A hydrogenation process for producing ethylene cracking material from coking gasoline

Abstract

The present invention discloses a hydrogenation process for producing ethylene cracking material from coker gasoline, comprising the following steps: a hydrodesiliconization and deolefination reaction zone and a hydrodesulfurization reaction zone are arranged from bottom to top within a hydrogenation reactor; the hydrodesiliconization and deolefination reaction zone is provided with at least one fixed-bed catalyst bed, loaded with a silicon scavenger and a hydrodeolefination catalyst from bottom to top; and the hydrodesulfurization reaction zone is provided with at least one fixed-bed catalyst bed, loaded with a hydrodesulfurization catalyst; vaporized coker gasoline feedstock is mixed with hydrogen and introduced from the bottom of the reactor, where it contacts and reacts with the silicon scavenger, hydrodeolefination catalyst, and hydrodesulfurization catalyst to undergo desiliconization, deolefination, and hydrodesulfurization reactions. The reaction products are subjected to gas-liquid separation, stripping, and fractionation to obtain ethylene cracking material. The process of the present invention maximizes the effectiveness of various catalysts in the reaction, reduces energy consumption of the device, and extends the operating cycle of the coker gasoline hydrogenation device.

Description

Technical Field

[0001] The invention belongs to the field of petroleum refining and relates to a method for producing high-quality ethylene material from coker naphtha, and specifically to a hydrogenation process for producing ethylene cracking material from coker gasoline. Background Art

[0002] Delayed coking is one of the important means for refining companies to process low-quality heavy oil, but delayed coking products generally require further impurity removal. For example, coking naphtha contains high levels of impurities such as sulfur, nitrogen, olefins, gums, and silicon-containing compounds. If it is to be used as a raw material for units such as ethylene cracking, fertilizers, and reforming, it must first be hydrorefined and desiliconized. The silicon-containing compounds come from the silicon-containing defoaming agents added during the delayed coking process. The silicon in the silicon-containing defoaming agent will enter the reactor with the raw oil and deposit on the catalyst, causing the catalyst to be permanently poisoned and deactivated by silicon, and also resulting in a short single start-up cycle of the unit. Many documents such as "Direct Determination of Silicon Content in Coking Distillate Oil by ICP-OES" report that the main product generated after the thermal cracking of polydimethylsiloxane as a defoaming agent is cyclosiloxane.

[0003] Currently, the coker gasoline hydrogenation process involves a simple, one-step desiliconization, hydrodeolefinization, and desulfurization process. The coking reaction is a thermal cracking process, and the coker oil contains a high concentration of olefins. Conventional coker gasoline hydrogenation processes require that the reactants, upon entering the reactor, first react with a protective agent and a silicon scavenger to remove impurities and silicon, followed by hydrodeolefination and desulfurization, allowing the silicon scavenger to protect the main catalyst. Currently, widespread attention is focused on the pressure drop caused by diolefins. However, since olefin saturation facilitates desiliconization and desulfurization reactions and generates a high exotherm, even after low-temperature diolefin removal during coker gasoline hydrogenation, the catalyst bed must be fed at a lower inlet reaction temperature to avoid strong olefin exotherm and maintain a controllable reaction process. While the optimal reaction temperature for silicon scavenging catalysts is above 270°C, the initial temperature of coker gasoline hydrogenation is only around 200°C. This prevents the silicon scavenger from initially meeting the optimal reaction temperature, resulting in insufficient desiliconization capacity, premature silicon poisoning of the main catalyst, and a rapid decrease in activity, shortening the operating cycle.

[0004] CN103789020A discloses a coker gasoline hydrogenation method. Combining the advantages of both an ebullating bed and a fixed bed, circulating oil is used to dilute unsaturated hydrocarbons, such as diolefins, in the feedstock while simultaneously removing the heat of reaction. This prevents coking of the catalyst bed and reduces the reactor pressure differential. However, this method fails to mitigate the strong exothermic effects of olefins, resulting in inappropriate temperature matching. It also requires significant equipment modifications.

