A process for the production of chloro-phthalic anhydride

Abstract

The application discloses a chlorinated phthalic anhydride production method, and relates to the chemical industry field, and aims at solving the problem of repeatedly increasing and decreasing pressure and temperature during the production of chlorinated phthalic anhydride. The chlorinated phthalic anhydride production method comprises the following steps: under the action of a catalyst, taking chlorinated o-xylene and acetic acid as raw materials, taking air as an oxidant, and performing an oxidation reaction in an oxidation reactor to obtain a gas phase product and an oxidation reaction liquid; the oxidation reaction liquid at least contains chlorinated phthalic acid, and the gas phase product at least contains acetic acid; performing acetic acid separation on the gas phase product to obtain acetic acid and a purified reaction liquid; then performing dehydration on the purified reaction liquid to obtain a crude anhydride product; and at least two chlorinated phthalic anhydride monomers are separated from the crude anhydride product. The chlorinated phthalic anhydride production method is used for producing chlorinated phthalic anhydride monomers.

Description

Technical Field

[0001] This invention relates to the field of chemical engineering, and more particularly to a method for producing chlorophthalic anhydride. Background Technology

[0002] Chlorophthalic anhydride is an important organic chemical raw material, mainly used in the synthesis of high-performance polyimide engineering plastics. It is also used as an intermediate in the production of dyes, pharmaceuticals and pesticides, and has been widely applied in high-tech fields such as aviation, aerospace, and electronics.

[0003] There are many methods for synthesizing chlorophthalic anhydride. Among them, the most economical and promising method is the one using o-xylene as a raw material, which is chlorinated to obtain monochloroo-xylene, and then catalytically oxidized to chlorophthalic anhydride under liquid-phase conditions with air. This method generally employs batch oxidation or batch distillation processes, requiring repeated pressure and temperature adjustments to the oxidation tower, which can easily lead to equipment fatigue and damage, shortening its service life. Furthermore, incomplete oxidation reactions result in a large amount of partially oxidized products and byproducts, which are difficult to separate using distillation. Since the reflux ratios required to remove different fractions vary significantly during distillation, using the same batch distillation unit is both difficult to operate and energy-intensive.

[0004] In addition, halogens, which act as electron-withdrawing substituents, are present on the benzene ring of chloroxylene. Under similar conditions, the oxidation reaction of chloroxylene is more difficult than that of xylene (such as p-xylene), the oxidation reaction has lower selectivity, and there are more partial oxidation products and byproducts. These partial oxidation products and byproduct impurities cannot be cleanly separated from the acid anhydride even by traditional distillation or recrystallization. Summary of the Invention

[0005] The purpose of this invention is to provide a production method for a fully continuous flow, high-purity chlorophthalic anhydride production system.

[0006] To achieve the above objectives, the present invention provides a method for producing chlorophthalic anhydride, comprising:

[0007] An oxidation reaction is carried out in an oxidation reactor using chloro-o-xylene and acetic acid as raw materials and air as an oxidant under the action of a catalyst to obtain gaseous products and an oxidation reaction liquid. The oxidation reaction liquid contains at least chloro-o-phthalic acid, and the gaseous products contain at least acetic acid.

[0008] The gaseous product is separated by acetic acid to obtain acetic acid and a purified reaction solution. The purified reaction solution is then dehydrated to obtain a crude anhydride product. At least two chlorophthalic anhydride monomers are then separated from the crude anhydride product.

[0009] Compared with existing technologies, the method for producing chlorophthalic anhydride provided by this invention uses chloro-o-xylene and acetic acid as raw materials and air as an oxidant under the action of a catalyst in an oxidation reactor to obtain a gaseous product and an oxidation reaction liquid. The oxidation reaction liquid contains at least chlorophthalic acid, and the gaseous product contains at least acetic acid. This allows for the separation of a portion of the acetic acid and the product, thereby increasing the reaction rate and yield. Simultaneously, a post-treatment unit can be used to separate acetic acid from the gaseous product, obtaining acetic acid and a purified reaction liquid. The purified reaction liquid is then dehydrated in an anhydride-forming reactor to obtain crude anhydride. Furthermore, the anhydride-forming reaction liquid can be transported to a multi-stage separation unit to separate different types of chlorophthalic anhydride monomers. Each separation unit is used to separate different types of chlorophthalic anhydride monomers. Therefore, this embodiment of the invention employs a fully continuous chlorophthalic anhydride production method, eliminating the need for repeated pressure and temperature adjustments to the oxidation reactor, thus reducing the risk of damage during production and extending the equipment's lifespan. Attached Figure Description

[0010] The accompanying drawings, which are included to provide a further understanding of the invention and form part of this invention, illustrate exemplary embodiments of the invention and are used to explain the invention, but do not constitute an undue limitation of the invention. In the drawings:

[0011] Figure 1 A basic block diagram of the chlorophthalic anhydride production system in this embodiment is shown;

[0012] Figure 2 A flowchart of the production method of chlorophthalic anhydride in this embodiment is shown;

[0013] Figure 3 A basic block diagram of the post-processing unit in this embodiment is shown;

[0014] Figure 4 A flowchart of the production method in the post-processing stage of this embodiment is shown;

[0015] Figure 5 A basic block diagram of the exhaust gas treatment unit in this embodiment is shown;

[0016] Figure 6 A flowchart of the production method for the exhaust gas treatment stage in this embodiment is shown;

[0017] Figure 7 A basic block diagram of the exhaust gas condensation unit in this embodiment is shown;

[0018] Figure 8 A basic block diagram of the acetic acid separation unit in this embodiment is shown;

[0019] Figure 9A flowchart of the production method for the acetic acid separation stage in this embodiment is shown;

[0020] Figure 10 A basic block diagram of the anhydride-forming apparatus in this embodiment is shown;

[0021] Figure 11 A basic block diagram of the multi-stage separation unit in this embodiment is shown;

[0022] Figure 12 A flowchart of the production method for the product separation stage in this embodiment is shown;

[0023] Figure 13 A schematic diagram of the oxidation reactor in this embodiment is shown;

[0024] Figure 14 A schematic diagram of the anhydride-forming apparatus of this embodiment is shown;

[0025] Figure 15 A process flow diagram of the chlorophthalic anhydride production system in this embodiment is shown.

[0026] Figure label:

[0027] 100 - Oxidation reactor; 200 - Post-treatment unit; 210 - Tail gas treatment unit; 211 - Tail gas condensation unit; 2111 - First condenser; 2112 - Second condenser; 2113 - Third condenser; 212 - Tail gas absorption unit; 220 - Acetic acid separation unit; 221 - Flash tank; 222 - Stripping tower; 230 - Acetic acid recovery unit; 300 - Anhydride formation device; 310 - Anhydride formation reactor; 311 - Oxidation reaction liquid distributor; 312 - First stirrer; 313 - Second stirrer; 320 - Catalyst separation unit; 321 - Evaporator; 322 - Recovery container; 330 - Crude anhydride condenser; 340 - Crude anhydride storage tank; 350 - Anhydride condenser; 400 - Multi-stage separation unit; 410 - First separation unit; 420 - Second separation unit; 430 - Third separation unit; a - First feed pipe; b - Second feed pipe. Detailed Implementation

[0028] To make the technical problems to be solved, the technical solutions, and the beneficial effects of the present invention clearer, the present invention will be further described in detail below with reference to the accompanying drawings and embodiments. It should be understood that the specific embodiments described herein are merely illustrative of the present invention and are not intended to limit the present invention.

[0029] Furthermore, the terms "first" and "second" are used for descriptive purposes only and should not be construed as indicating or implying relative importance or implicitly specifying the number of indicated technical features. Thus, a feature defined as "first" or "second" may explicitly or implicitly include one or more of that feature. In the description of this invention, "a plurality of" means two or more, unless otherwise explicitly specified. "Several" means one or more, unless otherwise explicitly specified.

