrogenation step for further reaction.

[0052] wherein, in step (4), the conditions for the olefin cracking reaction include: a reaction temperature of 550-750°C, preferably 580-720°C; a weight hourly space velocity of the fraction rich in olefin of 15-95 h -1 , preferably 35-85 h -1 ; and a weight ratio of the catalytic cracking catalyst to the fraction rich in olefin of (3-25): 1, preferably (5-20): 1.

[0053] In the specific embodiment of the present application, the reactor used for the olefin cracking reaction is selected from one or a combination of two or more of a riser, an isokinetic fluidized bed, an isodiameter or variable-diameter fluidized bed, an upflowing conveying line, and a downflowing conveying line.

[0054] wherein, in step (4), the catalytic cracking catalyst includes 1-50 wt% of zeolite, 5-99 wt% of inorganic oxide, and 0-70 wt% of clay, based on the total weight of the catalytic cracking catalyst; the zeolite includes mesoporous zeolite and / or small-pore SAPO zeolite, based on the total weight of the zeolite; the mesoporous MFI zeolite has a silicon-aluminum ratio of (50-400): 1, preferably (80-300): 1; and the conversion rate of the fraction rich in olefin is 85-100 mol%, preferably 90-100 mol%.

[0055] In the present application, the liquefied gas and the light hydrocarbon oil separated in step (4) can be introduced into an olefin separation device for separation, and ethylene, propylene, and C4+ olefin are obtained, and the C4+ olefin separated is returned to step (4) as the fraction rich in olefin for the olefin cracking reaction.

[0056] The present application provides a method for converting middle distillate oil into low carbon olefins, which first converts the middle distillate oil into an alkane-rich naphtha fraction through hydrofining and hydrocracking, and then converts the alkane-rich naphtha fraction into an olefin-rich fraction by using an alkane dehydrogenation process; and separates the product of the alkane dehydrogenation reaction into the olefin-rich fraction and the non-dehydrogenated alkane fraction, and then circulates the non-dehydrogenated alkane fraction to the alkane dehydrogenation reactor, and then introduces the olefin-rich fraction into an olefin cracking reactor to react with a cracking catalyst, so as to convert the olefins into low carbon olefins such as ethylene and propylene. The method of the present application not only can convert the middle distillate oil into low carbon olefins, but also can obtain a higher yield of low carbon olefins and a lower yield of dry gas, especially reduces the yield of methane in the product, and realizes the high value-added utilization of middle distillate oil resources.

[0057] It should be noted that "rich" in the present application can mean that the amount of one or a kind of compounds in the stream is usually at least 50 mol% or at least 65 mol%, preferably 90 mol%. In a broad sense, the "rich" refers to the fact that the outlet stream from the unit device has a greater percentage of a certain component than that present in the inlet feed of the unit device.

[0058] Figure 1 The flow of one specific embodiment of the method provided by the present application is shown as follows, Figure 1 As shown in the figure, the preheated middle distillate oil is mixed with hydrogen from pipeline 6 through pipeline 1 and then enters the hydrofining unit 2 to contact with the hydrofining catalyst to perform a hydrofining reaction, and the hydrofined middle distillate oil is obtained. The reaction conditions of the hydrofining are as follows: the reaction temperature is 235-455℃, preferably 335-425℃; the reaction pressure is 2.5-25.0 MPa, preferably 3.5-18.5 MPa; the volume ratio of the first hydrogen donor to the middle distillate oil is (230-2600):1, preferably (220-2100):1; the liquid hourly space velocity of the middle distillate oil is 0.05-20.0 h -1 , preferably 0.15-10.5 h -1 . The reaction oil gas leaves the hydrofining unit 2 and then downwardly enters the hydrocracking unit 3, and the hydrofined distillate oil contacts with the hydrocracking catalyst in the hydrocracking unit 3 to perform a hydrocracking reaction, and an alkane-rich naphtha fraction is obtained; the reaction conditions of the hydrocracking are as follows: the reaction temperature is 335-425℃, preferably 355-395℃, the reaction pressure is 2.5-25.0 MPa, preferably 3.5-18.5 MPa; the volume ratio of the second hydrogen donor to the hydrofined middle distillate oil is (220-2600):1, preferably (310-2100):1; the liquid hourly space velocity of the hydrofined middle distillate oil is 0.05-20.0 h -1, preferably 0.15-10.5h -1 The alkane-rich naphtha fraction is introduced into the alkane dehydrogenation reactor 5 through line 4 and is contacted with the dehydrogenation catalyst therein to perform alkane dehydrogenation reaction, thereby obtaining hydrogen, an olefin-rich fraction and unconverted naphtha fraction; the alkane dehydrogenation reaction is performed under the following conditions: the reaction pressure is 0.03-2.0 MPa, preferably 0.05-2.5 MPa; the reaction temperature is 355-650℃, preferably 400-550℃; the mass space velocity of the alkane-rich naphtha is 0.5-50.0h -1 , preferably 1.5-9.5h -1 .

