Method for recycling by-product hcl in isocyanate production process

Abstract

The present application provides a kind of to the by-product HCl in isocyanate production process is recycled method, based on the method of the present application, can effectively improve compressor unit blockage, oxidation reactor catalyst poisoning deactivation, product chlorine gas is entrained solid and leads to equipment blockage and other problems, to improve the running stability of recycling process, thereby favoring.This method includes purifying by-product HCl, after by-product HCl is purified, by compressor pressure boosting, cooler cooling and low-temperature absorption tower absorption, send to oxidation reactor in the presence of catalyst by oxidation reaction to prepare chlorine gas, its characterized in that, the by-product HCl is purified and includes removing amine impurities in by-product HCl, and the amine impurities are the substance containing NH2-group and molecular weight less than 200.

Description

Technical Field

[0001] This invention relates to a technology for recycling HCl, a byproduct of isocyanate production processes. Background Technology

[0002] In the production of isocyanates, CO and chlorine react to produce phosgene. Phosgene then reacts with amines to produce isocyanates and the byproduct HCl. This process consumes a large amount of chlorine and releases a significant amount of HCl. Therefore, an HCl oxidation process is typically used to convert HCl into chlorine, thus achieving a chlorine cycle in the isocyanate industry. Currently, a relatively mature HCl oxidation process is the catalytic oxidation process developed by Sumitomo Chemical in Japan.

[0003] To address the issues of impurities in the byproduct HCl, which can lead to catalyst deactivation and blockage of subsequent units, the industry offers several other methods to achieve chlorine recycling:

[0004] 1. Hydrochloric acid analysis method

[0005] This method involves using purified water to absorb the byproduct HCl to produce concentrated hydrochloric acid with a concentration of 32-34%. The 32-34% hydrochloric acid is then heated using a distillation column to obtain gaseous HCl and dilute hydrochloric acid with a concentration of 18-22%. The gaseous HCl is dried and dehydrated before being sent for catalytic oxidation to produce chlorine gas. The dilute acid with a concentration of 18-22% is returned to the hydrochloric acid absorption unit to absorb the byproduct HCl together with purified water.

[0006] The advantage of this method is that the byproduct HCl, after being absorbed by pure water, can fix the impurities in it within the dilute hydrochloric acid, resulting in high-purity gaseous HCl that does not require further impurity removal. However, the hydrochloric acid desorption process consumes a large amount of steam, leading to high production costs and poor economic efficiency.

[0007] 2. Chlor-alkali electrolysis method

[0008] This method involves using purified water to absorb the byproduct HCl to obtain concentrated hydrochloric acid with a concentration of 32-34%. The concentrated hydrochloric acid is then neutralized with sodium hydroxide to obtain a sodium chloride solution. The organic matter in the solution is then disposed of by oxidative decomposition with sodium hypochlorite. The concentration of sodium chloride is adjusted to saturation by adding alkali and water, and then sent for electrolysis to obtain chlorine gas and sodium hydroxide solution.

[0009] This method converts hydrochloric acid into sodium chloride through alkali neutralization, and then electrolyzes it to obtain chlorine gas. However, due to the high energy consumption and alkali consumption of electrolysis, the production cost is high and the economic efficiency is poor. In addition, impurities such as organic matter and metal ions in the sodium chloride solution have a significant impact on the electrolytic membrane. Therefore, it is necessary to strictly control the impurities in the sodium chloride solution.

[0010] 3. Byproduct HCl oxychlorination process

[0011] This method removes impurities from the by-product HCl, mainly phosgene and solvents, and then sends it to an oxidation reactor to react and produce chlorine gas, thus achieving the recycling of chlorine.

[0012] This method frequently encounters problems such as solid blockage in the post-processing section and catalyst poisoning during operation. To address these issues, the industry typically uses adsorbents such as activated carbon to pretreat the by-product HCl and absorb trace amounts of chlorobenzene, dichlorobenzene, or toluene solvents. However, these methods cannot completely solve the blockage and poisoning problems.

[0013] According to publicly available patent reports, the following methods are currently used to address issues such as high HCl impurity content, low reaction yield, and equipment blockage:

[0014] Patent CN213505963 describes a device for preparing electronic-grade high-purity hydrogen chloride. The method removes impurities from dry HCl through an HCl absorption tower, an HCl desorption tower, a dehydration drying tower, a particulate impurity filter, and a distillation tower. However, due to the use of desorption process to purify HCl, the energy consumption is large and the economic efficiency is poor.

[0015] Patent CN102101651 describes a method for refining hydrogen chloride, a byproduct of difluorochloromethane, and for recovering trifluoromethane. The refining of HCl includes coarse separation, absorption, desorption, condensation, acid mist capture, and adsorption. However, the use of desorption process to refine HCl results in high energy consumption and poor economic efficiency.

[0016] Patent CN110198914A introduces a method for flexibly controlling the continuous chemical production of hydrochloric acid. This method involves sending hydrochloric acid to a neutralization station and neutralizing it with concentrated alkali solution, especially concentrated sodium hydroxide solution, to form a salt solution, which is then sent to a chlor-alkali electrolysis station. This achieves the recycling of HCl to hydrochloric acid, hydrochloric acid to sodium chloride solution, and then to chlorine gas. This method consumes a large amount of electrical energy and alkali solution, the process is relatively complex, and the route has poor reliability and economy.