[0030] Chlorophthalic anhydride is a high-performance polyimide engineering plastic. It can be used to prepare chlorophthalic acid from o-xylene and air under catalysis, followed by dehydration to produce chlorophthalic anhydride. The preparation process of chlorophthalic anhydride often employs batch oxidation or batch distillation processes, both requiring repeated pressure and temperature adjustments to the oxidation tower, which can easily lead to equipment fatigue and damage, shortening its lifespan. Furthermore, incomplete oxidation reactions produce a large amount of partially oxidized reaction products and byproducts, which are difficult to separate using distillation. Since the reflux ratios required to remove different fractions vary significantly during distillation, using the same batch distillation unit is both difficult to operate and energy-intensive.

[0031] Traditional processes employ multi-stage, stirred reactors as oxidation reactors, each requiring a separate cooling system. Using multi-stage stirred anhydride-forming reactors results in a long and inefficient process, numerous pieces of equipment, and significant investment. Furthermore, measures are not taken to address water generation during the oxidation reaction. High water content weakens the oxidation reaction, leading to the formation of large amounts of byproducts and intermediates. Even slight fluctuations in the system can cause significant instability, and the process is technically unstable. The removal of acetic acid and water from the oxidation products, as well as the separation of acetic acid and water, are extremely energy-intensive. Moreover, energy recovery is not considered, as the oxidation of o-chloroxylene releases a large amount of heat (1170.32 kJ / mol for 3-chloroo-xylene and 1167.3 kJ / mol for 4-chloroo-xylene). Finally, the final product of this method is a mixture of 3-chlorophthalic anhydride and 4-chlorophthalic anhydride, failing to yield the high-purity chlorophthalic anhydride monomer required for industrial applications.

[0032] Meanwhile, due to the presence of halogens as electron-withdrawing substituents on the benzene ring of chloroxylene, the oxidation reaction of chloroxylene is more difficult than that of xylene (such as p-xylene) under similar conditions. The oxidation reaction has lower selectivity, more partial oxidation and more byproducts, including four chloro-2-benzo[C]furanones, four chlorobenzoic acids and phthalic acid. These partially oxidized and byproduct impurities cannot be cleanly separated from the acid anhydride even by traditional distillation or recrystallization.

[0033] To address the aforementioned problems, embodiments of the present invention provide a method for producing chlorophthalic anhydride, which offers a fully continuous production method with high product purity, low intermediate and by-product levels, high acetic acid recovery rate, and minimal equipment wear. This method is applicable to a chlorophthalic anhydride production system comprising an oxidation reactor, a post-treatment unit, an anhydride-forming reactor, and a multi-stage separation unit.

[0034] Figure 1 A basic block diagram of the chlorophthalic anhydride production system in an embodiment of the present invention is shown. Figure 1 As shown, the chlorophthalic anhydride production system of this invention includes: an oxidation reactor 100, a post-treatment unit 200, an anhydride-forming device 300, and a multi-stage separation unit 400, wherein the oxidation reactor 100, the post-treatment unit 200, the anhydride-forming device 300, and the multi-stage separation unit 400 are connected in sequence.

[0035] like Figure 1 As shown, the oxidation reactor 100 is connected to a first feed pipe a and a second feed pipe b, respectively. The first feed pipe a is used to transport chloro-o-xylene, acetic acid, and a catalyst, while the second feed pipe b is used to transport air. Therefore, chloro-o-xylene, acetic acid, and a catalyst can be supplied to the oxidation reactor 100 via the first feed pipe a, and air can be supplied to the oxidation reactor 100 via the second feed pipe b. Under the action of the catalyst, chloro-o-xylene and air undergo an oxidation reaction in the oxidation reactor 100 to obtain an oxidation reaction liquid containing chlorophthalic acid and gaseous products.

[0036] like Figure 1 As shown, the oxidation reactor 100 is connected to the post-treatment unit 200, which is connected to the anhydride-forming device 300. The oxidation reaction liquid can be post-treated by the post-treatment unit 200 and then transported to the anhydride-forming device 300. In the anhydride-forming device 300, chlorophthalic acid can be dehydrated and converted into at least two kinds of chlorophthalic anhydride monomers. The outlet of the oxidation reaction liquid of the anhydride-forming device 300 is connected to the inlet of the multi-stage separation unit 400. The multi-stage separation unit 400 includes multiple separation units connected in series. Each separation unit can separate a corresponding kind of chlorophthalic anhydride monomer. The different kinds of chlorophthalic anhydride monomers are separated by each separation unit.

[0037] Figure 2 A flowchart of a method for producing chlorophthalic anhydride according to an embodiment of the present invention is shown. Figure 2 As shown, the method for producing chlorophthalic anhydride provided by the present invention includes:

[0038] Step 201: Under the action of a catalyst, using chloro-o-xylene and acetic acid as raw materials and air as an oxidant, an oxidation reaction is carried out in an oxidation reactor to obtain gaseous products and an oxidation reaction liquid. The oxidation reaction liquid contains at least chloro-o-xylene, and the gaseous products contain at least acetic acid.

[0039] In practice, the first feed pipe a and the second feed pipe b can supply chloro-o-xylene, acetic acid and catalyst to the oxidation reactor 100 according to the actual reaction molar ratio, so that chloro-o-xylene and oxygen can undergo an oxidation reaction in the oxidation reactor 100 at a certain temperature, and chloro-o-xylene is almost completely converted into chlorophthalic acid, obtaining an oxidation reaction liquid and a gaseous product. The oxidation reaction liquid contains at least chlorophthalic acid and the gaseous product contains at least acetic acid.

[0040] For example: the air can be compressed air with an oxygen content of 21% v%, the catalyst can be at least one of cobalt acetate, manganese acetate and tetrabromoethane, the ion mass ratio of chloro-o-xylene, acetic acid, cobalt ions, manganese ions and bromide ions can be 1:(3-6):(0.002-0.006):(0.001-0.004):(0.0005-0.004), the mass ratio of air (molar content of oxygen 21%) to chloro-o-xylene is (3.6-5):1, the reaction temperature is 185℃-230℃, and the reaction pressure is 1.2MPa(G)-2.3MPa(G). The oxidation reaction solution may also contain chloro-o-xylene, acetic acid, a catalyst, and intermediate products. The water content in the oxidation reaction solution is 5 wt% to 10 wt%. The gaseous products may contain at least gaseous acetic acid and water vapor, and may also contain waste gases such as oxygen, nitrogen, carbon monoxide, and carbon dioxide. The oxygen volume content in the gaseous products is 3% to 7%. The acetic acid used in this invention is pure acetic acid, preferably acetic acid with a purity higher than 98 wt%, such as acetic acid with a purity of 98.5 wt% or 99 wt%. Acetic acid serves as both a reactant and a solvent in the reaction.

[0041] Step 202: Separate the gaseous product with acetic acid to obtain acetic acid and a purified reaction solution. Then, dehydrate the purified reaction solution to obtain a crude anhydride product. Separate at least two chlorophthalic anhydride monomers from the crude anhydride product.

[0042] In practice, the gaseous products can enter a post-treatment unit to separate acetic acid and waste gas. The oxidation reaction liquid can also enter the post-treatment unit for acetic acid separation. The separated oxidation reaction liquid then enters an anhydride-forming unit, where the chlorophthalic acid contained in the oxidation reaction liquid is dehydrated and converted into at least two types of chlorophthalic anhydride monomers. These different chlorophthalic anhydride monomers are then separated through a multi-stage separation unit. At this point, different types of chlorophthalic anhydride monomers can be obtained, such as 4-chlorophthalic anhydride and 3-chlorophthalic anhydride.

[0043] As can be seen, in the method for producing chlorophthalic anhydride provided by this invention, the oxidation reactor, post-treatment unit, anhydride-forming device, and multi-stage separation unit in the production system are sequentially connected. Under the action of a catalyst, using chloro-o-xylene and acetic acid as raw materials and air as the oxidant, an oxidation reaction is carried out in the oxidation reactor to obtain a gaseous product and an oxidation reaction liquid. The oxidation reaction liquid contains at least chlorophthalic acid, and the gaseous product contains at least acetic acid. This allows for the separation of a portion of the acetic acid and the product, thereby increasing the reaction rate and yield. Simultaneously, the oxidation reactor is connected to the post-treatment unit, which is connected to the anhydride-forming device. The post-treatment unit can be used to separate acetic acid from the gaseous product to obtain acetic acid and a purified reaction liquid. The purified reaction liquid is then dehydrated in the anhydride-forming reactor to obtain crude anhydride. Furthermore, since the anhydride-forming device is connected to the multi-stage separation unit, the anhydride-forming reaction liquid can be transported to the multi-stage separation unit. The multi-stage separation unit is used to separate different types of chlorophthalic anhydride monomers, with each separation unit separating different types of chlorophthalic anhydride monomers. Therefore, the embodiments of the present invention employ a fully continuous chlorophthalic anhydride production method, which eliminates the need for repeated pressure and temperature adjustments to the oxidation reactor, thereby reducing the risk of damage during production and extending the service life of the equipment.

