A polycarbonate by-product recovery process

Abstract

The invention relates to the technical field of polycarbonate, and discloses a polycarbonate byproduct recovery process, which is characterized by comprising the following steps: a light removal step, a phenol removal step, a phenol refining step and a DPC refining step. A certain amount of water is added to the original byproduct system, and all phenol and water are first separated from the system in the form of azeotropy, and then after cooling and condensation, the phenol and water are separated by liquid-liquid stratification, and the mixture of DPC and water is separated by crystallization cooling. Compared with the prior art, the process of the invention can not only obtain high-purity phenol, but also high-purity DPC.

Description

Technical Field

[0001] The invention relates to the technical field of polycarbonate, in particular to a polycarbonate byproduct recovery process. Background Art

[0002] Polycarbonate (PC) is an important engineering plastic widely used in the electronics, automotive, optical and construction industries. In the traditional melt process of producing polycarbonate, diphenyl carbonate (DPC) and bisphenol A (BPA) are mainly used for polycondensation. However, in this production process, in addition to the target product polycarbonate, a series of by-products are also produced, mainly including phenol, and DPC and high-boiling substances (such as oligomers) are also entrained. As an important chemical raw material, phenol has a high market value. Usually, phenol needs to be recovered from the production process and utilized to reduce production costs.

[0003] There are two main methods for recovering polycarbonate byproducts:

[0004] The first method is distillation. However, due to the high boiling points of DPC and high boiling products (the boiling point of DPC is 309°C), even under vacuum distillation, the temperature at the bottom of the tower is still over 200°C (because the system contains high boiling products and oligomers, the distillation tower often uses a plate tower. In addition, due to the height of the liquid column at the bottom of the tower, even if the top of the tower is completely vacuum, the pressure at the bottom of the tower is still about 10kPa and the temperature is about 222°C due to the pressure drop of the tower plate and the influence of the liquid column at the bottom of the tower). DPC and high boiling products will decompose to varying degrees at high temperatures for a long time, forming some substances with boiling points close to phenol (such as phenyl ether, benzoquinone, cresol, acetophenone, etc.), making it difficult to obtain high-purity phenol products at the top of the tower (phenol purity <99.6%). In addition, the phenol recovery tower is a vacuum environment, and air will leak in to oxidize phenol into benzoquinone, causing phenol to change color. If DPC products are to be obtained, the bottom temperature of the DPC distillation tower needs to be higher (normal boiling point is greater than 350°C). At such a high temperature, some DPC and high-boiling substances will decompose, making it difficult to recover DPC. Therefore, the existing process will directly send DPC and high-boiling substances to waste liquid treatment (such as incineration).

[0005] For example, patent CN1676512A discloses a method for separating and purifying diaryl carbonate by vacuum distillation. The process is a method for recovering by-products of a process for synthesizing diaryl (phenyl) carbonate by melt ester exchange. The patent mentions that when the reaction liquid is exposed to a high temperature environment with a catalyst for a long time, the generated diaryl carbonate is prone to reverse reaction or decomposition, which will greatly reduce the yield.

[0006] For example, patent CN113577814B discloses a diphenyl carbonate recovery device and method for industrial production of polycarbonate. The process is to feed the refined phenol material into the DPC refining system, and then feed the bottom material of the DPC refining tower into the membrane vaporizer to further separate DPC and high-boiling products. Unfortunately, the patent does not disclose the relevant operating parameters. Part of the DPC and high-boiling products will decompose at high temperatures, and the composition is complex. The side line cannot obtain such high-purity DPC (99.95%). In addition, the total inlet and outlet materials in the embodiments disclosed in the patent are not conserved in mass, whether for a single substance or the total amount, so the authenticity of the patent is questionable.

[0007] The second is the crystallization method, such as patent CN100500640C discloses a diphenyl carbonate product separation preparation method, which is also a product separation method for the production of diphenyl carbonate. The reaction mixture is dissolved in methanol, ethanol or n-butanol as a solvent, and then crude diphenyl carbonate is obtained by cooling, crystallization, and vacuum drying, and then the steps are repeated to obtain refined diphenyl carbonate (more than 99.5%).