[0059] The hydrogen in the alkane dehydrogenation reaction product is introduced into the hydrofining unit 2 through line 6, and the olefin-rich fraction and the unconverted naphtha fraction are introduced into the alkene separation unit through line 7 for separation; the separated unconverted naphtha fraction is returned to the dehydrogenation reactor 5 through line 9 for dehydrogenation reaction, and the olefin-rich fraction is introduced into the olefin cracking unit 11 through line 10 and is contacted with the catalytic cracking catalyst therein to perform olefin cracking reaction, thereby obtaining dry gas, liquefied gas, light hydrocarbon oil and coke and other products; the olefin cracking reaction is performed under the following conditions: the reaction temperature is 550-750℃, preferably 580-720℃; the weight hourly space velocity of the olefin-rich fraction is 15-95h -1 , preferably 35-85h -1 ; the weight ratio of the catalytic cracking catalyst to the olefin-containing fraction is (3-25): 1, preferably (5-20): 1. The dry gas in the olefin cracking product is introduced through line 12, the liquefied gas is introduced through line 13, and the light hydrocarbon oil is introduced through line 14 into the olefin separation device for separation, thereby obtaining low-carbon olefins and C4+ olefins. The low-carbon olefins are sent to the low-carbon olefin treatment device through line 16, and the C4+ olefins are returned to the olefin cracking unit 11 through line 17 for further catalytic cracking reaction.

[0060] The present application is further illustrated in detail by the following examples, but the present application is not limited by this. The raw materials used in the examples can be obtained by commercial channels. Among them, the catalytic conversion catalyst used in the examples and comparative examples is TCC-1, which is produced by Qilu Branch Company of Sinopec Catalyst Co., Ltd. The properties of the catalytic conversion catalyst TCC-1 are shown in Table 1.

[0061] Table 1 Properties of catalyst TCC-1

[0062]

[0063]

[0064] The properties of the intermediate distillate oil used in the examples and comparative examples are shown in Table 2.

[0065] Table 2 Properties of Middle Distillate Oils

[0066] Feed oil name Middle distillate oil Density, g / cm 3 ]] 0.8429 Initial boiling point, °C 180 Final boiling point, °C 350 Viscosity (20°C), mm 2 / s]] 5.955 Freezing point, °C -6 Saturated hydrocarbons, wt% 76.3 Aromatics, wt% 23.7 Carbon, % 86.48 Hydrogen, % 13.52 Sulfur, % 1.10 Nitrogen, mg / L 74

[0067] Example 1

[0068] The experiment in this embodiment is... Figure 1 The process is carried out on the unit device shown.

[0069] Middle distillate oil enters hydrorefining unit 2 via pipeline 1 and hydrogen from pipeline 6, where it contacts the hydrorefining catalyst for a hydrorefining reaction. The conditions for hydrorefining the middle distillate oil include: a reaction pressure of 15.0 MPa, a reaction temperature of 375 °C, and a hydrogen-to-oil volume ratio of 900 Nm³. 3 / m 3 Liquid hourly space velocity 2.5 h⁻¹ -1 .

[0070] The hydrorefined middle distillate and hydrogen obtained from the reaction are fed downwards into hydrocracking unit 3 and contacted with the hydrocracking catalyst therein to carry out a hydrocracking reaction. The conditions for the hydrocracking reaction include: a reaction pressure of 15.0 MPa, a reaction temperature of 375 °C, and a hydrogen-to-oil volume ratio of 900 Nm³. 3 / m 3 The liquid hourly space velocity is 2.5 h⁻¹. -1 .