[0044] Figure 3 A basic block diagram of the post-processing unit according to an embodiment of the present invention is shown, as follows: Figure 3 As shown, the post-processing unit 200 includes: a tail gas treatment unit 210, an acetic acid separation unit 220, and an acetic acid recovery unit 230. The gas phase component discharge port of the oxidation reactor 100 is connected to the inlet of the tail gas treatment unit 210, the oxidation reaction liquid outlet of the oxidation reactor 100 is connected to the inlet of the acetic acid separation unit 220, and the oxidation reaction liquid outlet of the acetic acid separation unit 220 is connected to the inlet of the anhydride forming device 300. The oxidation reaction liquid outlet of the tail gas treatment unit 210 and the gas outlet of the acetic acid separation unit 220 are respectively connected to the inlet of the acetic acid recovery unit 230, which can recover the acetic acid solution in the gas phase product to the acetic acid recovery unit 230. The acetic acid outlet of the acetic acid recovery unit 230 is connected to the first feed pipe a of the oxidation reactor 100, and the oxidation reaction liquid outlet of the tail gas treatment unit 210 is also connected to the reflux port of the oxidation reactor 100.

[0045] Figure 4 A flowchart of the production method for the post-processing stage in this embodiment is shown. For example... Figure 4 As shown, the gaseous product is separated by acetic acid to obtain acetic acid and a purified reaction solution, including:

[0046] Step 401: The gaseous product is condensed to obtain an acetic acid solution, and then the acetic acid in the acetic acid solution is separated and sent to the oxidation reactor.

[0047] Step 402: Separate the acetic acid from the oxidation reaction solution to obtain a purified reaction solution and a gaseous component containing acetic acid, and then separate the acetic acid from the gaseous component.

[0048] In specific implementation, the aforementioned oxidation reactor 100 can be a bubbling oxidation tower, and the acetic acid recovery unit 230 can be an acetic acid recovery tower. The gaseous products discharged from the oxidation reactor 100 enter the tail gas treatment unit 210 to separate the waste gas from the acetic acid gas. The waste gas can be discharged into the waste gas treatment system. After the acetic acid gas is condensed, most of it is refluxed back to the oxidation reactor 100 through valve control, and a small portion is recovered to the acetic acid recovery unit 230 to control the water mass content in the oxidation reactor 100 to be 5wt% to 10wt%, with a reflux ratio of (20 to 40):1. At the same time, the oxidation reaction liquid of the oxidation reactor 100 enters the acetic acid separation unit 220 to separate acetic acid. The acetic acid in the oxidation reaction liquid is recovered to the acetic acid recovery unit 230, where it is treated to obtain purified acetic acid. The purified acetic acid is then recovered to the oxidation reactor 100 for further reaction. Based on this, the acetic acid volatilized during the reaction and the acetic acid discharged from the reaction solution can be recovered and reused, which can reduce the amount of acetic acid used and thus reduce production costs.

[0049] For example, Figure 5 A basic block diagram of the exhaust gas treatment unit according to an embodiment of the present invention is shown, as follows: Figure 5 As shown in the embodiment of the present invention, the tail gas treatment unit 210 in the chlorophthalic anhydride production system includes a tail gas condensation unit 211 and a tail gas absorption unit 212. The tail gas discharge port of the oxidation reactor 100 is connected to the inlet of the tail gas condensation unit 211, the tail gas discharge port of the tail gas condensation unit 211 is connected to the inlet of the tail gas absorption unit 212, the oxidation reaction liquid outlet of the tail gas absorption unit 212 is connected to the inlet of the acetic acid recovery unit 230, and the oxidation reaction liquid outlet of the tail gas condensation unit 211 is divided into two paths, one path is connected to the reflux port of the oxidation reactor 100, and the other path is connected to the inlet of the acetic acid recovery unit 230.

[0050] Figure 6 A flowchart of the production method for the exhaust gas treatment stage in this embodiment is shown. Figure 6 As shown, the gaseous product is condensed to obtain an acetic acid solution, comprising:

[0051] Step 601: The gaseous product is condensed to obtain a first acetic acid solution and tail gas;

[0052] Step 602: Use the exhaust gas absorption unit to absorb the exhaust gas and separate the residual second acetic acid solution in the exhaust gas;

[0053] Step 603: Separate the acetic acid from the first acetic acid solution and the second acetic acid solution.

[0054] In practice, the gaseous components discharged from the oxidation reactor 100 enter the tail gas condensation unit 211 for condensation treatment. After condensing a portion of the acetic acid gas in the gaseous components into a first acetic acid solution, most of the condensed acetic acid solution and water vapor are returned to the oxidation reactor 100 via valve control, while a small portion is sent to the tail gas absorption unit 212. The uncondensed gas in the gaseous components is discharged into the tail gas absorption unit 212. The tail gas absorption unit 212 can be a high-pressure tail gas absorption tower, equipped with packing material and demineralized water. It can convert the residual acetic acid gas in the uncondensed gas into an acetic acid solution, which is then discharged to the acetic acid recovery unit for recycling, significantly reducing acetic acid consumption and production costs.

[0055] For example, Figure 7 A basic block diagram of the exhaust gas condensation unit according to an embodiment of the present invention is shown, as follows: Figure 7 As shown, the method for producing chlorophthalic anhydride provided in this embodiment of the invention uses a tail gas condensation unit 211 in the chlorophthalic anhydride production system, which includes a first condenser 2111, a second condenser 2112, and a third condenser 2113, which are connected in sequence.

[0056] The method for producing chlorophthalic anhydride according to an embodiment of the present invention further includes, after condensing the gaseous product to obtain an acetic acid solution, and before separating acetic acid from the acetic acid solution through an acetic acid recovery unit, dividing the acetic acid solution into a first acetic acid solution and a second acetic acid solution, refluxing the first acetic acid solution to the oxidation reactor, and sending the second acetic acid solution into the acetic acid recovery unit.

[0057] The oxidation reaction liquid outlets of the first condenser 2111, the second condenser 2112, and the third condenser 2113 are sequentially connected and divided into two paths: one path connects to the reflux port of the oxidation reactor 100, and the other path connects to the inlet of the acetic acid recovery unit 230. It should be understood that there can be multiple tail gas condensation units or just one; three are used here as an example.

[0058] In practical applications, the shell sides of the first condenser 2111 and the second condenser 2112 are connected to the boiler feedwater pipe and steam network, while the shell side of the third condenser 2113 is connected to the demineralized water. The tube-side gas phase inlet and outlet of the first condenser 2111, the second condenser 2112, and the third condenser 2113 are connected sequentially. The tube-side gas phase inlet and outlet of the third condenser 2113 are also connected to the tail gas absorption unit 212. The oxidation reaction liquid outlets of the first condenser 2111, the second condenser 2112, and the third condenser 2113 are connected to the gas phase component discharge pipe and the condensate return pipe of the oxidation reactor 100. The upper part of the tail gas absorption unit 212 is connected to the demineralized water inlet, the top outlet of the tail gas absorption unit 212 is connected to the tail gas treatment device of the entire plant, and the bottom outlet of the tail gas absorption unit 212 is connected to the acetic acid recovery unit 230.