[0008] For example, patent CN117776921A discloses a separation method and separation system for polycarbonate polycondensation byproducts, which precipitates all DPC by gradually cooling and crystallizing, and then distills the remaining phenol to obtain high-purity phenol. Although this method can obtain phenol with higher purity, according to its embodiment, the purity of DPC is only 79-86%, that is, a lot of phenol is discharged as waste liquid along with DPC, and the patent does not explain the treatment plan of high-boiling substances.

[0009] For example, the literature "Study on the Purification of Diphenyl Carbonate by Water Crystallization" (Bai Rongxian et al., Chemical Engineering, Vol. 37, No. 3) used a mixed solvent of water and ethanol to purify crude DPC and obtained DPC with a purity greater than 99.5wt%.

[0010] It can be seen that the existing technology for the recovery of by-products is mostly limited to phenol, and the purity of the recovered phenol is not very high (usually around 98.5~99wt%), so it can only be sold as a qualified product, and DPC can usually only be treated as waste liquid (incineration). Summary of the invention

[0011] The technical problem to be solved by the present invention is to provide a polycarbonate byproduct recovery process in view of the current status of the prior art, which can not only obtain high-purity phenol, but also high-purity DPC.

[0012] The technical solution adopted by the present invention to solve the above technical problems is: a polycarbonate byproduct recovery process, characterized by comprising the following steps:

[0013] Process 1: Light removal process:

[0014] The by-product mixture from the polycarbonate production unit is mixed with circulating water and then enters the light component removal tower T1 to remove light components;

[0015] Step 2: Phenol removal step:

[0016] The bottom stream of the light-removal tower T1 enters the phenol separation tower T2. The vapor phase at the top of the phenol separation tower T2 is pressurized by the compressor C1 and enters the economizer H1 to heat the stream from the bottom of the phenol separation tower T2. After heat exchange and condensation, a part of the stream is returned to the top of the phenol separation tower T2 as reflux liquid, and the other part enters the phenol refining process. A part of the bottom stream of the phenol separation tower T2 enters the economizer H1 and then enters the reboiler H2. After being heated, it returns to the bottom of the phenol separation tower T2, and the other part enters the DPC refining process.

[0017] Step 3: Phenol refining step:

[0018] Part of the distillate from the top of the phenol separation tower T2 enters the cooler H4 for cooling, and then enters the separation tank F1 for liquid-liquid separation operation. The upper part of the separation tank F1 is a phenol-containing water phase, which is returned to the process 1 as circulating water. The lower part of the separation tank F1 is a water-containing phenol phase, which enters the phenol refining tower T3. The top of the phenol refining tower T3 is a phenol-containing water phase, which is returned to the process 1 as circulating water. The bottom of the phenol refining tower T3 produces refined phenol.

[0019] Process 4: DPC refining process:

[0020] Part of the material from process 2 enters the crystallization cooler H3 through the filter F2. DPC is precipitated as a solid and enters the liquid-solid separation equipment X1 for liquid-solid separation operation. The liquid phase output by the liquid-solid separation equipment X1 is the mother liquor wastewater. The solid phase output by the liquid-solid separation equipment X1 is the crude DPC product, which enters the liquid-solid washing equipment X2. At the same time, the supplemented pure water is added to the liquid-solid washing equipment X2, and the liquid-solid separation operation is performed again. The solid phase output by the liquid-solid washing equipment X2 is the refined DPC.

[0021] The mass content of phenol in the top stream of the phenol separation tower T2 in step 2 is recorded as x1, and the mass content of phenol in the upper layer liquid in the separation tank F1 in step 3 is recorded as x2. The relationship between x1 and x2 satisfies: x1>x2.

[0022] In order to effectively separate phenol water and take comprehensive energy consumption into consideration, the relationship between x1 and x2 satisfies: 2.5×x2≥x1≥1.5×x2.