[0071] The alkane-rich naphtha fraction obtained from the hydrocracking reaction enters the alkane dehydrogenation reaction unit 5 via pipeline 5 for alkane dehydrogenation. The conditions for the alkane dehydrogenation reaction include: a reaction pressure of 0.15 MPa, a reaction temperature of 520 °C, and a mass hourly space velocity (MSV) of 9 h⁻¹ for the alkane-rich naphtha. -1 .

[0072] Hydrogen gas obtained from the alkane dehydrogenation reaction is introduced into the hydrorefining unit via pipeline 6, and the mixed hydrocarbon oil product is introduced into the alkane-olefin separation unit via pipeline 7. The conditions for the alkane-olefin separation reaction include: reaction pressure of 1.0 MPa, reaction temperature of 120 °C, and space velocity of 3 h⁻¹. -1 The olefin-rich fraction obtained from the separation is fed into olefin cracking unit 11 and contacted with the catalytic cracking catalyst therein to carry out an olefin cracking reaction. The conditions for the olefin cracking reaction include: a reaction temperature of 620°C, a catalyst-to-oil ratio of 20, and a space velocity of 35 h⁻¹. -1 .

[0073] The integrated process obtained the olefin-rich fraction was tested on a pilot-scale catalytic cracking unit in a variable diameter fluidized bed reactor. The catalyst used in the catalytic cracking reaction was TCC-1. The reaction oil gas and spent catalyst were separated in a settler. The products were cut according to the distillation range in a separation system, thereby obtaining ethylene, propylene, butylene and light oil and the like. The specific reaction conditions and product distribution are shown in Table 3.

[0074] Example 2

[0075] The method of this example was performed with reference to the method of Example 1, which was different from Example 1 in that the reaction temperature was different. The specific reaction conditions and product distribution are shown in Table 3.

[0076] Comparative Example 1

[0077] This comparative example used middle distillate oil as the raw material and directly tested on a pilot-scale catalytic cracking unit in a variable diameter fluidized bed reactor. The catalytic cracking catalyst used was TCC-1 catalyst. The reaction oil gas and spent catalyst were separated in a settler. The reaction products were cut according to the distillation range in a separation system, thereby obtaining ethylene, propylene, butylene and light oil and the like. The specific reaction conditions and product distribution are shown in Table 3.

[0078] Comparative Example 2

[0079] This comparative example was performed with reference to the method of Comparative Example 1, which was different from Comparative Example 1 in that the reaction temperature was different. The specific reaction conditions and product distribution are shown in Table 3.

[0080] Table 3 Reaction conditions and product distribution of examples and comparative examples

[0081]

[0082]

[0083] As can be seen from the data of Examples 1-2 in Table 3, the method of the present application obtained ethylene yields of 26.36 wt% and 28.29 wt%, respectively, and propylene yields of 45.32 wt% and 44.74 wt%, respectively, by using middle distillate oil as the raw material and performing process integration. In Comparative Example 1 and Comparative Example 2, the ethylene yields obtained by directly catalytically converting middle distillate oil were 7.52 wt% and 9.42 wt%, respectively, and the propylene yields were 14.77 wt% and 13.06 wt%, respectively. The ethylene and propylene yields obtained in Examples 1 and 2 were significantly higher than the ethylene and propylene yields obtained by directly catalytically converting middle distillate oil in Comparative Example 1 and Comparative Example 2.

[0084] The methane yield in the reaction oil gas obtained in Example 1-2 is 2.29 wt% and 3.06 wt%, respectively, and the methane yield in the direct catalytic conversion of the middle distillate oil in Comparative Example 1 and Comparative Example 2 is 6.20 wt% and 8.65 wt%, respectively, so it can be seen that the methane yield in the product obtained by the method of the present application in Example 1 and Example 2 is significantly lower than the methane yield in the direct catalytic conversion of the middle distillate oil in Comparative Example 1 and Comparative Example 2.