[0059] In specific implementation, the gaseous components discharged from the top of the oxidation reactor 100 are sequentially passed through the first condenser 2111, the second condenser 2112, and the third condenser 2113, and the oxidation reaction liquid outlet is divided into two paths: one path is returned to the oxidation reactor 100, and the other path is sent to the acetic acid recovery unit 230. It should be understood that the oxidation reaction liquid in the outlet is an acetic acid solution, and the water content in the oxidation reactor 100 is controlled by valves to ensure that the water content is 5wt% to 10wt%, and the reflux ratio is (20 to 40):1. In this embodiment of the invention, the first condenser 2111 produces steam at a pressure of 0.4 MPa (G) to 0.6 MPa (G), the second condenser 2112 produces steam at a pressure of 0.1 MPa (G) to 0.3 MPa (G), and the third condenser 2113 cools the gaseous components with demineralized water. The temperatures of the gaseous components exiting the tubes of the first condenser 2111, the second condenser 2112, and the third condenser 2113 are 165°C to 180°C, 130°C to 155°C, and 50°C to 110°C, respectively. Based on this, since different components have different vaporization points, components that reach their vaporization point at a certain pressure can be vaporized. Different gaseous components can be vaporized separately through the first, second, and third condensers. Furthermore, the oxidized gaseous components exiting from the top of the oxidation reactor can recover heat in stages by passing through the first, second, and third condensers sequentially. The third condenser uses a preheating method for the demineralized water, enabling effective and rational cascade utilization of the reaction heat.

[0060] The gas phase component exiting the third condenser in this embodiment of the invention still contains a large amount of acetic acid. If this acetic acid is not recovered, it will result in a waste of resources and an increase in product costs. Therefore, a high-pressure tail gas absorption tower is set up. The demineralized water contained in the high-pressure tail gas absorption tower absorbs the acetic acid in the gas phase component. The resulting acetic acid aqueous solution is directly sent to the subsequent acetic acid recovery device for recovery, which greatly reduces the consumption of acetic acid and production costs.

[0061] In one alternative approach, Figure 8 A basic block diagram of the acetic acid separation unit according to an embodiment of the present invention is shown, as follows: Figure 8 As shown in the embodiment of the present invention, the method for producing chlorophthalic anhydride uses an acetic acid separation unit 220 in its chlorophthalic anhydride production system, which includes a flash tank 221 and a stripping tower 222. The outlet of the oxidation reaction liquid of the oxidation reactor 100 is connected to the inlet of the flash tank 221. The outlet of the oxidation reaction liquid of the flash tank 221 and the exhaust port of the anhydride forming device 300 are both connected to the inlet of the stripping tower 222. The gas outlets of the flash tank 221 and the stripping tower 222 are both connected to the inlet of the acetic acid recovery unit 230. The outlet of the oxidation reaction liquid of the stripping tower 222 is connected to the inlet of the anhydride forming device 300.

[0062] Figure 9 A flowchart of the production method for the acetic acid separation stage in this embodiment is shown. Figure 9 As shown, an acetic acid separation unit is used to separate acetic acid from the oxidation reaction solution to obtain a purified reaction solution and a gaseous component containing acetic acid, including:

[0063] Step 901: Separate the oxidizing reaction solution with acetic acid to obtain a first purified reaction solution and a first gas phase component containing acetic acid;

[0064] Step 902: Separate the first purified reaction solution with acetic acid to obtain a second purified reaction solution and a second gas phase component containing acetic acid;

[0065] Step 903: The second purified reaction solution is sent into the anhydride formation reactor to carry out the anhydride formation reaction.

[0066] In practice, the oxidation reaction liquid from the oxidation reactor 100 is fed into the flash tank 221 for atmospheric pressure flash evaporation. The high temperature and pressure energy of the oxidation products are fully utilized to vaporize the acetic acid, which accounts for about 40% of the total acetic acid. The temperature of the liquid after flash evaporation is 115℃~130℃. The acetic acid gas flashed out directly enters the acetic acid recovery unit 230 for distillation, saving the steam consumption of the subsequent stripping tower 222 and the steam consumption of the acetic acid recovery unit 230. Meanwhile, the flash-evaporated oxidation reaction liquid enters the stripping tower 222. The bottom of the stripping tower 222 is connected to a stripping tower reboiler. The stripping tower reboiler uses external forced circulation heating to heat and vaporize most of the remaining acetic acid and water. The stripping tower is also equipped with a built-in condenser near the top. The vaporized acetic acid vapor and water vapor are partially condensed by the built-in condenser to obtain purified acetic acid vapor and water vapor, which then enter the acetic acid recovery unit 230 for distillation. The oxidation reaction liquid inlet of the anhydride forming unit 300 is connected to the pump outlet of the stripping tower 222, so that the oxidation reaction liquid in the stripping tower 222 enters the anhydride forming unit 300. The stripping tower operates at atmospheric pressure, with a bottom temperature of 160℃~200℃ and a reflux ratio of 0.08~0.3. The acetic acid recovery tower separates acetic acid and water by distillation, operating at atmospheric pressure with a reflux ratio of 2~6. The acetic acid concentration separated at the bottom of the tower is 98wt%~99.9wt%, which is circulated into the oxidation reactor through the reflux pipe. The acetic acid content in the wastewater separated at the top of the tower is 0.1wt%~2wt%, which is sent to the wastewater treatment device. Non-condensable gases are sent to the factory's tail gas treatment device for treatment to meet standards before being discharged into the atmosphere. The acetic acid recovery tower can be a plate tower or a composite tower combining plates and packing, with a theoretical number of 27~45 plates.

[0067] In this embodiment of the invention, the acetic acid vapor exiting the top of the stripping tower does not need to be completely condensed into an oxidation reaction liquid. A portion of the gas directly enters the acetic acid recovery tower for distillation, saving steam consumption in the reboiler of the acetic acid recovery tower. At the same time, the top of the stripping tower is equipped with a built-in condenser, which can condense a small amount of acetic acid vapor to wash and purify the organic matter carried in the acetic acid vapor, resulting in high purity of the recovered acetic acid.

[0068] In one example, Figure 10 A basic block diagram of the anhydride-forming apparatus according to an embodiment of the present invention is shown, as follows: Figure 10As shown in the embodiment of the present invention, the method for producing chlorophthalic anhydride uses an anhydride-forming device 300 in its production system. This device includes an anhydride-forming reactor 310, a catalyst separation unit 320, and a crude anhydride condenser 330. The outlet of the oxidation reaction liquid from the anhydride-forming reactor 310 is connected to a multi-stage separation unit 400 via the catalyst separation unit 320 and the crude anhydride condenser 330. The catalyst separation unit 320 includes an evaporator 321 and a recovery container 322. The outlet of the oxidation reaction liquid from the anhydride-forming reactor 310 is connected to the inlet of the evaporator 321. The gas outlet of the evaporator 321 is connected to the multi-stage separation unit 400 via the crude anhydride condenser 330, and the gas outlet of the evaporator 321 is connected to the recovery container 322. The recovery container 322 can be a residual liquid storage tank for storing the separated catalyst and heavy component residual liquid. The evaporator 321 can be a scraped film evaporator.

[0069] The method for producing chlorophthalic anhydride according to this invention, when dehydrating the purified reaction solution in an anhydride-forming reactor to obtain a crude anhydride product, further includes: dehydrating a second purified reaction solution to obtain a crude anhydride product containing at least acetic acid and chlorophthalic anhydride; and sending the acetic acid separated from the crude anhydride product to a post-processing unit. The second purified reaction solution is the reaction solution after acetic acid separation in a stripping tower.

[0070] In practice, the oxidation reaction liquid exiting from the bottom of the stripping tower 222 enters the anhydride-forming reactor 310. The oxidation reaction liquid in the anhydride-forming reactor 310 undergoes evaporation and dehydration to convert chlorophthalic acid into chlorophthalic anhydride. The crude anhydride exiting the anhydride-forming reactor 310 enters the evaporator 321 of the catalyst separation unit 320. Since the outlet of the oxidation reaction liquid from the anhydride-forming reactor 310 is sequentially connected to the multi-stage separation unit 400 through the catalyst separation unit 320 and the crude anhydride condenser 330, and the crude anhydride condenser 330 is connected to the multi-stage separation unit 400 through the crude anhydride storage tank 340, the crude anhydride is continuously vaporized. The vaporized crude anhydride is then liquefied by the crude anhydride condenser 330 and enters the crude anhydride storage tank 340, thereby separating the catalyst and polycyclic aromatic hydrocarbon residues and storing them in the recovery container 322. The crude anhydride is then pumped into the multi-stage separation unit 400 for the separation of various phthalic anhydride monomers. It should be understood that the multi-stage separation unit 40 can be multiple chlorophthalic anhydride separation towers connected in series. The evaporator 321 operates at a pressure of 0 kPa(A) to 10 kPa(A) and a temperature of 170℃ to 250℃.