[0023] In order to simplify the process, the lightness removal tower T1 in step 1 is operated under positive pressure.

[0024] In order to make DPC as the dispersed phase, the mass ratio of water to DPC in the bottom stream of the lightness removal tower T1 in step 1 is not less than 2.

[0025] In order to facilitate the subsequent liquid-liquid separation of the top stream and avoid the decomposition of DPC in the bottom stream, the top operating pressure of the phenol separation tower T2 in step 2 is 0.1-0.6 MPaG. If the top pressure of the phenol separation tower T2 is too low, the separation tank F1 in the subsequent step 3 cannot achieve liquid-liquid separation, and if the top pressure of the phenol separation tower T2 is too high, it will not only increase energy consumption, but more importantly, the temperature of the phenol separation tower T2 will become very high, so that the DPC and high boiling products at the bottom of the phenol separation tower T2 will decompose at high temperature, and the impurities will increase.

[0026] In order to facilitate the liquid-liquid separation of the phenol-water mixture and avoid its solidification, the cooling temperature of the cooler H4 in step 3 is -10~60°C.

[0027] In order to ensure that the composition of each phase in the separation tank meets the requirement of x1>x2, the operating temperature of the separation tank F1 in step 3 is -10~60°C.

[0028] In order to simplify the process, the phenol refining tower T3 in step 3 is operated at positive pressure or slightly positive pressure.

[0029] In order to ensure the smooth precipitation of DPC and save energy, the cooling temperature of the crystallization cooler H3 in step 4 is 70-78°C.

[0030] In order to effectively utilize wastewater, part of the liquid phase produced by the liquid-solid separation device X1 in step 4 is discharged as wastewater, and the other part is returned to step 1 as circulating water;

[0031] The liquid phase produced by the liquid-solid washing equipment X2 in step 4 is returned to step 1 as circulating water.

[0032] Compared with the prior art, the advantages of the present invention are:

[0033] (1) A certain amount of water is added to the original by-product system, and all phenol and water are first separated from the system in the form of azeotropy. Then, after cooling and condensing, phenol and water are separated by liquid-liquid stratification. At the same time, the mixture of DPC and water is separated by crystallization cooling. This process can not only obtain high-purity phenol, but also high-purity DPC;

[0034] (2) The bottom temperature of the phenol separation tower of the present invention is only slightly higher than the top temperature of the tower, and the heat pump technology is used to pressurize the steam at the top of the phenol separation tower and use it to heat the bottom flow of the phenol separation tower, which hardly increases the additional energy consumption. BRIEF DESCRIPTION OF THE DRAWINGS

[0035] Figure 1The process flow chart of an embodiment of the polycarbonate by-product recovery process of the present invention. DETAILED DESCRIPTION

[0036] The present invention is further described in detail below with reference to the accompanying drawings.

[0037] The system of phenol and water is quite special. When the temperature is above about 65°C, phenol and water can be miscible in any ratio. When the temperature is below 65°C, they can separate into phases quickly, and the composition changes greatly from miscible to separated phases (for example, the water content of the water layer at 64°C is 82.8%). The following are the liquid-liquid equilibrium data of phenol and water at various temperatures:

[0038] Table 1 Water content in phenol-water phase at different temperatures

[0039] .

[0040] In addition, phenol and water are an azeotropic system, that is, phenol and water will form a minimum azeotropic mixture in a certain ratio and evaporate from the liquid phase into the gas phase. The following is the azeotropic composition table of phenol and water:

[0041] Table 2 Water content of phenol-water azeotrope at different pressures

[0042] .

[0043] DPC and water are insoluble. Water and DPC can be separated simply by crystallization. By controlling the crystallization temperature, high-purity DPC can be obtained. However, if phenol is present, a small amount of phenol will co-crystallize with DPC. Therefore, if high-purity DPC is to be obtained, phenol cannot be contained in the system.