[0017] Patent CN109455670A describes a purification process for HCl byproducts from an F32 unit, including cooling, separation, adsorption, drying, and re-run-off steps. It can produce products with purities of 99.9% and 99.9999%. This method purifies HCl through separation, adsorption, and drying. However, the adsorbents used are 3-5 mm 5A molecular sieves and 3-5 mm activated carbon, which have low efficiency and do not have a significant removal effect on the amines unique to isocyanate byproduct HCl.

[0018] Patent CN103922286A describes a method for recovering HCl during polycrystalline silicon production. The method involves sending the tail gas into an absorption tower, using silicon tetrachloride as an absorbent to absorb the HCl gas, and extracting the HCl as a liquid phase. The gas phase containing a small amount of light components and HCl extracted from the cryogenic reactor is returned to the absorption tower as feed. This method requires complete condensation of HCl through compression and condensation, followed by distillation to remove impurities. It is the most energy-intensive and least economical method among the above methods, and it cannot effectively remove amines due to their wide boiling point distribution.

[0019] In summary, existing methods for removing impurities from isocyanate byproduct HCl suffer from poor economic efficiency, incomplete identification of the impact of impurity composition and content on HCl pressurization and oxidation to chlorine processes, and a lack of effective measures for impurity control and identification. A new process is needed to achieve goals such as reduced energy consumption, increased yield, and reduced clogging. Summary of the Invention

[0020] In view of this, the present invention provides a method for recycling HCl, a byproduct of isocyanate production. Based on the method of the present invention, problems such as compressor blockage, catalyst poisoning and deactivation in oxidation reactors, and equipment blockage caused by solids entrained in the product chlorine gas can be effectively improved, thereby improving the operational stability of the recycling process and reducing energy consumption.

[0021] To achieve its objective, the present invention provides the following technical solution:

[0022] This invention provides a method for recycling HCl, a byproduct of isocyanate production. The method includes purifying the byproduct HCl, pressurizing the purified byproduct HCl with a compressor, cooling it with a cooler, absorbing it in a low-temperature absorption tower, and then sending it to an oxidation reactor where it undergoes an oxidation reaction in the presence of a catalyst to produce chlorine gas. The purification of the byproduct HCl includes removing amine impurities from the byproduct HCl, wherein the amine impurities are substances containing NH2- groups and having a molecular weight of less than 200.

[0023] Preferably, the by-product HCl is purified until the amine impurity content in the by-product HCl is <100 ppm. Reducing the amine impurity content to <100 ppm can significantly improve the operational stability of the recycling process.

[0024] Furthermore, the amine impurities include, but are not limited to, one or more of ammonia, ammonium chloride, aminomethane, and aminoethane.

[0025] Furthermore, an amine adsorbent is used to adsorb and remove amine impurities from the by-product HCl. The preparation steps of the amine adsorbent include:

[0026] 1) Using styrene as the skeleton raw material for the amine absorber, styrene, stilbene, liquid paraffin and an initiator are added to an aqueous sodium chloride solution for suspension polymerization to obtain copolymer beads; wherein, the mass ratio of styrene to stilbene is 3:1-9:1, preferably 3.5:1-6.5:1; the mass ratio of styrene to liquid paraffin is 4:1-20:1, preferably 7.5:1-8.0:1; the mass ratio of styrene to initiator is 70:1-150:1, preferably 75:1-150:1, more preferably 90:1-125:1;

[0027] 2) The copolymer beads obtained in step 1) are purified;

[0028] 3) The copolymer beads obtained in step 2) are subjected to sulfonation reaction with sulfonating agent, and the resulting product is washed with water for later use.

[0029] In some specific embodiments, in step 1), the sodium chloride aqueous solution can be a saturated sodium chloride aqueous solution, or the concentration of the sodium chloride aqueous solution is 10-22 wt%, and the mass ratio of the sodium chloride aqueous solution to styrene is 5:1-20:1, preferably 8.0:1-8.5:1.

[0030] In some specific embodiments, in step 1), the initiator is benzoyl peroxide.

[0031] In some specific embodiments, in step 1), the styrene, stilbene, liquid paraffin and initiator are added to an aqueous sodium chloride solution and stirred to completely disperse the oil droplets. Then, the reaction is carried out at 60-80℃, preferably 70-75℃, for 1-4 hours, preferably 2-2.5 hours. Then, the temperature is raised to 80-100℃, preferably 90-95℃, and the reaction is carried out for 1-4 hours, preferably 2-2.5 hours. After the reaction is completed, the product is cooled, filtered and washed with hot water, dried and sieved to obtain the copolymer beads, which are macroporous crosslinked styrene-stilbene resin skeletons.

[0032] In some specific embodiments, in step 2), the purification process includes soaking and swelling the copolymer beads in a solvent, wherein the mass ratio of the copolymer beads to the solvent is 1:20-1:100, preferably 1:40-1:60, and the soaking time is 2-48 hours, preferably 12-20 hours; the solvent is preferably one or more of dichloromethane, dimethyl sulfoxide, tetrahydrofuran, and chlorobenzene, more preferably dichloromethane.