[0085] The above describes the preferred embodiments of the present application in detail, but the present application is not limited to the specific details in the above-described embodiments, and various simple modifications can be made to the technical solutions of the present application within the technical concept of the present application, and these simple modifications all belong to the protection scope of the present application.

[0086] In addition, it should be noted that each specific technical feature described in the above-described specific embodiments can be combined in any appropriate manner without contradiction. In order to avoid unnecessary repetition, the present application will not describe various possible combinations again.

[0087] In addition, various different embodiments of the present application can also be combined in any manner without departing from the idea of the present application, and they should also be considered as disclosed by the present application.

Claims

1. A method for producing low-carbon olefins from a middle distillate, characterized by, The method comprises the following steps: (1) contacting the middle distillate oil and a first hydrogen donor with a hydrofining catalyst to perform a hydrofining reaction, to obtain a hydrofined middle distillate oil and a spent hydrofining catalyst; the initial boiling point of the middle distillate oil is any temperature within the range of 180-230℃, and the final boiling point is any temperature within the range of 330-380℃; (2) contacting the hydrofined middle distillate oil and a second hydrogen donor with a hydrocracking catalyst and performing a hydrocracking reaction, to obtain a naphtha fraction rich in alkanes and unconverted hydrofined middle distillate oil; (3) contacting the naphtha fraction rich in alkanes with an alkane dehydrogenation catalyst and reacting under alkane dehydrogenation reaction conditions, to obtain hydrogen and a mixed hydrocarbon oil product; (4) separating olefins in the mixed hydrocarbon oil product to obtain an olefin-rich fraction, and then contacting the olefin-rich fraction with a catalytic cracking catalyst to perform an olefin cracking reaction, to separate dry gas, liquefied gas and light hydrocarbon oil; (5) separating the liquefied gas to obtain ethylene, propylene and C4+ olefins; the separated C4+ olefins are returned to step (4) to perform an olefin cracking reaction as the olefin-rich fraction; In step (4), based on the total weight of the catalytic cracking catalyst, the catalytic cracking catalyst comprises 1-50wt% of zeolite, 5-99wt% of inorganic oxide and 0-70wt% of clay; based on the total weight of the zeolite, the zeolite comprises mesoporous MFI zeolite and / or microporous SAPO zeolite.

2. The method of claim 1, wherein, In step (1), the hydrofining catalyst comprises a first carrier and a metal component supported on the first carrier; the first carrier is selected from at least one of alumina and amorphous silica-alumina; the metal component is selected from at least one of Group VIB metal and Group VIII metal; Based on the total weight of the hydrofining catalyst, the hydrofining catalyst comprises the following components: 0-11wt% of an additive, 1-38wt% of Group VIII metal, 1-51wt% of Group VIB metal and the balance of the carrier; the additive is selected from at least one of fluorine, phosphorus, titanium and platinum.

3. The method of claim 2, wherein, The conditions of the hydrofining reaction include: a reaction temperature of 235-455 ℃; a reaction pressure of 2.5-25.0 MPa; a volume ratio of the first hydrogen donor to the intermediate distillate oil of (230-2600): 1; a liquid hourly space velocity of the intermediate distillate oil of 0.05-20.0 h -1 .

4. The method of claim 3, wherein, The conditions of the hydrofining reaction include: a reaction temperature of 335-425 ℃; a reaction pressure of 3.5-18.5 MPa; a volume ratio of the first hydrogen donor to the intermediate distillate oil of (220-2100): 1; a liquid hourly space velocity of the intermediate distillate oil of 0.15-10.5 h -1 .

5. The method of claim 1, wherein, In step (2), the content of C5-C9 saturated hydrocarbon components in the naphtha fraction rich in alkanes is 50-100wt%; The hydrocracking catalyst comprises a second carrier and an active component supported on the second carrier; the second carrier is silica-alumina and / or amorphous silica-alumina; the active component is a hydride component of one or more metals selected from Group VIB metal and / or Group VIII metal; The content of the molecular sieve in the hydrocracking catalyst is 4-31wt%; The molecular sieve is Y molecular sieve and / or β molecular sieve.