[0044] Through an in-depth analysis of the above system, the present application proposes that a certain amount of water is added to the original byproduct system, and all phenol and water are first separated from the system in the form of azeotropy, and then after cooling and condensation, the phenol and water are separated by liquid-liquid stratification. It should be noted here that the content of phenol in the azeotropic composition should be greater than the content of phenol in the water phase (upper liquid) in the liquid-liquid stratification, the water phase can be returned to the system for recycling, and the phenol phase is further distilled to obtain high-purity phenol.

[0045] In order to solve the problem of high bottom temperature in the original process, the present invention adopts the method of adding water in an amount higher than the azeotropic ratio of phenol and water to the feed, that is, after separation in the phenol separation tower, the bottom of the tower is a mixture of DPC and water, so that the bottom temperature of the tower is only slightly higher than the top temperature of the tower (depending on the amount of added water, the temperature is about 1-5°C higher).

[0046] It can be seen from Table 1 that the phenol-water azeotrope has a low phenol content, so a large amount of water will be azeotropically vaporized, resulting in high energy consumption. However, the bottom temperature of the phenol separation tower of the present invention is only slightly higher than the top temperature, and the heat pump technology is used to pressurize the steam at the top of the phenol separation tower to heat the bottom flow of the phenol separation tower, which hardly increases additional energy consumption. In addition, the pressure ratio of the compressor is very small (usually only a pressure ratio of about 1.2 is required to meet the heat transfer requirements).

[0047] like Figure 1 As shown, the device used in the polycarbonate byproduct recovery process of the present invention includes a lightness removal tower T1, a delivery pump P1, a phenol separation tower T2, a compressor C1, an economizer H1, a delivery pump P2, a reboiler H2, a cooler H4, a separation tank F1, a phenol refining tower T3, a filter F2, a crystallization cooler H3, a liquid-solid separation device X1, a liquid separator B1 and a liquid-solid washing device X2. It should be noted that Figure 1 It is a schematic diagram, and it is not particularly indicated whether the conveying power of the gas phase and liquid phase flow is sufficient. If necessary, a pump or fan can be added to increase the conveying power. The heat preservation and cold preservation of equipment and pipelines are handled according to general regulations, and the present invention does not emphasize it. In addition, the lightness removal tower T1 and the phenol separation tower T2 in the accompanying drawings belong to conventional distillation, so the auxiliary equipment of the distillation tower is simply indicated or not indicated in the figure, such as reflux tank, overhead condenser, reflux pump, tower bottom reboiler and other equipment.

[0048] The polycarbonate byproduct recovery process of the present invention mainly consists of four steps: light component removal step, phenol removal step, phenol refining step and DPC refining step. The light component referred to in the present invention refers to a component with a normal pressure boiling point less than 99.5°C, and the high boiling point referred to refers to a substance with a normal pressure boiling point greater than the normal pressure boiling point of DPC.

[0049] Process 1 is a light component removal process, in which the by-product mixture from the polycarbonate production unit and the circulating water (from process 3 and process 4) are mixed and then enter the light component removal tower T1 to remove the light components.

[0050] The above by-product mixture contains phenol, DPC, a small amount of light components, high boiling points (oligomers), etc. The typical composition is shown in the following table:

[0051] Table 3 Typical composition of by-product mixture

[0052] ;

[0053] Among them, the impurities in Table 3 refer to substances that are relatively difficult to separate from phenol (for example, substances with a boiling point difference of about ±30°C), such as anisole, cresol, acetophenone, benzoquinone, etc.

[0054] The above-mentioned circulating water comes from process 3 and process 4, and contains a small amount of phenol.

[0055] The by-product mixture and the circulating water are mixed and then enter the light-removal tower T1. The top distillate of the light-removal tower T1 is mainly light components, such as carbon dioxide, water, catalyst decomposition products and a small amount of phenol. The light components can be removed from the top of the light-removal tower T1 in the vapor phase or in the vapor-liquid two-phase manner.

[0056] Unlike conventional processes, in which the light components are removed by vacuum distillation due to the bottom temperature of the tower, the present invention can be operated under positive pressure because water is added to the system. For example, the tower top pressure can be operated at 5-30 kPaG.