[0071] For example, Figure 11 A basic block diagram of the multi-stage separation unit according to an embodiment of the present invention is shown, as follows: Figure 11As shown, the method for producing chlorophthalic anhydride provided in this embodiment of the invention uses a multi-stage separation unit 400 in the chlorophthalic anhydride production system, which includes a first separation unit 410, a second separation unit 420, and a third separation unit 430. The crude anhydride storage tank in the anhydride forming device 300 is sequentially connected to the first separation unit 410, the second separation unit 420, and the third separation unit 430.

[0072] Figure 12 A flowchart of the production method for the product separation stage in this embodiment is shown. For example... Figure 12 As shown, at least two chlorophthalic anhydride monomers were separated from the crude anhydride product using a multi-stage separation unit, including:

[0073] Step 1201: Separate the light component from the crude anhydride using the first separation unit;

[0074] Step 1202: Separate 4-chlorophthalic anhydride from the crude anhydride product using the second separation unit;

[0075] Step 1203: Separate 3-chlorophthalic anhydride from the crude anhydride product using the third separation unit.

[0076] For example, the multi-stage separation unit 400 can be three chlorophthalic anhydride separation towers: a phthalic anhydride separation tower, a 4-chlorophthalic anhydride separation tower, and a 3-chlorophthalic anhydride separation tower. The phthalic anhydride separation tower has 35-50 theoretical plates, an operating pressure of 2-7 kPa, a top temperature of 178-185°C, and a reflux ratio of 220-300. The 4-chlorophthalic anhydride separation tower has 30-45 theoretical plates, an operating pressure of 0-10 kPa, a top temperature of 170-205°C, and a reflux ratio of 2-10. The 3-chlorophthalic anhydride separation tower has 20-35 theoretical plates, an operating pressure of 0-10 kPa, a top temperature of 200-230°C, and a reflux ratio of 1-10. The phthalic anhydride separation tower, the 4-chlorophthalic anhydride separation tower, and the 3-chlorophthalic anhydride separation tower all contain a reboiler. The reboiler can be a falling film reboiler or other types of reboilers, which are not limited here.

[0077] In practice, the crude anhydride is continuously separated into products in three separation towers. First, light components such as phthalic anhydride are removed in the phthalic anhydride separation tower. Then, 4-chlorophthalic anhydride is separated in the second tower, the 4-chlorophthalic anhydride separation tower, with a purity greater than 99.6 wt%. Finally, 3-chlorophthalic anhydride is separated in the 3-chlorophthalic anhydride separation tower, with a purity greater than 99.8 wt%. This embodiment of the invention, through multi-stage separation unit design and optimized distillation process parameters, can continuously obtain high-purity chlorophthalic anhydride monomer products with relatively low energy consumption.

[0078] In one alternative approach, Figure 13 A schematic diagram of the structure of the oxidation reactor according to an embodiment of the present invention is shown, as follows: Figure 13 As shown in the embodiment of the present invention, the method for producing chlorophthalic anhydride uses an oxidation reactor 100 in its production system. The reactor includes an outer cylinder 110, an inner cylinder 120, packing material 130, and a gas distributor 140. The packing material 130 includes a first packing material 131 and a second packing material 132. The gas distributor 140 includes a first gas distributor 141 and a second gas distributor 142. The first packing material 131 is located in the outer cylinder 110 near the top of the inner cylinder 120. The first gas distributor 141 is arranged around the outer bottom of the inner cylinder 120. The second packing material 132 and the second gas distributor 142 are both located in the inner bottom of the inner cylinder 120.

[0079] like Figure 13 As shown, the oxidation reactor 100 can be cylindrical, with the inner cylinder 120 having a diameter 0.4 to 0.8 times that of the outer cylinder 110, and the inner cylinder 120 having a height 0.5 to 0.9 times that of the liquid level in the outer cylinder 110. The outer cylinder 110 is divided into a water concentration section, a gas-liquid separation section, and a gas-liquid reaction section from top to bottom. The water concentration section accounts for 20 to 40% of the total height of the oxidation reactor, the gas-liquid separation section accounts for 15 to 25%, and the gas-liquid reaction section accounts for 40 to 70%. The water concentration section is equipped with a first packing material 131. A first gas distributor 141 is installed at the bottom of the gas-liquid reaction section. The first gas distributor 141 can be a loop gas distributor or a disc gas distributor. The gas distributor is connected to the compressed air inlet of the oxidation reactor. A first feed pipe a is located above the gas distributor on the gas-liquid reaction section, and a second feed pipe b is located near the gas distributor. The top of the oxidation reactor 100 is equipped with a tail gas outlet pipe c and a condensate return pipe d. The inner cylinder 120 is entirely a gas-liquid reaction section. A second gas distributor 142 is installed at the bottom of the inner cylinder 120. The second gas distributor 142 can be a loop gas distributor or a disc gas distributor. The upper middle part of the inner cylinder 120 is equipped with a second packing material 132, which can be wire mesh packing. The bottom of the inner cylinder 120 is equipped with a discharge pipe. In this embodiment, the theoretical number of stages for the packing or trays in the water extraction and concentration section is 1 to 4. The liquid-phase reaction residence time in the outer cylinder 110 is 60 to 90 minutes, and the liquid-phase reaction residence time (the residence time after oxidation, i.e., the reaction time for secondary oxidation) in the inner cylinder 120 is not less than 50 minutes. The production system of this invention uses packing material in the upper part of the outer and inner cylinders of the oxidation reactor to increase the gas-liquid contact area, thereby ensuring the complete oxidation of chloro-o-xylene and its intermediate products. This achieves higher conversion and yield, and significantly reduces partial oxidation and the generation of by-product impurities.

[0080] In specific implementation, the outer cylinder 110 is the main reaction area between the raw material o-chloroxylene and air. The raw material o-chloroxylene, catalyst and acetic acid are fed into the outer cylinder 110 through the first feed pipe a contained in the oxidation reactor 100. Compressed air is fed into the first gas distributor 141 contained in the outer cylinder 110 through the second feed pipe b contained in the oxidation reactor 100. Thus, the first gas distributor 141 is used to evenly distribute the air transported by the second feed pipe b to the outer cylinder 110. At this point, most of the chloro-o-xylene is converted into chlorophthalic acid in the outer cylinder 110. The reacted liquid flows downward from the upper part of the inner cylinder 120 in a near-horizontal plug flow manner. Then, compressed air is sent through the second feed pipe b in the oxidation reactor 100 into the second gas distributor 142 in the inner cylinder. The second gas distributor 142 evenly distributes the air supplied by the second feed pipe b into the inner cylinder 120, where it undergoes another oxidation reaction with the reaction liquid in the inner cylinder 120 to obtain the final oxidation reaction liquid. The final oxidation reaction liquid flows out from the bottom of the inner cylinder 120. The first and second gas distributors 141 increase the reaction area between the air and the chloro-o-xylene, thereby accelerating the reaction rate. It should be understood that both the oxidation reaction liquid in the outer cylinder 110 and the final oxidation reaction liquid include the raw material chloro-o-xylene, catalyst, acetic acid, water, chlorophthalic acid, and some intermediate products. The gas distributor is at least one of a ring-shaped gas distributor and a disc-shaped gas distributor.

[0081] When the upper part of the outer cylinder 110 is filled with the first packing 131 and the upper part of the inner cylinder 110 is filled with the second packing 132, since the feed liquid is an oxidation reaction liquid, the oxidation reaction liquid needs to contact with air. The packing can disperse the oxidation reaction liquid, thereby increasing the contact area between the oxidation reaction liquid and air, so that chloro-o-xylene and its intermediate products are fully oxidized to chlorophthalic acid, thereby achieving a higher conversion rate and yield, and thus significantly reducing the generation of some oxidation products and by-product impurities. In the embodiments of the present invention, the acetic acid and water vapor refluxed from the tail gas condensation unit to the oxidation reactor are packed with packing in the upper part of the outer cylinder of the oxidation reactor. The packing section allows the cold reflux liquid to come into countercurrent contact with the rising acetic acid vapor and water vapor in the outer cylinder, which reduces the volatilization of acetic acid and increases the content of discharged water vapor, making the reaction more continuous and controllable.