6. The method of claim 5, wherein, The content of C5-C9 saturated hydrocarbon components in the naphtha fraction rich in alkanes is 55-100wt%; the content of the molecular sieve in the hydrocracking catalyst is 16-28wt%.

7. The method of claim 1, wherein, In step (2), the conditions of the hydrocracking reaction include a reaction temperature of 335-425°C, a reaction pressure of 2.5-25.0 MPa; the volume ratio of the second hydrogen donor to the hydrorefined middle distillate oil is (220-2600): 1; the liquid hourly space velocity of the hydrorefined middle distillate oil is 0.05-20.0 h -1 .

8. The method of claim 7, wherein, The conditions of the hydrocracking reaction include: a reaction temperature of 355-395°C, a reaction pressure of 3.5-18.5 MPa; a volume ratio of the second hydrogen donor to the hydrofinished middle distillate oil of (310-2100):1; a liquid hourly space velocity of the hydrofinished middle distillate oil of 0.15-10.5 h -1 .

9. The method of claim 1, wherein, In step (3), the alkane dehydrogenation catalyst comprises a third carrier and a dehydrogenation active component supported on the third carrier; the third carrier is selected from one or more of alumina, porous zeolite, kaolin, ZSM-5 and SAPO-34 molecular sieve; and the dehydrogenation active component is selected from one or more of platinum, palladium, ruthenium, vanadium, titanium, cerium, molybdenum and niobium; The content of olefins in the mixed hydrocarbon oil product is 50-100 wt%.

10. The method of claim 9, wherein, The content of olefins in the mixed hydrocarbon oil product is 60-100 wt%.

11. The method of claim 1, wherein, In step (3), the conditions of the alkane dehydrogenation reaction include: the reaction pressure is 0.03-2.0 MPa; the reaction temperature is 355-650 ℃; the mass space velocity of the alkane-rich naphtha fraction is 0.5-50.0 h -1 .

12. The method of claim 11, wherein, The conditions of the alkane dehydrogenation reaction include: a reaction pressure of 0.05-2.5 MPa; a reaction temperature of 400-550 ℃; a mass space velocity of the alkane-rich naphtha fraction of 1.5-9.5 h -1 .

13. The method of claim 1, wherein, In step (4), an alkene separation adsorbent is used to separate the olefins in the mixed hydrocarbon oil product, and the alkene separation adsorbent is X-type zeolite and / or Y-type zeolite; The conditions for separating the olefins in the mixed hydrocarbon oil product using the alkene separation adsorbent include a temperature of 85-455°C; a pressure of 0.5 to 3.0 MPa; a mass space velocity of the mixed hydrocarbon oil product of 1.5 to 5.5 h -1 .

14. The method of claim 13, wherein, Conditions for separating olefins in the mixed hydrocarbon oil product using an alkylene separation adsorbent include a temperature of 105-225°C; a pressure of 1.0-2.5 MPa; a mass space velocity of the mixed hydrocarbon oil product of 2.5-4.5 h -1 .

15. The method of claim 1, wherein, In step (4), the conditions of the olefin cracking reaction include: the reaction temperature is 550-750°C; the weight hourly space velocity of the olefin-rich fraction is 15-95 h -1 ; the weight ratio of the catalytic cracking catalyst to the olefin-rich fraction is (3-25): 1; The reactor used for the olefin cracking reaction is selected from one or a combination of two or more of a riser, an isokinetic fluidized bed, an isodiameter or variable-diameter fluidized bed, an upflowing conveying line, a downflowing conveying line.

16. The method of claim 15, wherein, The conditions of the olefin cracking reaction include: a reaction temperature of 580-720°C; a weight hourly space velocity of the olefin-rich fraction of 35-85 h -1 ; and a weight ratio of the catalytic cracking catalyst to the olefin-rich fraction of (5-20):

1.

17. The method of claim 1, wherein, The mesoporous MFI zeolite has a silicon-aluminum ratio of (50-400):

1. The conversion rate of the olefin-rich fraction is 85-100 mole%.

18. The method of claim 17, wherein, The mesoporous MFI zeolite has a silicon-aluminum ratio of (80-300):

1. The conversion rate of the olefin-rich fraction is 90-100 mole%.

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