[0005] CN112852480A discloses a coking gasoline hydrogenation catalyst grading and hydrogenation treatment method, comprising: sequentially arranging a replaceable zone and a main reaction zone in series along the logistics direction, wherein the replaceable zone is sequentially loaded with a diene saturation catalyst in a first reactor and a silicon scavenger in a second reactor along the logistics direction, wherein the second reactor is a parallel dual reactor capable of online switching; and the main reaction zone is sequentially loaded with an arsenic removal agent and a gasoline hydrogenation catalyst along the logistics direction.

[0006] CN107955642A discloses a method and system for hydrogenating a mixture of C5 raffinate oil and coker gasoline. The method comprises: subjecting the C5 raffinate oil to a pre-hydrogenation reaction to saturate the diolefins in the C5 raffinate oil with hydrogen; and then using the pre-hydrogenation product as a diluent and hydrogenating it together with the coker gasoline.

[0007] The processes used in the aforementioned patents fail to consider the matching of reaction temperature rise with the desiliconization catalyst. They also fail to integrate the exothermic heat of the coker gasoline reaction with the active temperature of the silicon scavenger. Because coker gasoline hydrogenation is highly exothermic, the initial reaction temperature of the silicon scavenger is often too low to match the activity of the scavenger. Summary of the Invention

[0008] In response to the shortcomings of the existing technology, the present invention provides a hydrogenation process for producing ethylene cracking material from coker gasoline. The process of the present invention deeply couples the phase changes of the reactants, the reaction forms and various reactions, controls the depth of the desulfurization reaction, balances the reaction heat, maximizes the efficiency of various catalysts in the reaction, reduces the energy consumption of the device, and extends the operating cycle of the coker gasoline hydrogenation device.

[0009] The hydrogenation process for producing ethylene cracking material from coker gasoline of the present invention comprises the following steps: a hydrodesiliconization and deolefination reaction zone and a hydrodesulfurization reaction zone are arranged from bottom to top in a hydrogenation reactor; the hydrodesiliconization and deolefination reaction zone is provided with at least one fixed-bed catalyst bed layer, in which a silicon scavenger and a hydrodeolefination catalyst are loaded from bottom to top; and the hydrodesulfurization reaction zone is provided with at least one fixed-bed catalyst bed layer, in which a hydrodesulfurization catalyst is loaded; gasified coker gasoline feedstock is mixed with hydrogen and enters from the bottom of the reactor, passes through the hydrodesiliconization and deolefination reaction zone and the hydrodesulfurization reaction zone in sequence, and contacts and reacts with the silicon scavenger, the hydrodeolefination catalyst and the hydrodesulfurization catalyst to carry out desiliconization, deolefination and hydrodesulfurization reactions; and the reaction products are subjected to gas-liquid separation, steam stripping and fractionation to obtain ethylene cracking material.

[0010] In the process of the present invention, in the hydrodesiliconization and deolefination reaction zone, the volume ratio of the silicon scavenger and the hydrodeolefination catalyst is 10:90 to 95:5, preferably 35:65 to 80:20; wherein the silicon scavenger is well known to those skilled in the art. As one of the specific embodiments, the active metals of the silicon scavenging catalyst are oxides of Group VIB metals and oxides of Group VIII metals, and based on the total weight of the catalyst, the Group VIB metal is calculated as oxide at 2-15%, preferably 5-10%, and the Group VIII metal is calculated as oxide at 2-10%, preferably 4-8%; the pore volume of the silicon scavenging catalyst is 0.5-1.0 mL / g, preferably 0.7-1.0 mL / g; the specific surface area is 250-500 m 2 / g, preferably 350-500m 2 / g; the acid content of B is 0.05-0.3mmol / g, preferably 0.08-0.2mmol / g; the preparation method of the catalyst is well known in the art, and one or more of the impregnation method, co-extrusion method, and co-precipitation method can be used. The hydrodeolefination catalyst is well known to those skilled in the art. As one of the specific embodiments, the active metal of the hydrodeolefination catalyst is a metal oxide of Group VIB and / or a metal oxide of Group VIII, and the carrier is alumina or alumina modified with an auxiliary agent. Based on the weight of the hydrodeolefination catalyst, the content of the metal oxide of Group VIB is 5%-30%, preferably 5%-15%, and the content of the metal oxide of Group VIII is 1%-15%, preferably 2%-6%; the specific surface area is 100-500m 2 / g, preferably 300-500m 2 / g, and a pore volume of 0.3-1.2 mL / g, preferably 0.4-0.8 mL / g. Methods for preparing the catalyst are well known in the art, and one or more of impregnation, co-extrusion, and co-precipitation methods may be employed. The hydrodeolefination catalyst may be a suitable commercial catalyst selected according to the requirements of the process flow, or may be prepared according to existing methods, or may be a regenerated catalyst obtained by regenerating a deactivated catalyst.