[0033] In some specific embodiments, in step 3), the sulfonating agent is selected from one or more of sulfur trioxide, concentrated sulfuric acid, and fuming sulfuric acid, with sulfur trioxide being preferred.

[0034] In some specific embodiments, the mass ratio of the copolymer beads to the sulfonating agent is 1:2-1:10, preferably 1:3.5-1:6.

[0035] In some specific embodiments, the sulfonation reaction is carried out at a temperature of 20-60°C, preferably 30-50°C, and for a reaction time of 3-20 h, preferably 6-9 h.

[0036] The average particle size of the amine adsorbent is 400-800 μm, preferably 450-750 μm, and more preferably 550-600 μm; the volume of resin with a particle size of 200-600 μm in the amine adsorbent is no more than 5%, preferably no more than 5%; the volume of resin with a particle size of 600-1200 μm is no less than 75%, preferably no less than 75%; and the volume of resin with a particle size of 700-800 μm is no less than 75%. Adopting the preferred particle size requirements can effectively avoid the phenomena of flow deviation and short-circuiting of local amine adsorbent adsorption during the amine adsorption process.

[0037] In some specific embodiments, the volume exchange capacity of the amine absorber is not less than 0.8-1.15 eq / L, preferably not less than 1.0-1.15 eq / L, and the determination method adopts GB / T 8144-2008 Method for Determining the Exchange Capacity of Ion Exchange Resins.

[0038] In some specific embodiments, the uniformity coefficient of the amine absorbent is not greater than 1.1-1.4, preferably not greater than 1.2-1.25, and the determination method adopts GB5758-86 Ion Exchange Resin Particle Size Distribution Determination Method.

[0039] In some specific embodiments, the purification of the by-product HCl also includes removing the solvent from the by-product HCl. Removing the solvent from the by-product HCl helps improve the operational stability of the recycling process.

[0040] In some embodiments, the method for recycling the by-product HCl specifically includes the following steps:

[0041] The by-product HCl is fed into an acidic resin adsorption tower filled with the amine adsorbent to adsorb and remove amine impurities from the by-product HCl; preferably, the acidic resin tower operates at a pressure of 1.1-3.0 barg and an operating temperature of -20°C to 40°C.

[0042] The HCl treated by the acidic resin adsorption tower is sent to the solvent adsorption tower to remove the solvent; preferably, the operating pressure of the solvent adsorption tower is 1.1-3.0 barg and the operating temperature is -20℃ to 40℃.

[0043] After being processed by the solvent adsorption tower, the HCl is pressurized by a compressor and cooled by a cooler before being sent to a low-temperature absorption tower for absorption. Then, it is sent to an oxidation reactor where it undergoes an oxidation reaction with oxygen under the action of a catalyst to produce chlorine.

[0044] The processes of pressurizing HCl with a compressor, cooling it with a cooler, absorbing it in a low-temperature absorption tower, and then catalytically oxidizing it in an oxidation reactor are all conventional processes in the art and will not be described in detail. In some specific embodiments, the pressure of the gas at the compressor inlet is 0.1 barg-5 barg, preferably 1.5-2.5 barg, and the pressure of the gas at the compressor outlet is 3.0-12 barg, preferably 6-10 barg; the cooler cools the gas to about 40°C.

[0045] In some specific embodiments, the solvent adsorption tower is filled with activated carbon and / or molecular sieves for adsorbing the solvent, preferably activated carbon. The solvent absorption tower can be a solid bed filled with packing material that has solvent adsorption capabilities (such as activated carbon and / or molecular sieves). The solvents removed by the solvent absorption tower include chlorobenzene, dichlorobenzene, toluene, etc., which are solvents used in the phosgenation reaction and are carried by HCl gas. In some specific embodiments, the solvent adsorption tower processes 20,000-50,000 tons of HCl per ton of packing material, preferably 40,000-45,000 tons. The HCl treated by the acid resin adsorption tower flows through the solvent absorption tower, where the solvent in the HCl is adsorbed and removed.

[0046] In some specific embodiments, the absorbent used in the low-temperature absorption tower is liquid HCl; preferably, the liquid HCl used as the absorbent is derived from the cooled solvent adsorption tower outlet HCl (i.e., the HCl output from the solvent absorption tower), and the HCl output from the solvent absorption tower is cooled to a temperature of -30 to -5°C, preferably -25 to -20°C, before being used as the absorbent in the low-temperature absorption tower; the operating pressure of the low-temperature absorption tower is 3.0-12 barg, preferably 6-10 barg.

[0047] In some specific embodiments, the amine adsorbent in the acidic resin adsorption tower is filled in a loose pile, and each ton of amine adsorbent can process 5,000-10,000 tons of HCl, preferably 7,000-8,500 tons, and then regenerated.

[0048] The amine absorbent provided by this invention is an acidic resin with ion exchange function. Through the strong acidity and adsorption / interception of this resin, amines are fixed onto the resin. The resin can then be reused through regeneration and desorption. The regeneration method of the amine absorbent in this invention is as follows: The saturated amine absorbent is mixed with hydrochloric acid at a mass ratio (20:1 to 1:1), with a hydrochloric acid concentration of 5wt%-20wt%. After mixing, the mixture is heated to 40-75℃ for 0.5-4 hours.