[0082] In one example, Figure 14 A schematic diagram of the anhydride-forming apparatus according to an embodiment of the present invention is shown, as follows: Figure 14As shown in the embodiment of the present invention, the method for producing chlorophthalic anhydride uses an anhydride-forming device 300 in the chlorophthalic anhydride production system, which further includes an oxidation reaction liquid distributor 311 and an anhydride-forming condenser 350. The oxidation reaction liquid distributor 311 is located inside the anhydride-forming reactor 310. The outlet of the oxidation reaction liquid of the post-treatment unit is connected to the inlet of the anhydride-forming reactor 310, and the exhaust port of the anhydride-forming reactor 310 is connected to the inlet of the anhydride-forming condenser 350. The outlet of the anhydride-forming condenser 350 is divided into two paths, one connected to the inlet of the post-treatment unit and the other connected to the inlet of the anhydride-forming reactor 310.

[0083] When the oxidation reaction liquid from the bottom of the stripping tower enters the anhydride-forming reactor 310, it first flows into the oxidation reaction liquid distributor 311, which has a heating function. At this time, the distributor 311 can evenly distribute the oxidation reaction liquid into the anhydride-forming reactor 310. The operating pressure of the anhydride-forming reactor 310 is atmospheric pressure, and the reaction temperature is 205℃~250℃. Inside the anhydride-forming reactor 310, the oxidation product chlorophthalic acid is dehydrated to chlorophthalic anhydride. Simultaneously, at this temperature, the reaction liquid in the anhydride-forming reactor 310 is in a gaseous state and enters the anhydride-forming condenser 350, whose temperature is 90℃~130℃.

[0084] Therefore, the organic matter in the gaseous reaction liquid can be condensed and refluxed into the anhydride-forming reactor. Acetic acid vapor and water vapor in the gaseous reaction liquid are refluxed back to the stripping tower through a reflux pipe. The acetic anhydride from the anhydride-forming reactor is pumped into the evaporator. In this embodiment of the invention, the acetic acid in the anhydride-forming reactor is recovered to the stripping tower, and can be further recovered to the oxidation reactor through the acetic acid recovery unit for recycling, reducing production costs. Simultaneously, the oxidation reaction liquid feed is first heated on the oxidation reaction liquid distributor, and then evenly sprayed onto the surface of the crude anhydride liquid in the reactor through the distributor, which facilitates the evaporation of acetic acid and water, resulting in a faster anhydride formation rate.

[0085] For example, such as Figure 14 As shown, the anhydride-forming reactor 310 also includes a first stirrer 312 and a second stirrer 313, which are used to stir the oxidation reaction liquid in the anhydride-forming reactor 310. The first stirrer 312 can be an axial flow stirrer, and the second stirrer 313 can be a radial flow stirrer. For example, after the oxidation reaction liquid distributor 311 uniformly distributes the oxidation reaction liquid into the anhydride-forming reactor 310, the axial flow stirrer and the radial flow stirrer can uniformly stir the oxidation reaction liquid, thereby increasing the reaction rate while ensuring uniform distribution of the oxidation reaction liquid within the anhydride-forming reactor 310.

[0086] Figure 15 A process flow diagram of the chlorophthalic anhydride production system in this embodiment is shown.

[0087] Example 1

[0088] 1) Oxidation reaction

[0089] A mixture of 67.5 kg / h of chloro-o-xylene (mass composition: 54.8% 4-chloro-o-xylene, 44.8% 3-chloro-o-xylene, 0.3% o-xylene, 0.1% dichloro-o-xylene), 284.9 kg / h of solvent acetic acid, and catalysts of 0.85 kg / h of cobalt acetate, 0.42 kg / h of manganese acetate, and 0.50 kg / h of tetrabromoethane is introduced into the gas-liquid reaction section of the outer cylinder of the bubble column reactor through the feed pipe. Simultaneously, 260.4 kg / h of compressed air (21% oxygen molar content) enters the gas-liquid reaction section of the outer cylinder of the bubble column reactor through the air inlet and a ring-shaped gas distributor. The air reacts with the chloro-o-xylene to produce chlorophthalic acid. The reacted liquid enters the inner cylinder and flows downwards from the top of the inner cylinder in a near-horizontal flow manner. Simultaneously, compressed air (21% oxygen molar content) at a rate of 13.0 kg / h enters from the bottom of the inner cylinder, passes through a ring-shaped gas distributor, and flows counter-currently with the liquid in the inner cylinder to undergo a secondary oxidation reaction. The oxidation reaction temperature is 227.6℃, the pressure is 2.2 MPa(G), the reaction residence time in the outer cylinder is 70 min, and the reaction residence time in the inner cylinder is 60 min. The water content of the reactants in the oxidation tower is 5.3 wt%, and the oxygen volume content of the tail gas is 4.36%. The final oxidation product is discharged from the bottom outlet pipe of the inner cylinder and sent to the flash tank of the acetic acid recovery unit. The oxidation tail gas enters the tail gas treatment unit from the top gas phase outlet of the bubbling tower.

[0090] In this embodiment, the bubbling tower reactor has an outer diameter of 700 mm, an inner diameter of 500 mm, an outer cylinder liquid level height of 2350 mm, an inner cylinder height of 1700 mm, an inner cylinder wire mesh packing height of 700 mm, a theoretical number of 4 packing plates in the upper part of the outer cylinder, and a total height of 8230 mm for the bubbling tower reactor (outer cylinder). The liquid level and equipment height baselines are both lower tangent end caps.

[0091] 2) Exhaust gas treatment

[0092] The oxidation tail gas exiting the top of the bubbling reactor undergoes sequential heat recovery through primary, secondary, and tertiary tail gas condensers. The primary condenser produces 0.4 MPa(G) steam as a byproduct, the secondary condenser produces 0.1 MPa(G) steam as a byproduct, and the tertiary condenser cools the tail gas with demineralized water at 40°C. The tail gas temperatures exiting the tube side of the tertiary condenser are 165°C, 140°C, and 102°C respectively. It then enters a high-pressure tail gas absorption tower, where the acetic acid in the tail gas is absorbed by the demineralized water. The high-pressure tail gas absorption tower is a packed tower with an operating pressure of 2.0 MPa(G), an operating temperature of 83–94°C, a theoretical number of trays of 3, and a demineralized water flow rate of 24.5 kg / h, producing an acetic acid aqueous solution with a concentration of 25.02 wt%, which is then sent to the subsequent acetic acid recovery tower. The acetic acid content in the tail gas exiting the top of the absorption tower is 6.1 × 10⁻⁴ ppm. 5 The exhaust gas enters the factory and is discharged after being treated to meet standards by the exhaust gas treatment device. The condensate from the three-stage condenser is an aqueous solution containing acetic acid. Most of it is returned to the oxidation tower, and a small portion is sent to the subsequent acetic acid recovery tower. The condensate reflux ratio is 24.

[0093] 3) Acetic acid recovery

[0094] The material exiting from the bottom of the oxidation tower's inner cylinder enters the flash tank for flash evaporation. The operating pressure is atmospheric pressure, and the operating temperature is 120.6℃. The acetic acid gas generated from the flash evaporation enters the acetic acid recovery tower for rectification, and the liquid material generated from the flash evaporation enters the stripping tower. The stripping tower is a packed tower, operating at atmospheric pressure, with a bottom temperature of 182℃, a top temperature of 114.2℃, a reflux ratio of 0.098, and a theoretical number of plates of 6. The reboiler uses external forced circulation heating to heat and vaporize most of the remaining acetic acid and water. The acetic acid and water vapor are partially condensed and purified by the built-in condenser at the top before entering the acetic acid recovery tower for rectification. The feed to the acetic acid recovery tower includes both gas and liquid phases: an aqueous acetic acid solution from the bottom of the high-pressure tail gas absorption tower enters from the top; condensate from the third-stage tail gas condenser enters from the bottom; flash vapor from the flash tank enters from the bottom; and acetic acid and water vapor from the top of the stripping tower enter from the bottom. Acetic acid recovery tower distills and separates acetic acid and water. The operating pressure is atmospheric pressure, the top temperature is 99.3℃, the reflux ratio is 3.12, the theoretical number of plates is 34, and the concentration of acetic acid separated at the bottom of the tower is 98.2wt%. It is recycled into the oxidation reactor through the recycling pipe. The wastewater separated at the top of the tower contains 0.1wt% acetic acid and is sent to the wastewater treatment device. The non-condensable gas is sent to the factory's tail gas treatment device for treatment to meet the standards before being discharged into the atmosphere.