[0057] In addition, the present invention adds water to the by-products, which has another function that water can hydrolyze oligomers to generate bisphenol A and carbon dioxide, which is beneficial to improve the purity of DPC precipitated in step 4. The bottom of the light removal tower T1 is a mixture of DPC and water (including a small amount of high boiling products). In order to make DPC as a dispersed phase, the mass ratio of water to DPC should be not less than 2, and the bottom temperature is about 104-167°C.

[0058] Process 2 is a phenol removal process. The bottom stream of the light-removal tower T1 enters the phenol separation tower T2 through the delivery pump P1. The vapor phase at the top of the phenol separation tower T2 is pressurized by the compressor C1 and enters the economizer H1 to heat the stream from the bottom of the phenol separation tower T2. After heat exchange and condensation, a part of it is returned to the top of the phenol separation tower T2 as reflux liquid, and the other part enters the phenol refining process. After the bottom stream of the phenol separation tower T2 passes through the delivery pump P2, a part enters the economizer H1 and then enters the reboiler H2. After being heated, it returns to the bottom of the phenol separation tower T2, and the other part enters the DPC refining process.

[0059] The top of the phenol separation tower T2 is an azeotropic vapor-phase phenol-water mixture, and the operating pressure of the top of the phenol separation tower T2 is 0.1~0.6 MPaG. If the top pressure of the phenol separation tower T2 is too low, the separation tank F1 in the subsequent process 3 cannot achieve liquid-liquid stratification. If the top pressure of the phenol separation tower T2 is too high, it will not only increase the energy consumption, but more importantly, the temperature of the phenol separation tower T2 will become very high, causing the DPC and high boiling products at the bottom of the phenol separation tower T2 to decompose at high temperature, resulting in more impurities.

[0060] The vapor phase phenol-water mixture at the top of the phenol separation tower T2 then enters the compressor C1 for pressurization. There is no special limit on the pressurization pressure, but for energy saving considerations, the logarithmic heat transfer temperature difference of the subsequent economizer H1 is generally ≥5°C. For example, when the top pressure of the phenol separation tower T2 is 0.4 MPaG and the outlet pressure of the compressor C1 is 0.5 MPaG (pressure ratio is 1.2), the heat transfer temperature difference of the economizer H1 can reach 6.5°C.

[0061] The pressurized stream of compressor C1 is heated by economizer H1 from the bottom stream of phenol separation tower T2. After heat exchange, the stream is changed from vapor phase phenol-water mixture to liquid phase phenol-water mixture, and part of the stream at the bottom of phenol separation tower T2 is vaporized. A part of the liquefied top stream, i.e., the liquid phase phenol-water mixture, is returned to the top of phenol separation tower T2 as reflux liquid, and the other part enters the phenol refining process. The reflux liquid flow rate (reflux ratio) is determined by the phenol content in the bottom of phenol separation tower T2. For example, the appropriate reflux rate is adjusted so that the phenol content in the bottom stream of phenol separation tower T2 is less than 0.01wt%.

[0062] After part of the bottom flow of the phenol separation tower T2 enters the economizer H1, it cannot fully meet the reboiler requirement, so another reboiler H2 is needed to meet the heat requirement and return to the bottom of the phenol separation tower T2. The other part of the bottom flow enters the DPC refining process.

[0063] Process 3 is a phenol refining process. Part of the distillate from the top of the phenol separation tower T2 enters the cooler H4 for cooling, and then enters the separation tank F1 for liquid-liquid separation operation. The upper part of the separation tank F1 is a phenol-containing water phase, which is returned to the process 1 as circulating water. The lower part of the separation tank F1 is a water-containing phenol phase, which enters the phenol refining tower T3. The top of the phenol refining tower T3 is a phenol-containing water phase, which is returned to the process 1 as circulating water. Refined phenol is produced at the bottom of the phenol refining tower T3.