[0049] In some specific embodiments, the oxidation reaction is carried out in a fixed-bed reactor or a fluidized-bed reactor; preferably in a fluidized-bed reactor, with a feed ratio of HCl to oxygen of 2:1.02-2:1.1 by molar weight, an oxidation temperature of 260-480°C, preferably 370-440°C, an oxidation pressure of 1.5-12 barg, preferably 2.0-8.0 barg, and an oxidation conversion rate of 80-95%, preferably 85-91%. The catalyst used in the oxidation reaction can be a copper-based catalyst, such as copper chloride or copper oxide, specifically commercially available products, such as copper-based catalysts from Sumitomo Chemical Co., Ltd.

[0050] The inventors have surprisingly discovered that small-molecule amine impurities, such as ammonia, aminomethane, aminoethane, and aniline, contained in the byproduct HCl produced during isocyanate formation, cause catalyst deactivation and blockages in equipment like compressors when introduced into subsequent oxidation reactors. This results in the catalytic oxidation unit's operating cycle and catalyst consumption failing to meet expectations. However, existing technologies do not recognize the impact of amines on oxidation reactors, compression refrigeration systems, and post-processing stages, and lack methods for amine removal and detection. Consequently, efficient, long-term, and stable utilization of byproduct HCl is impossible, hindering optimal process stability and economy. The inventors have investigated the effects of amines on catalyst activity and the stability of compression refrigeration systems, and have achieved stable system operation by employing specific methods to control amine impurities. This invention improves upon existing methods for recycling byproduct HCl in isocyanate production processes by removing amine impurities from the byproduct HCl, thereby enhancing the operational stability of the recycling process.

[0051] The technical solution provided by this invention has the following beneficial effects:

[0052] By using the method of the present invention, small molecule amines in HCl are removed, especially to a concentration of less than 100 ppm, which can help improve problems such as scaling and clogging of compressor units, poisoning and deactivation of catalysts in oxidation reactors, and blockage caused by solids entrained in the product chlorine gas in the recycling process.

[0053] By employing the recycling method of this invention, high-purity HCl (e.g., purity not less than 99.99%) can be obtained by purifying the by-product HCl. The small-molecule amines in this HCl are less than 100 ppm. By controlling the content of the above-mentioned amines, problems such as scaling and clogging of compressor units, catalyst poisoning and deactivation in oxidation reactors, and blockage caused by solids entrained in the product chlorine gas can be improved.

[0054] By applying this byproduct HCl recycling method, it is possible to reduce HCl processing energy consumption (e.g., by 80%), significantly reduce processing technology investment (e.g., by 60%), extend catalyst operating cycle (e.g., by 75%), significantly reduce the solid content in the product chlorine (e.g., by more than 99%), and extend the operating cycle of equipment (e.g., to more than two years). Attached Figure Description

[0055] Figure 1 This is a process flow diagram illustrating the recycling of HCl, a byproduct of isocyanate production, in one embodiment. Detailed Implementation

[0056] To better understand the technical solution of the present invention, the following embodiments further illustrate the content of the present invention, but the content of the present invention is not limited to the following embodiments. The test methods not specifically described in the following embodiments are conventional test methods in the art.

[0057] In the following examples, the amine impurities in the byproduct HCl are ammonia, ammonium chloride, aminomethane, and aminoethane. These amine impurities were determined using gas chromatography. The determination methods are described below:

[0058] HCl was absorbed by methanol, and the total amount of amine impurities was determined by gas chromatography with an ECD detector. The chromatographic conditions were as follows: column oven temperature was maintained at 35℃ for 8 min, then increased to 100℃ at a rate of 5℃ / min, and then increased to 200℃ at a rate of 10℃ / min and held for 5 min; column flow rate: 1.5 ml / min; injection port temperature: 220℃; detector temperature: 320℃; split ratio: 5:1; make-up gas flow rate: 60 ml / min.

[0059] Standard curve plotting: Use a microsyringe to transfer 5.0 μL and 10.0 μL of the standard working solution, respectively. The concentration of the standard working solution is 100 μg / ml. Plot the standard curve with the target concentration as the x-axis and the corresponding response value as the y-axis.

[0060] The total amount is calculated by matching the response value of each substance with the standard curve to obtain the concentration of each substance, and summing them up to obtain the total amount of amines.

[0061] Example 1 of preparation of amine absorbent

[0062] An example of preparing an amine absorbent includes the following reagents measured in parts by weight:

[0063]

[0064] 1) Styrene was selected as the skeleton material for the amine absorber. Styrene, stilbene, liquid paraffin, and the initiator benzoyl peroxide were added to an aqueous sodium chloride solution. The oil droplets were evenly dispersed under stirring at 1200 rpm, and then the reaction was carried out at 72°C for 2 hours. The reaction was then carried out at 90°C for 2 hours. After the reaction, the product was cooled to 30°C, filtered, washed with hot water at 40°C, dried, and sieved to obtain the copolymer beads (the copolymer's number-average molecular weight was 5500). In this step, the concentration of the aqueous sodium chloride solution used was 15 wt%.