[0011] In the process of the present invention, the reaction conditions of the hydrodesiliconization and deolefination reaction zone are as follows: the gas phase inlet reaction temperature is 270℃~350℃, preferably 280℃~320℃, the hydrogen partial pressure is 1.0MPa~16MPa, preferably 3MPa~10MPa, and the volume space velocity is 2h -1 ~16h -1 The volume ratio of hydrogen to oil is 100:1 to 2000:1.

[0012] In the process of the present invention, the hydrodesulfurization catalyst is generally supported by refractory porous oxides, such as aluminum oxide, silicon oxide, titanium oxide, and composite oxides or mixed oxides of several elements. Generally, non-acidic or weakly acidic materials are used as carriers, and the hydrogenation activity of the catalyst is one or more of W, Mo, Ni and Co. The content of hydrogenation active components in terms of oxides is generally 3wt% to 55wt%, preferably 10wt% to 45wt%. The shape is spherical or bar-shaped, with a spherical diameter of 0.04 to 10mm, preferably 0.04 to 5mm; the bar is 2 to 10mm in length, preferably 2 to 8mm, and a diameter of 1 to 6mm, preferably 1.5 to 3.5mm. The pore volume of the hydrodesulfurization catalyst is 0.4-1.0mL / g, preferably 0.5-0.8mL / g; the specific surface area is 200-400m 2 / g, preferably 220-320m 2 The hydrodesulfurization catalyst can be a commercial catalyst selected according to the needs of the process, or can be prepared according to existing methods, or can be a regenerated catalyst obtained by regenerating a deactivated catalyst.

[0013] In the process of the present invention, the sulfur removal rate of the gaseous effluent from the hydrodesiliconization and deolefination reaction zone is 10% to 50%.

[0014] In the process of the present invention, the hydrodesulfurization reaction zone is a gas phase upflow reaction process, and the hydrodesulfurization reaction process conditions are as follows: the inlet reaction temperature is 250-400°C, preferably 260-310°C, the hydrogen partial pressure is 1.0 MPa-16.0 MPa, preferably 3.0 MPa-10.0 MPa, and the volume space velocity is 0.3 h -1 ~16.0h -1 The volume ratio of hydrogen to oil is 100:1 to 2000:1.

[0015] In the process of the present invention, the volume ratio of the catalysts in the hydrodesulfurization reaction zone and the hydrodesiliconization and deolefination reaction zone is 20:80 to 80:20, preferably 35:65 to 65:35.

[0016] In the process of the present invention, the coker gasoline feedstock has an initial boiling point of 30-100°C, an end boiling point of 160-240°C, an olefin content of 15-60% by weight, a silicon content of 2-500µg / g, and a sulfur content of 5-15000µg / g, preferably 6-12000µg / g. At least 40-100% by weight of the coker gasoline feedstock is coker gasoline, and one or more of catalytic gasoline, ethylene cracking gasoline, and straight-run naphtha may also be mixed.

[0017] In the process of the present invention, circulating oil enters the reactor between the hydrodesiliconization and deolefination reaction zones and the hydrodesulfurization reaction zones (in the middle of the reactor), exits the reactor bottom, and circulates back to the middle of the reactor. The circulating oil does not participate in the reaction and serves as a heat carrier.

[0018] In the process of the present invention, the feed mass ratio of the liquid phase circulating oil to the coking gasoline is 15:1 to 1.5:1.

[0019] In the process of the present invention, the initial boiling point of the liquid phase circulating oil is 340-450°C, and is generally a heavy diesel fraction or a wax oil fraction, such as one or more of heavy diesel, vacuum wax oil, hydrogenated tail oil, coker wax oil, etc.

[0020] In the process of the present invention, the organic sulfur content in the hydrodesulfurization reaction product flowing out of the hydrodesulfurization reaction zone is 50µg / g to 600µg / g.

[0021] In the process of the present invention, the gas separated from the gas-liquid is subjected to hydrogen sulfide removal and then mixed with new hydrogen and raw materials.