[0095] 4) Anhydride formation

[0096] The oxidation products from the bottom of the stripping tower enter the anhydride-forming reactor, where the oxidation product, chlorophthalic acid, is dehydrated to form anhydride. The anhydride-forming reactor uses a double-layer agitator: an upper axial-flow impeller agitator and a lower radial-flow paddle agitator. A liquid distribution disc is located above the agitator shaft. The liquid feed first flows to the top of the disc and is then sprayed onto the liquid surface below. The anhydride-forming reactor operates at atmospheric pressure and a temperature of 240°C. A significant amount of organic matter is carried away by the evaporated acetic acid and water, which is partially condensed in a condenser at a temperature of 130°C–140°C. This condensed organic matter is returned to the anhydride-forming reactor. The acetic acid and water vapor are purified in the stripping tower and then sent to the acetic acid recovery tower for distillation to recover acetic acid. The crude anhydride from the anhydride-forming reactor enters a scraped-film evaporator, where it is continuously vaporized, separating the catalyst and polycyclic heavy component residue. The operating pressure is 5 kPa(A) and the operating temperature is 231°C. The residual liquid from evaporation is sent to the catalyst recovery unit to recover the catalyst. The crude anhydride vapor is condensed to 160°C by the crude anhydride condenser, and the liquid enters the crude anhydride storage tank for buffering, and then is pumped into the distillation unit for purification.

[0097] 5) Distillation

[0098] The crude anhydride underwent continuous product separation in three vacuum distillation columns. First, light components such as phthalic anhydride were removed in a phthalic anhydride separation column. Then, 4-chlorophthalic anhydride was separated in a second column, the 4-chlorophthalic anhydride separation column, and finally, 3-chlorophthalic anhydride was separated in the 3-chlorophthalic anhydride separation column. The operating pressure at the top of the phthalic anhydride separation column was 5 kPa(A), the temperature was 184.6℃, the reflux ratio was 230, and the theoretical plate number was 39. The product obtained at the top of the column was mainly a mixture containing phthalic anhydride, with a phthalic anhydride content of 70.9 wt%, chlorophthalic anhydride 27.8 wt%, and chlorobenzoic acid, etc., 1.3 wt%. The 4-chlorophthalic anhydride separation column operates at a top pressure of 2.1 kPa(A), a temperature of 174.6℃, a reflux ratio of 4.2, a theoretical plate number of 34, a 4-chlorophthalic anhydride yield of 44.2 kg / h, and a purity of 99.6 wt%. The 3-chlorophthalic anhydride separation column operates at a top pressure of 2.1 kPa(A), a temperature of 208℃, a reflux ratio of 3.8, a theoretical plate number of 28, a 3-chlorophthalic anhydride yield of 35.0 kg / h, and a purity of 99.8 wt%. All three vacuum distillation columns use falling film reboilers.

[0099] Example 2

[0100] 1) Oxidation reaction

[0101] A mixture of 122.2 kg / h of raw material chloro-o-xylene (mass composition: 54.93% 4-chloro-o-xylene, 45.02% 3-chloro-o-xylene, 0.05% o-xylene), 516 kg / h of solvent acetic acid (mass composition: 98.0% acetic acid, 2% water), and catalysts cobalt acetate 1.44 kg / h, manganese acetate 0.48 kg / h, and tetrabromoethane 1.2 kg / h is introduced into the gas-liquid reaction section of the outer cylinder of the bubble column reactor through the raw material feed pipe. Simultaneously, 476 kg / h of compressed air (oxygen molar content 21%) enters the gas-liquid reaction section of the outer cylinder of the bubble column reactor through the air inlet and a ring-shaped gas distributor, where it reacts with chloro-o-xylene to produce chloro-o-benzene. Diformic acid, after reaction, enters the inner cylinder and flows downward from the top of the inner cylinder in a near-horizontal plug flow manner. At the same time, 19 kg / h of compressed air (oxygen molar content 21%) enters from the bottom of the inner cylinder, passes through the annular gas distributor, and flows counterclockwise with the liquid in the inner cylinder to undergo a secondary oxidation reaction. The oxidation reaction temperature is 199℃, the pressure is 1.2 MPa(G), the reaction residence time in the outer cylinder is 90 min, and the reaction residence time in the inner cylinder is 60 min. The water content of the reactants in the oxidation tower is 5.8 wt%, and the oxygen content of the tail gas is 4.49% (volume). The final oxidation product is discharged from the bottom outlet pipe of the inner cylinder and sent to the flash tank of the acetic acid recovery unit. The oxidation tail gas enters the tail gas treatment unit from the top gas phase outlet of the bubbling tower.

[0102] In this embodiment, the bubbling tower reactor has an outer diameter of 1000mm, an inner diameter of 700mm, an outer cylinder liquid level height of 2450mm, an inner cylinder height of 1590mm, an inner cylinder wire mesh packing height of 800mm, a theoretical number of 3 packing plates in the upper part of the outer cylinder, and a total height of 9200mm for the bubbling tower reactor (outer cylinder). The liquid level and equipment height baselines are both tangent end caps.

[0103] 2) Exhaust gas treatment

[0104] The oxidation tail gas exiting the top of the bubbling reactor undergoes sequential heat recovery through primary, secondary, and tertiary tail gas condensers. The primary condenser produces 0.4 MPa(G) steam as a byproduct, the secondary condenser produces 0.1 MPa(G) steam as a byproduct, and the tertiary condenser cools the tail gas with 40°C demineralized water. The tail gas temperatures exiting the tube side of the tertiary condenser are 160°C, 135°C, and 60°C respectively. It then enters the high-pressure tail gas absorption tower, where the acetic acid in the tail gas is absorbed by the demineralized water. The high-pressure tail gas absorption tower is a packed tower. The operating pressure is 1.0 MPa(G), the operating temperature is 53-56℃, the theoretical number of plates is 6, the demineralized water consumption is 15 kg / h, and the acetic acid aqueous solution with a concentration of 23.3 wt% is prepared and sent to the subsequent acetic acid recovery tower. The acetic acid content in the tail gas coming out of the top of the absorption tower is less than 1 PPM. After being treated to meet the standards by the tail gas treatment device in the factory, it is discharged into the atmosphere. The condensate from the three-stage condenser is an aqueous solution containing acetic acid. Most of it is refluxed into the oxidation tower, and a small portion is sent to the subsequent acetic acid recovery tower. The condensate reflux ratio is 32.

[0105] 3) Acetic acid recovery

[0106] The material exiting from the bottom of the inner cylinder of the oxidation tower enters the flash tank for flash evaporation. The operating pressure is atmospheric pressure and the operating temperature is 118.4℃. The acetic acid gas generated by flash evaporation enters the acetic acid recovery tower for rectification, and the liquid material generated by flash evaporation enters the stripping tower. The stripping tower is a composite tower with upper packing and lower sieve plates. The operating pressure is atmospheric pressure, the bottom temperature is 190℃, the top temperature is 114℃, the reflux ratio is 0.25, and the theoretical number of plates is 6. The reboiler uses external forced circulation heating to heat and vaporize most of the remaining acetic acid and water. After partial condensation and purification by the upper built-in condenser, the acetic acid and water vapor enter the acetic acid recovery tower for rectification. The acetic acid recovery tower operates at atmospheric pressure, with a top temperature of 99.9℃, a reflux ratio of 4.56, and a theoretical number of trays of 30. The acetic acid concentration separated at the bottom of the tower is 99.0 wt%, which is circulated into the oxidation reactor through the recycling pipe. The wastewater separated at the top of the tower contains 0.3 wt% acetic acid and is sent to the wastewater treatment unit. The non-condensable gas is sent to the factory's tail gas treatment unit for treatment to meet standards before being discharged into the atmosphere.