[0065] 2) Add copolymer beads and a solvent for swelling to a three-necked flask. The solvent is dichloromethane. The mass ratio of copolymer beads to solvent is 1:40. Soak for 20 hours to allow the copolymer beads to fully swell.

[0066] 3) Then, sulfur trioxide was added for sulfonation reaction. The mass ratio of copolymer beads to sulfur trioxide was 1:6, and the reaction was carried out at 50°C for 6 hours. After the reaction was completed, the product was transferred to a beaker and washed with deionized water at 20°C for later use.

[0067] Example 2 of preparation of amine absorbent

[0068] An example of preparing an amine absorbent includes the following reagents measured in parts by weight:

[0069]

[0070]

[0071] 1) Styrene was selected as the skeleton material for the amine absorber. Styrene, stilbene, liquid paraffin, and the initiator benzoyl peroxide were added to an aqueous sodium chloride solution. The oil droplets were evenly dispersed under stirring at 1000 rpm, and then the reaction was carried out at 75°C for 2.5 h. The reaction was then carried out at 95°C for another 2.5 h. After the reaction, the product was cooled to 30°C, filtered, washed with hot water at 40°C, dried, and sieved to obtain the copolymer beads (the copolymer's number-average molecular weight was 6300). In this step, the concentration of the aqueous sodium chloride solution used was 15 wt%.

[0072] 2) Add copolymer beads and a solvent for swelling to a three-necked flask. The solvent is dichloromethane. The mass ratio of copolymer beads to solvent is 1:60. Soak for 12 hours to allow the copolymer beads to fully swell.

[0073] 3) Then, sulfur trioxide was added for sulfonation reaction. The mass ratio of copolymer beads to sulfur trioxide was 1:3.5, and the reaction was carried out at 30°C for 9 hours. After the reaction was completed, the product was transferred to a beaker and washed with deionized water at 20°C for later use.

[0074] Example 1

[0075] This embodiment uses the amine absorbent from Preparation Example 1.

[0076] The raw material used in this embodiment is by-product HCl, which comes from the MDI production process and has an amine content of 820 ppm.

[0077] The amine absorber obtained in Preparation Example 1 was used; the average particle size of the amine absorber was 520 μm, the volume of resin with a particle size distribution of 200-600 μm was not greater than 5%, the volume of resin with a particle size distribution of 600-1200 μm was not less than 75%, the volume exchange capacity of the amine absorber was 1.20 eq / L, and the uniformity coefficient was not greater than 1.23.

[0078] The aforementioned byproduct HCl is passed into an acidic resin adsorption tower filled with an amine absorbent. The amine absorbent is randomly packed, and the volume of the acidic resin adsorption tower is 20 m³. 3The acidic resin adsorption tower operates at a pressure of 2.3 barg and a temperature of 15°C. Refined HCl with an amine content of 42 ppm is obtained through the acidic resin adsorption tower. This refined HCl is then sent to a solvent adsorption tower to remove the solvent. The solvent adsorption tower is filled with activated carbon and operates at a pressure of 2.2 barg and a temperature of 15.5°C. The HCl is then sent to a compressor with an inlet pressure of 2.1 barg and an outlet pressure of 6.5 barg. The compressor outlet HCl is then cooled to 40°C by a cooler before being sent to a low-temperature absorption tower. The absorbent in the low-temperature absorption tower is the HCl from the solvent adsorption tower outlet, cooled to -22°C for use as the absorbent. The operating pressure of the low-temperature absorption tower is 6.5 barg. The obtained HCl purity was 99.99%. The treated HCl was then sent to an oxidation reactor for catalytic oxidation. The oxidation reactor was a fluidized bed reactor, and the oxychlorination catalyst used was a copper-based catalyst (CuO-CeO-2MO type catalyst) supplied by Sumitomo, Japan. The reactor temperature was 390℃, the pressure was 2.95 barg, and the ratio of HCl to oxygen (molarity) was 2:1.04. The solid content in the product chlorine was 0.05%, the oxidation conversion rate was 88%, the oxychlorination catalyst operating cycle was 6 years, the energy consumption of the oxychlorination reaction process was 60 KW / tHCl, and the compressor cleaning cycle was 30 months.

[0079] Example 2

[0080] This embodiment uses the amine absorbent from Preparation Example 1.

[0081] The raw material used in this embodiment is by-product HCl, which comes from the MDI production process and has an amine content of 1750 ppm.

[0082] The amine absorber obtained in Preparation Example 1 was used; the average particle size of the amine absorber was 520 μm, the volume of resin with a particle size distribution of 200 μm-600 μm was not greater than 5%, the volume of resin with a particle size distribution of 600-1200 μm was not less than 75%, the volume exchange capacity of the amine absorber was 1.20 eq / L, and the uniformity coefficient was not greater than 1.23.