[0022] The process of the present invention has the following advantages compared with the prior art:

[0023] 1. The existing technology for treating coker gasoline generally first performs hydrodesiliconization and then desulfurization. The purpose is to prevent the silicon in the crude oil from contaminating the desulfurization catalyst and causing silicon poisoning of the desulfurization catalyst. However, due to the influence of olefins in coker gasoline, matching the catalyst activity with the reaction temperature is a difficulty. The inventors broke the convention and gasified the coker gasoline and mixed it with hydrogen. A countercurrent reaction was carried out in a hydrodesiliconization and deolefination reactor. The gas phase product of the reaction was then subjected to a desulfurization reaction to control the depth of the desulfurization reaction. By making full use of the characteristics of the countercurrent bed reaction and controlling the reaction temperature of the silicon scavenger, the desiliconization, deolefination and desulfurization reactions were achieved. The reaction heat and desulfurization depth were effectively controlled to produce high-quality ethylene cracking materials and reduce the amount of sulfur injected into the ethylene unit.

[0024] 2. The coker gasoline hydrogenation process provided by the present invention is based on the fact that the optimal reaction temperature for the hydrodesiliconization reaction is above 270°C. The coker gasoline is vaporized and reacted with circulating oil in the hydrodesiliconization and deolefination reaction zone. Due to the countercurrent reaction, the coker gasoline enters the reaction zone at the highest reaction temperature of the process. As the gas phase rises and the reaction temperature decreases under the action of the circulating oil, the temperature in the desiliconization reaction zone is more easily controlled, achieving a perfect match between the activity of the silicon capture agent and the reactants. This is determined by the inherent characteristics of the countercurrent reaction, effectively controlling the reaction temperature of the silicon capture catalyst bed, effectively resolving the problem of reaction exotherm in the original process, which causes the reaction temperature to mismatch during the initial operation of the silicon capture catalyst.

[0025] 3. The present invention provides circulating oil that does not participate in the reaction and serves as a heat carrier. Under the action of the heat carrier oil, the coker gasoline undergoes a controllable desiliconization and deolefination reaction, reducing the reaction temperature rise before entering the desulfurization reaction zone for the hydrodesulfurization reaction. This avoids the strong exotherm that occurs after the concentrated reaction of a large amount of olefins in the coker gasoline, facilitates controlling the temperature rise, and artificially reduces the difficulty of the reaction operation. The reactor bed temperature shows a trend of first cooling and then heating. This process is gentle and highly controllable, and can effectively reduce thermal cracking caused by excessive temperatures, which can lead to a decrease in liquid yield and target product yield.

[0026] 4. The present invention provides a method for controlling the reaction process in the reactor, setting up a hydrodesulfurization reaction zone and a hydrodesiliconization and deolefination reaction zone, respectively installing a hydrodesulfurization catalyst, a silicon scavenger, and a hydrodeolefination catalyst, and controlling the reaction temperature to fully exert the desulfurization, deolefination, and desiliconization activities of the catalyst, thereby avoiding the problem of the conventional coker gasoline hydrogenation process in which the initial silicon scavenger bed reaction temperature is too low, resulting in the silicon scavenger being unable to function and causing premature activity loss of the hydrogenation catalyst, and can avoid passive over-desulfurization and reduce the sulfur injection amount of the downstream ethylene unit.

[0027] 5. This invention utilizes progressive desulfurization, focusing on desulfurization and desiliconization reactions. Silicon, olefins, and simple sulfides are first removed, and then the desulfurization depth is controlled to suppress desulfurization, avoid excessive desulfurization during the reaction, and effectively control the desulfurization depth. The desiliconization temperature requirement is aligned with the control of the desulfurization reaction, achieving efficient matching of the reaction performance of the desulfurization catalyst, the silicon capture catalyst, and the deolefination catalyst.

[0028] 6. The present invention can be used to slightly modify existing coker gasoline hydrogenation or other similar hydrogenation equipment, which reduces the modification cost. At the same time, the operating steps are slightly changed, which also reduces the operational risks. The coker gasoline hydrogenation process provided by the present invention has a simple process, convenient operation, good safety, environmental friendliness, and a long operating cycle. BRIEF DESCRIPTION OF THE DRAWINGS

[0029] Figure 1 It is a schematic diagram of the process flow of producing ethylene cracking material from coking gasoline according to the present invention. DETAILED DESCRIPTION

[0030] The method disclosed in the present invention is described in more detail below with reference to the accompanying drawings and specific embodiments. Figure 1 Many necessary equipment are omitted, such as oil pumps, valves, heating furnaces, etc.