[0107] 4) Anhydride formation

[0108] The oxidation products from the bottom of the stripping tower enter the anhydride-forming reactor, where the oxidation product, chlorophthalic acid, is dehydrated to form anhydride. The anhydride-forming reactor uses a double-layer agitator: an upper axial-flow impeller agitator and a lower radial-flow paddle agitator. A liquid distribution disc is located above the agitator shaft. The liquid feed first flows to the top of the disc and is then sprayed onto the liquid surface below. The anhydride-forming reactor operates at atmospheric pressure and a temperature of 230°C. A significant amount of organic matter is carried away by the evaporated acetic acid and water, which is partially condensed in a condenser at 130°C. The condensed organic matter is returned to the anhydride-forming reactor. The acetic acid and water vapor are purified in the stripping tower and then sent to the acetic acid recovery tower for distillation to recover acetic acid. The crude anhydride from the anhydride-forming reactor enters a scraped film evaporator, where it is continuously vaporized to separate the catalyst and polycyclic heavy component residue. The operating pressure is 2.1 kPa(A) and the operating temperature is 210℃. The evaporation residue is sent to a catalyst recovery unit to recover the catalyst. The crude anhydride vapor is condensed to 90℃ in a crude anhydride condenser, and the liquid enters a crude anhydride storage tank for buffering before being pumped into a distillation unit for purification.

[0109] 5) Distillation

[0110] The crude anhydride undergoes continuous product separation in three vacuum distillation columns. First, light components such as phthalic anhydride are removed in a phthalic anhydride separation column. Then, 4-chlorophthalic anhydride is separated in a second column, namely the 4-chlorophthalic anhydride separation column. Finally, 3-chlorophthalic anhydride is separated in the 3-chlorophthalic anhydride separation column. The phthalic anhydride separation column operates at a top pressure of 5 kPa(A), a temperature of 183.5℃, a reflux ratio of 270, and 40 theoretical plates. The product obtained at the top is mainly a mixture containing phthalic anhydride, with a phthalic anhydride content of 86.8 wt%, chlorophthalic anhydride of 12.6 wt%, and chlorobenzoic acid, etc., of 0.6 wt%. The 4-chlorophthalic anhydride separation column operates at a top pressure of 2.1 kPa(A), a temperature of 174.4℃, a reflux ratio of 5, and 38 theoretical plates. The yield of 4-chlorophthalic anhydride is 80.9 kg / h, and the purity is 99.8 wt%. The 3-chlorophthalic anhydride separation column operates at a top pressure of 2.1 kPa(A), a temperature of 208.3℃, a reflux ratio of 4.5, and 35 theoretical plates. The yield of 3-chlorophthalic anhydride is 62.1 kg / h, and the purity is 99.8 wt%. All three vacuum distillation columns use falling film reboilers.

[0111] Comparative Example 1

[0112] The operating parameters of Comparative Example 1 are the same as those of Example 1, but a traditional batch process is used, and only one oxidation reactor is used. There is no outer cylinder, inner cylinder and continuous device of this application, nor is there the post-processing unit and multi-stage separation unit of this application.

[0113] Comparative Example 2

[0114] The operating parameters of Comparative Example 2 are the same as those of Example 2, but multiple oxidation reactors connected in series are used. There are no outer and inner cylinders of this application, nor are there the post-processing unit and multi-stage separation unit of this application.

[0115] The following table compares the effects of an embodiment of the chlorophthalic anhydride production system of the present invention with those of a comparative example:

[0116]

[0117] As shown in the table above, the product yield, quality stability, and purity of Examples 1 and 2 of the present invention are higher than those of Comparative Examples 1 and 2. Furthermore, Examples 1 and 2 can separate 4-chlorophthalic anhydride and 3-chlorophthalic anhydride and obtain higher purity. Moreover, the amounts of chloro-o-xylene and acetic acid consumed in Examples 1 and 2 are significantly less than those in the comparative examples. Therefore, the chlorophthalic anhydride production system of the present invention can achieve higher yield and purity with less raw material and acetic acid consumption. Meanwhile, the batch method cannot maintain the same operation for each batch; variations due to different personnel, locations, and times result in poor repeatability, leading to inconsistent purity and poor stability in each batch. Since the chlorophthalic anhydride production system of the present invention is a fully continuous flow system with consistent raw materials, equipment, and operating conditions, the product quality purity varies less and the stability is stronger.

[0118] The above description is merely a specific embodiment of the present invention, but the scope of protection of the present invention is not limited thereto. Any variations or substitutions that can be easily conceived by those skilled in the art within the technical scope disclosed in the present invention should be included within the scope of protection of the present invention. Therefore, the scope of protection of the present invention should be determined by the scope of the claims.

Claims

1. A method for producing chloro-phthalic anhydride, characterized by, The method comprises: carrying out an oxidation reaction in an oxidation reactor with chloro-o-xylene and acetic acid as raw materials, air as an oxidant under the action of a catalyst, distributing air to the area between the outer cylinder and the inner cylinder of the oxidation reactor, and carrying out a first oxidation reaction in the outer cylinder; distributing air to the inner cylinder, carrying out a second oxidation reaction in the inner cylinder, obtaining a gas phase product and an oxidation reaction liquid, the oxidation reaction liquid containing at least chloro-phthalic acid, and the gas phase product containing at least acetic acid; carrying out condensation treatment on the gas phase product to obtain an acetic acid solution, then separating acetic acid in the acetic acid solution and sending the acetic acid into the oxidation reactor for continuous oxidation reaction; carrying out acetic acid separation on the oxidation reaction liquid to obtain a first purified reaction liquid and a first gas phase component containing acetic acid, carrying out acetic acid separation on the first purified reaction liquid to obtain a second purified reaction liquid and a second gas phase component containing acetic acid, sending the second purified reaction liquid into an anhydride forming reactor for anhydride forming reaction, and then separating acetic acid from the first gas phase component and the second gas phase component; carrying out dehydration on the second purified reaction liquid to obtain an anhydride crude product, and then separating at least two chloro-phthalic anhydride monomers from the anhydride crude product.

2. The method for producing chloro-phthalic anhydride according to claim 1, characterized by, The method further comprises: carrying out condensation treatment on the gas phase product to obtain a first acetic acid solution and tail gas, absorbing the tail gas with a tail gas absorption unit, and separating a second acetic acid solution remaining in the tail gas; carrying out condensation treatment on the gas phase product to obtain a first acetic acid solution and tail gas, absorbing the tail gas with a tail gas absorption unit, and separating a second acetic acid solution remaining in the tail gas; separating acetic acid in the first acetic acid solution and the second acetic acid solution.

3. The method for producing chloro-phthalic anhydride according to claim 1, characterized by, The method further comprises: dividing the acetic acid solution into a first acetic acid solution and a second acetic acid solution, refluxing the first acetic acid solution into the oxidation reactor for continuous reaction, and sending the second acetic acid solution into an acetic acid recovery unit for recovery.

4. The method for producing chloro-phthalic anhydride according to claim 1, characterized by, The method further comprises: carrying out dehydration on the second purified reaction liquid to obtain an anhydride crude product containing at least acetic acid and chloro-phthalic anhydride.

5. The method of producing chloro-phthalic anhydride according to claim 1, characterized by, The anhydride crude product is a gaseous anhydride crude product, and the method further comprises: separating acetic acid contained in the anhydride crude product, and condensing the anhydride crude product to obtain a condensed anhydride crude product; sending the separated acetic acid in the anhydride crude product into a post-processing unit for post-processing.

6. The method of producing chloro-phthalic anhydride according to claim 1, characterized by, The method further comprises: separating catalyst contained in the anhydride crude product, and then condensing the anhydride crude product after the catalyst is separated.

7. The method of producing chloro-phthalic anhydride according to claim 1, characterized by, The method further comprises: separating a light component from the anhydride crude product with a first separation unit; separating 4-chloro-phthalic anhydride from the anhydride crude product with a second separation unit; separating 3-chloro-phthalic anhydride from the anhydride crude product with a third separation unit.

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