[0083] The aforementioned byproduct HCl is passed into an acidic resin adsorption tower filled with an amine absorbent. The amine absorbent is randomly packed, and the volume of the acidic resin adsorption tower is 20 m³. 3The acidic resin adsorption tower operates at a pressure of 2.3 barg and a temperature of 15°C. Refined HCl with an amine content of 65 ppm is obtained through the acidic resin adsorption tower. This refined HCl is then sent to a solvent adsorption tower to remove the solvent. The solvent adsorption tower is filled with activated carbon and operates at a pressure of 2.2 barg and a temperature of 15.5°C. The HCl is then sent to a compressor with an inlet pressure of 2.5 barg and an outlet pressure of 4.5 barg. The compressor outlet HCl is then cooled to 45°C by a cooler before being sent to a low-temperature absorption tower. The absorbent in the low-temperature absorption tower is the HCl from the solvent adsorption tower outlet, which is cooled to -25°C for use as an adsorbent. The operating pressure of the low-temperature absorption tower is 4.5 barg. The HCl obtained had a purity of 99.99%. The treated HCl was then sent to an oxidation reactor for catalytic oxidation. A fluidized bed reactor was used, and the oxychlorination catalyst was a copper-based catalyst (CuO-CeO-2MO type catalyst) supplied by Sumitomo Chemical Co., Ltd. of Japan. The reactor temperature was 410℃, the pressure was 2.5 barg, and the HCl to oxygen ratio (molarity) was 2:1.03. The oxidation conversion rate was 87%, the solid content in the product chlorine was 0.057%, the oxychlorination catalyst operating cycle was 6 years, the energy consumption of the oxychlorination reaction process was 61 kW / tHCl, and the cleaning cycle of the HCl compressor was 25 months.

[0084] Example 3

[0085] This embodiment uses the amine absorbent from Preparation Example 2.

[0086] The raw material used in this embodiment is by-product HCl, which comes from the MDI production process and has an amine content of 850 ppm.

[0087] The amine absorber obtained in Preparation Example 1 was used; the average particle size of the amine absorber was 590 μm, the volume of resin with a particle size distribution of 200-600 μm was not greater than 5%, the volume of resin with a particle size distribution of 600-1200 μm was not less than 75%, the volume exchange capacity of the amine absorber was 1.0 eq / L, and the uniformity coefficient was not greater than 1.2.

[0088] The aforementioned byproduct HCl is passed into an acidic resin adsorption tower filled with an amine absorbent. The amine absorbent is randomly packed, and the volume of the acidic resin adsorption tower is 20 m³. 3The acidic resin adsorption tower operates at a pressure of 3.0 barg and a temperature of -20°C. Purified HCl with an amine content of 63 ppm is obtained through the acidic resin adsorption tower. This purified HCl is then sent to a solvent adsorption tower to remove the solvent. The solvent adsorption tower is filled with activated carbon and operates at a pressure of 2.9 barg and a temperature of -18.5°C. The HCl is then sent to a compressor with an inlet pressure of 2.5 barg and an outlet pressure of 4.5 barg. The HCl exiting the compressor is cooled to 42°C by a cooler and then sent to a low-temperature absorption tower. The absorbent in the low-temperature absorption tower is the HCl from the solvent adsorption tower outlet, which is cooled to -30°C for use as an adsorbent. The operating pressure of the low-temperature absorption tower is 10.0 barg. The HCl obtained had a purity of 99.99%. The treated HCl was then sent to an oxidation reactor for catalytic oxidation. A fluidized bed reactor was used, and the oxychlorination catalyst was a copper-based catalyst (CuO-CeO-2MO type catalyst) supplied by Sumitomo Chemical Co., Ltd. of Japan. The reactor temperature was 440℃, the pressure was 8 barg, and the ratio of HCl to oxygen (molarity) was 2:1.04. The oxidation conversion rate was 88%, the solid content in the product chlorine was 0.055%, the oxychlorination catalyst operating cycle was 6 years, the energy consumption of the oxychlorination reaction process was 60 kW / tHCl, and the compressor cleaning cycle was 29 months.

[0089] Comparative Example 1

[0090] The raw material used in this case is by-product HCl, which comes from the MDI production process. This by-product HCl is sent directly to the compressor without being treated by the acidic resin adsorption tower and solvent adsorption tower. The compressor inlet pressure is 2.5 barg, and the outlet pressure is 4.5 barg. The HCl exiting the compressor is then cooled to 40°C by a cooler before being sent to a cryogenic absorption tower. The absorbent in the cryogenic absorption tower is liquid HCl cooled to -25°C, and the operating pressure of the cryogenic absorption tower is 4.5 barg. Its amine content is 820 ppm, and the HCl purity is 99.97%. The byproduct HCl was sent to an oxidation reactor for reaction. A fluidized bed reactor was used, and the oxychlorination catalyst was a copper-based catalyst (CuO-CeO-2MO type catalyst) supplied by Sumitomo, Japan. The reactor temperature was 405℃, the pressure was 2.6 barg, and the ratio of HCl to oxygen (molarity) was 2:1.05. The solid content in the product chlorine was 1.2%, the oxychlorination conversion rate was 83%, the oxychlorination catalyst had an operating cycle of 3.5 years, the energy consumption of the oxychlorination reaction process was 270 KW / tHCl, and the compressor cleaning cycle was 4 months.