[0064] The cooling temperature of the above-mentioned cooler H4 is -10~60℃, preferably 10~40℃, and more preferably 20~30℃. Although the freezing point of phenol is about 40℃ and the freezing point of water is 0℃, studies have found that the phenol-water system will not solidify at -10℃, and will still solidify at too low a temperature. Therefore, the cooling temperature cannot be too low, and the cooling temperature cannot be too high. As mentioned above, phenol-water will be miscible in any ratio above 65℃, and cannot be separated into liquid-liquid layers in the separation tank F1.

[0065] In addition, the temperature of cooler H4 should also be related to the top pressure of phenol separation tower T2 in process 2. Specifically, the mass content of phenol in the top stream of phenol separation tower T2 in process 2 is recorded as x1, and the phenol content in the upper layer liquid in separation tank F1 in process 3 is recorded as x2. The relationship between x1 and x2 must satisfy: x1>x2. This is very important. If it is not satisfied, the separation of phenol and water cannot be achieved. The reason is that the substance at the top of phenol separation tower T2 in process 2 is an azeotrope of phenol and water. The composition of the azeotrope is related to the tower pressure. Adjusting the tower pressure changes the composition of the azeotrope. In addition, the operating temperature of separation tank F1 in process 3 also corresponds to the two-phase composition in liquid-liquid stratification. Changing the temperature of the separation tank can adjust the composition of each phase. For example, when the top pressure of the phenol separation tower T2 is 0 MPaG, the phenol content x1 in the top azeotrope is 0.083, and if the cooling temperature of the cooler H4 is set to 50°C, then according to Table 1, the phenol content x2 of the water phase at liquid-liquid equilibrium is 0.108, x1<x2, and separation cannot be achieved, that is, in the separation tank F1, phenol water still exists in a single phase. In order to effectively achieve the separation of phenol water and in combination with comprehensive energy consumption considerations, preferably, 2.5×x2≥x1≥1.5×x2.

[0066] The cooled phenol-water mixture enters the separation tank F1 for liquid-liquid separation. Because the density of water is smaller than that of phenol, the upper liquid phase or upper layer of liquid phase in the separation tank F1 mentioned in this article refers to the water phase. The water phase contains a large amount of water, which is returned to process 1 for recycling; the concentration of phenol in the phenol phase is, for example, 0.725 at 20°C. The phenol phase enters the phenol refining tower T3. Since the system of the tower only contains the separation of phenol and water and is relatively clean, it can be operated at positive pressure or slightly positive pressure (nitrogen sealing pressure is sufficient), and a packed tower is used for separation. The azeotrope of water and phenol is obtained at the top of the phenol refining tower T3, and the azeotrope is returned to process 1 as circulating water; the bottom of the phenol refining tower T3 is pure phenol with a purity of more than 99.9wt%. When the top pressure of the phenol refining tower T3 is 10 kPaG, the top temperature of the phenol refining tower T3 is 102°C and the bottom temperature is about 185°C.

[0067] Process 4 is a DPC refining process. Part of the material from process 2 passes through one or more filters F2 and enters the crystallization cooler H3. DPC is precipitated as a solid and then enters the liquid-solid separation equipment X1 for liquid-solid separation operation. The liquid phase produced by the liquid-solid separation equipment X1 is mother liquor wastewater. The mother liquor wastewater can be directly discharged as wastewater, or part of it can be discharged as wastewater, and the other part can be returned to process 1 as circulating water. The solid phase produced by the liquid-solid separation equipment X1 is a crude DPC product, which enters the liquid-solid washing equipment X2. At the same time, the supplementary pure water is added to the liquid-solid washing equipment X2, and the liquid-solid separation operation is performed again. The liquid phase produced by the liquid-solid washing equipment X2 can be directly discharged as wastewater, or it can be returned to process 1 as circulating water. The solid phase produced by the liquid-solid washing equipment X2 is refined DPC, which goes to the next step of drying or drying treatment.