[0031] like Figure 1As shown, after the coking gasoline is gasified, it is mixed with the new hydrogen in pipeline 2 and the circulating hydrogen in pipeline 3 through pipeline 1 and enters the reactor from the bottom of the hydrotreating reactor 5. The material passes through the lower bed and flows upward. The circulating oil passes through pipeline 4 and enters the middle of the reaction zone and flows downward to form a countercurrent to carry out hydrodesiliconization, hydrodesolefination and moderate desulfurization reactions. The circulating oil then flows out from the bottom of the reactor and returns to pipeline 4. The reaction effluent from the lower bed is still in a gaseous state and flows upward to enter the upper bed for hydrodesulfurization reaction. The reaction product flows out from the top of the reactor; enters the cold high fraction 7 through pipeline 6, and the liquid phase passes through pipeline 8 and enters the cold low fraction 9, and then enters the fractionation unit through pipeline 10; the circulating hydrogen separated by the cold high fraction 7 enters the circulating hydrogen desulfurization system through pipeline 11, and then returns to pipeline 3 through the circulating hydrogen compressor 12.

[0032] The present invention will be further described below in conjunction with preferred embodiments. It should be understood that the preferred embodiments described herein are only used to illustrate and explain the present invention and are not used to limit the present invention.

[0033] The hydrorefining catalyst used in the upper hydrodesulfurization reaction zone in the Examples and Comparative Examples of the present invention is FH-40C, developed by the Fushun Petrochemical Research Institute and produced by the Fushun Branch of Sinopec Catalyst Company. The silicon scavenger used in the lower hydrodesiliconization and deolefination reaction zone is FHRS-2, developed by the Fushun Petrochemical Research Institute and produced by the Fushun Branch of Sinopec Catalyst Company. The deolefination catalyst used in the lower hydrodesiliconization and deolefination reaction zone is regenerated FH-40A, developed by the Fushun Petrochemical Research Institute and produced by the Fushun Branch of Sinopec Catalyst Company. The FH-40C catalyst is supported on alumina and contains W-Mo-Ni as the active metal component. The FHRS-2 catalyst is supported on expanded pore alumina and acid-modified with Mo-Ni as the active metal component. The FH-40A catalyst is supported on amorphous silica-alumina and acid-modified with Mo-Ni as the active metal component. ICP was used to analyze the composition content of the raw oil and the product, and nuclear magnetic resonance 29Si MAS NMR analysis showed that the silicon types in the raw oil and the product were siloxane compounds such as dimethylsiloxane, trimethylsiloxane and tetramethylsiloxane.

[0034] The conventional coker gasoline hydrofining process used in the comparative example of the present invention is a trickle bed reactor, where coker gasoline is mixed with hydrogen and enters the reactor from the top for hydrodeolefination, desiliconization and desulfurization reactions. Figure 1 The raw oils used in the examples of the present invention and the comparative examples are shown in Table 1. The main operating process conditions and gasoline product properties in the examples of the present invention and the comparative examples are shown in Table 2.

[0035] Table 1 Properties of crude oil

[0036]

[0037] Table 2 Main operating process conditions and gasoline product properties at different times

[0038]

[0039] As can be seen from the data results of the Examples and Comparative Examples, the most significant feature of the method of the present invention is that it utilizes a conventional coker gasoline hydrorefining unit and a trickle-bed reactor with partial countercurrent reaction. By gasifying the feedstock, the phase changes of the reactants, the reaction patterns, and the various reactions are deeply coupled, thereby balancing the heat of reaction, changing the response of olefin exothermicity to the reaction temperature rise, controlling the reaction depth, and maximizing the effectiveness of the catalyst in each reaction zone. This reduces the energy and hydrogen consumption of the unit, improves the product liquid yield, reduces operational risks, and extends the operating cycle of the coker gasoline hydrorefining unit.