[0091] Comparative Example 2

[0092] The byproduct HCl used in this case comes from the MDI production process. The byproduct HCl is sent directly to the compressor without being treated by the acidic resin adsorption tower and solvent adsorption tower. The compressor inlet pressure is 2.5 barg, and the outlet pressure is 4.5 barg. The HCl from the compressor outlet is then cooled to 45°C by a cooler before being sent to a cryogenic absorption tower. The absorbent in the cryogenic absorption tower is liquid HCl cooled to -25°C, and the operating pressure of the cryogenic absorption tower is 4.5 barg. Its amine content is 1750 ppm, and the HCl purity is 99.97%. The byproduct HCl was sent to an oxidation reactor for reaction. A fluidized bed reactor was used, and the oxychlorination catalyst was a copper-based catalyst (CuO-CeO-2MO type catalyst) supplied by Sumitomo, Japan. The reactor temperature was 405℃, the pressure was 3.0 barg, and the ratio of HCl to oxygen (molarity) was 2:1.06. The solid content in the product chlorine was 3.3%, the oxychlorination conversion rate was 82%, the oxychlorination catalyst operating cycle was 2.0 years, the energy consumption of the oxychlorination reaction process was 315 KW / tHCl, and the cleaning cycle of the HCl compressor was 3 months.

[0093] Comparative Example 3

[0094] The raw material used in this case was HCl, a byproduct with an amine content slightly higher than 100 ppm.

[0095] The raw material used in this case is HCl, a byproduct from which phosgene has been removed. This byproduct HCl comes from the MDI production process. The byproduct HCl is sent directly to the compressor without being treated by the acidic resin adsorption tower and solvent adsorption tower. The compressor inlet pressure is 2.5 barg, and the outlet pressure is 4.5 barg. The compressor outlet HCl is then cooled to 43°C by a cooler before being sent to a cryogenic absorption tower. The absorbent in the cryogenic absorption tower is liquid HCl cooled to -25°C, and the operating pressure of the cryogenic absorption tower is 4.5 barg. Its amine content is 120 ppm, and the HCl purity is 99.97%. The byproduct HCl was sent to an oxidation reactor for reaction. A fluidized bed reactor was used, and the oxychlorination catalyst was a copper-based catalyst (CuO-CeO-2MO type catalyst) supplied by Sumitomo, Japan. The reactor temperature was 400℃, the pressure was 3.1 barg, and the ratio of HCl to oxygen (molarity) was 2:1.04. The solid content in the product chlorine was 0.9%, the oxychlorination conversion rate was 84%, the oxychlorination catalyst operating cycle was 3.2 years, the energy consumption of the oxychlorination reaction process was 265 KW / tHCl, and the cleaning cycle of the HCl compressor was 4.5 months.

[0096]

[0097]

[0098] Those skilled in the art will understand that modifications or adjustments can be made to the present invention based on the teachings of this specification. These modifications or adjustments should also be within the scope defined by the claims of the present invention.

Claims

1. A method for recycling by-product HCl in an isocyanate production process, the method comprising purifying the by-product HCl, and sending the purified by-product HCl to an oxidation reactor to produce chlorine gas by an oxidation reaction in the presence of a catalyst after the purified by-product HCl is pressurized by a compressor, cooled by a cooler, and absorbed by a low-temperature absorption tower, characterized in that, The purification of the by-product HCl includes removing amine impurities from the by-product HCl by using an amine adsorbent, and removing solvent from the by-product HCl, the amine impurities being substances containing NH2 groups and having a molecular weight less than 200; and the purification is to purify the by-product HCl to an amine impurity content of <100 ppm in the by-product HCl; The preparation step of the amine adsorbent includes: 1) adding styrene, diphenyl ethane, liquid paraffin and an initiator into an aqueous sodium chloride solution for suspension polymerization to obtain copolymer beads; wherein the mass ratio of styrene to diphenyl ethane is 3:1-9:1; the mass ratio of styrene to liquid paraffin is 4:1-20:1; and the mass ratio of styrene to initiator is 70:1-150:1; 2) purifying the copolymer beads obtained in step 1); 3) sulfonating the copolymer beads obtained in step 2) with a sulfonating agent, and washing the obtained product with water for standby; The method for recycling the by-product HCl specifically includes the following steps: feeding the by-product HCl into an acid resin adsorption tower filled with the amine adsorbent to remove amine impurities from the by-product HCl; the operating pressure of the acid resin tower is 1.1-3.0 barg, and the operating temperature is -20℃ to 40℃; feeding the HCl treated by the acid resin adsorption tower into a solvent adsorption tower to remove solvent, the operating pressure of the solvent adsorption tower is 1.1-3.0 barg, and the operating temperature is -20℃ to 40℃; feeding the HCl treated by the solvent adsorption tower into a low-temperature absorption tower after being pressurized by a compressor and cooled by a cooler, and then feeding it into an oxidation reactor to perform an oxidation reaction under the action of a catalyst to obtain chlorine.

2. The method of claim 1, wherein, The amine impurities are one or more of ammonia, ammonium chloride, amine methyl and amine ethane.

3. The method of claim 1, wherein, In step 1), the mass ratio of styrene to diphenyl ethane is 3.5:1-6.5:

1.

4. The method of claim 1, wherein, In step 1), the mass ratio of styrene to liquid paraffin is 7.5:1-8.0:

1.

5. The method of claim 1, wherein, In step 1), the mass ratio of styrene to initiator is 90:1-125:

1.

6. The method according to any one of claims 1 to 5, characterized in that, In step 1), the concentration of the aqueous sodium chloride solution is 10-22wt%, and the mass ratio of the aqueous sodium chloride solution to styrene is 5:1-20:1; In step 1), the initiator is dibenzoyl peroxide; In step 1), the styrene, diphenyl ethane, liquid paraffin and initiator are added into the aqueous sodium chloride solution for stirring to make the oil beads fully dispersed, and then the reaction is carried out at 60-80℃ for 1-4h; then the temperature is increased to 80-100℃, and the reaction is carried out for 1-4h; after the reaction is completed, the product is cooled, filtered and washed with hot water, and then is dried and sieved to obtain the copolymer beads.

7. The method of claim 6, wherein, In step 1), the mass ratio of the aqueous sodium chloride solution to styrene is 8.0:1-8.5:

1.

8. The method of claim 6, wherein, In step 1), the temperature of the incubation reaction is 70-75℃, and the reaction time is 2-2.5h.

9. The method of claim 6, wherein, In step 1), the temperature is increased to 90-95℃, and the reaction is carried out for 2-2.5h.

10. The method according to any one of claims 1 to 5, characterized in that, In step 2), the purification treatment comprises soaking and swelling the copolymer beads with a solvent, the mass ratio of the copolymer beads to the solvent is 1:20-1:100, and the soaking time is 2-48 hours; the solvent is one or more of dichloromethane, dimethyl sulfoxide, tetrahydrofuran, and chlorobenzene.

11. The method of claim 10, wherein, In step 2), the mass ratio of the copolymer beads to the solvent is 1:40-1:60, and the soaking time is 12-20 hours.

12. The method according to any one of claims 1 to 5, characterized in that, In step 3), the sulfonating agent is selected from one or more of sulfur trioxide, concentrated sulfuric acid, and fuming sulfuric acid; The mass ratio of the copolymer beads to the sulfonating agent is 1:2-1:10; The reaction temperature of the sulfonation reaction is 20-60℃, and the reaction time is 3-20 hours.

13. The method of claim 12, wherein, The mass ratio of the copolymer beads to the sulfonating agent is 1:3.5-1:

6.

14. The method of claim 12, wherein, The reaction temperature of the sulfonation reaction is 30-50℃, and the reaction time is 6-9 hours.

15. The method according to any one of claims 1-5, wherein, The average particle size of the amine absorbent is 400-800 μm; and / or, the volume of the amine absorbent with a particle size of 200-600 μm is not more than 5%; the volume of the amine absorbent with a particle size of 600-1200 μm is not less than 75%; and / or, the volume exchange capacity of the amine absorbent is not less than 0.8 eq / L, measured according to GB / T8144-2008 Ion Exchange Resin Exchange Capacity Determination Method; and / or, the uniformity coefficient of the amine absorbent is not more than 1.4, measured according to GB5758-86 Ion Exchange Resin Particle Size Distribution Determination Method.

16. The method of claim 15, wherein, The average particle size of the amine absorbent is 450-750 μm; and / or, the volume of the amine absorbent with a particle size of 400-500 μm is not more than 5%; The volume of the amine absorbent with a particle size of 700-800 μm is not less than 75%; and / or, the volume exchange capacity of the amine absorbent is not less than 1.0 eq / L; and / or, the uniformity coefficient of the amine absorbent is not more than 1.

25.

17. The method of claim 15, wherein, The average particle size of the amine absorbent is 550-600 μm.

18. The method of any one of claims 1-5, wherein, The solvent adsorption tower is filled with activated carbon and / or molecular sieve for adsorbing the solvent; and / or, the absorbent used in the low-temperature absorption tower is liquid HCl; the operating pressure of the low-temperature absorption tower is 3.0-12 barg.

19. The method of claim 18, wherein, The liquid HCl used as the absorbent is derived from the HCl output from the cooled solvent adsorption tower; the HCl output from the solvent adsorption tower is cooled to a temperature of -30 to -5℃, and used as the absorbent of the low-temperature absorption tower.

20. The method of claim 19, wherein, The HCl output from the solvent adsorption tower is cooled to a temperature of -25 to -20℃.

21. The method of claim 18, wherein, The operating pressure of the low-temperature absorption tower is 6-10 barg.

22. The method of any one of claims 1-5, wherein, The oxidation reaction is carried out in a fixed bed reactor or a fluidized bed reactor.

23. The method of claim 22, wherein, The oxidation reaction is carried out in a fluidized bed reactor; the feed ratio of HCl to oxygen is 2:1.02-2:1.1 by mole; the oxidation temperature is 260-480℃; the oxidation pressure is 1.5-12 barg; and the oxidation conversion rate is 80-95%.

24. The method of claim 23, wherein, The oxidation temperature is 370-440℃, the oxidation pressure is 2.0-8.0 barg, and the oxidation conversion rate is 85-91%. The oxidation temperature is 370-440℃, the oxidation pressure is 2.0-8.0 barg, and the oxidation conversion rate is 85-91%.

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