Claims

1. A coker gasoline hydrogenation process, characterized in that The invention comprises the following contents: a hydrodesiliconization and deolefination reaction zone and a hydrodesulfurization reaction zone are arranged from bottom to top in a hydrogenation reactor; at least one fixed-bed catalyst bed is arranged in the hydrodesiliconization and deolefination reaction zone, and a silicon scavenger and a hydrodeolefination catalyst are loaded from bottom to top; at least one fixed-bed catalyst bed is arranged in the hydrodesulfurization reaction zone, and a hydrodesulfurization catalyst is loaded; gasified coking gasoline feedstock is mixed with hydrogen and enters from the bottom of the hydrogenation reactor, passes through the hydrodesiliconization and deolefination reaction zone and the hydrodesulfurization reaction zone in sequence, and contacts and reacts with the silicon scavenger, the hydrodeolefination catalyst and the hydrodesulfurization catalyst to carry out desiliconization, deolefination and hydrodesulfurization reactions; after gas-liquid separation, the reaction products are stripped and fractionated to obtain ethylene cracking material; the hydrodesiliconization and deolefination reaction zone adopts a countercurrent bed reaction process, and the reaction conditions are: a gas phase inlet reaction temperature of 270°C to 350°C, a hydrogen partial pressure of 1.0MPa to 16MPa, and a volume space velocity of 2h -1 ~16h -1 The sulfur removal rate of the gaseous effluent from the hydrodesiliconization and deolefination reaction zone is 10% to 50%; the organic sulfur content of the hydrodesulfurization reaction product flowing out of the hydrodesulfurization reaction zone is 50µg / g to 600µg / g; the circulating oil enters the hydrogenation reactor from the position between the hydrodesiliconization and deolefination reaction zone and the hydrodesulfurization reaction zone and flows downward, flows out from the bottom of the hydrogenation reactor, and circulates back to the hydrogenation reactor.

2. The process according to claim 1, characterized in that: In the hydrodesiliconization and deolefination reaction zone, the volume ratio of the silicon scavenger to the hydrodeolefination catalyst is 10:90 to 95:

5.

3. The process according to claim 1, characterized in that: The active metals of the silicon scavenger are oxides of Group VIB metals and Group VIII metals. Based on the total weight of the catalyst, the Group VIB metals are calculated as oxides at 2%-15%, and the Group VIII metals are calculated as oxides at 2%-10%. The pore volume of the silicon scavenger is 0.5-1.0 mL / g, and the specific surface area is 250-500 m 2 / g, the B acid content is 0.05-0.3mmol / g; the active metal of the hydrodeolefination catalyst is a Group VIB metal oxide and / or a Group VIII metal oxide, the carrier is alumina or alumina modified with an additive, and the Group VIB metal oxide content is 5%-30% and the Group VIII metal oxide content is 1%-15% based on the weight of the hydrodeolefination catalyst; the specific surface area of ​​the hydrodeolefination catalyst is 100-500m 2 / g, and the pore volume is 0.3-1.2mL / g.

4. The process according to claim 1, wherein: The hydrodesulfurization reaction zone is a gas phase upflow reaction process, and the hydrodesulfurization reaction process conditions are as follows: inlet reaction temperature is 250-400 ° C, hydrogen partial pressure is 1.0 MPa-16.0 MPa, volume space velocity is 0.3h -1 ~16h -1 The volume ratio of hydrogen to oil is 100:1 to 2000:

1.

5. The process according to claim 1, characterized in that: The volume ratio of the catalysts in the hydrodesulfurization reaction zone and the hydrodesiliconization and deolefination reaction zone is 20:80 to 80:

20.

6. The process according to claim 1, characterized in that: The coker gasoline raw material has an initial boiling point of 30°C to 100°C, a final boiling point of 160°C to 240°C, an olefin content of 15wt% to 60wt%, a silicon content of 2µg / g to 500µg / g, and a sulfur content of 5000µg / g to 15000µg / g.

7. The process according to claim 1, characterized in that: At least 40 wt% to 100 wt% of the coker gasoline raw material is coker gasoline.

8. The process according to claim 1, characterized in that: The feed mass ratio of the circulating oil to the coking gasoline is 15:1 to 1.5:1; the initial boiling point of the circulating oil is 340 to 450° C., and the circulating oil is a heavy diesel fraction or a wax oil fraction.

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