[0068] The temperature of some materials from process 2 is about 100~165℃, and the precipitated bisphenol A or other substances (if any) are removed through one or more filters F2. Then the logistics enters the crystallization cooler H3. Due to the precipitation of solids, the crystallization cooler H3 needs to consider anti-blocking measures, such as using a shell-and-tube heat exchanger or a backup heat exchanger. The cooling temperature of the crystallization cooler H3 is 70~78℃, and the crystallization temperature of DPC is 78.9℃, so the cooling temperature of the crystallization cooler should be lower than 78.9℃. In addition, since DPC and water are incompatible, the cooling temperature does not need to be too low to crystallize most of the DPC. In order to ensure the purity of the precipitated DPC, the cooling temperature of the crystallization cooler H3 is preferably 75~77℃, and in order to ensure that the DPC in the liquid-solid separation equipment is precipitated in powder or small-size particles, a mixing device can be added to the inlet connecting pipe of the crystallization cooler H3, such as a static mixer instead of the inlet pipe.

[0069] After DPC is precipitated as a solid, it enters the liquid-solid separation device X1, which includes at least one centrifugal filter or filter press and a solid conveying device (such as a screw conveyor, etc.). The DPC solid enters the liquid-solid washing device X1, and the mother liquor is discharged for waste treatment (the high boiling points and other impurities in the by-product feed of the present invention are mainly discharged from here). An impurity cooler and a liquid separator B1 can also be set up, cooled to, for example, 40°C and then allowed to stand, and the clear liquid is returned to step 1 as circulating water, and the bottom oil phase and solid phase (if any) are regularly discharged.

[0070] The liquid-solid washing equipment X2 at least comprises a washing kettle with stirring, a centrifugal filter or filter press and a solid conveying device (such as a screw conveyor, etc.). The supplemented pure water is added to the washing kettle, and the washing operation is carried out under stirring. The bottom material is discharged to the centrifugal filter or filter press for separation again to obtain refined DPC. The DPC still has a certain amount of moisture, and is then dried or dried in the next step. The purity of the dried DPC can reach more than 99.8wt%. The mother liquor of the liquid-solid washing equipment X2 can be returned to process 1 as circulating water. The liquid-solid separation operation and the liquid-solid washing operation can be operated continuously or intermittently. If intermittent operation is adopted, a set of equipment can be shared.

[0071] Embodiment 1:

[0072] A certain amount of by-products from a polycarbonate production device was collected and tested, and the components were shown in the following table:

[0073] Table 4 Composition of by-product mixture

[0074] ;

[0075] The by-products of the above composition are mixed with circulating water at a flow rate of 14 kg / h and 81.35 kg / h, and then continuously pumped into the light-removing tower T1. The top cooler of the light-removing tower T1 is a partial condenser, which removes light components in vapor phase and liquid phase respectively. The total amount of light components is 0.49 kg / h (including water, phenol and other substances). The bottom material of the light-removing tower T1 is pumped to the phenol separation tower T2 at a flow rate of 94.85 kg / h.

[0076] The vapor phase at the top of the phenol separation tower T2 enters the compressor C1 and is pressurized to 0.7 MPaG (pressure ratio is 1.14), and then enters the shell of the economizer H1 for heat exchange. The logarithmic heat transfer temperature difference of the economizer is 4.6℃, and it is condensed into a liquid phase with a temperature of 170℃. One part of it is refluxed to the top of the phenol separation tower T2, and the other part enters the cooler H4 to be cooled to 20℃, and then enters the separation tank F1. The operating pressure of the separation tank F1 is 0.67 MPaG (there is a partial pressure drop in the economizer and pipeline). The upper layer liquid of the separation tank F1 is returned as circulating water, and the lower layer liquid of the separation tank F1 directly enters the phenol refining tower T3 with a flow rate of 18 kg / h. The top of the phenol refining tower T3 is a phenol-water azeotropic mixture, which is pumped back as circulating water. 12.3 kg / h of refined phenol is obtained at the bottom of the phenol refining tower T3.

[0077] The liquid phase at the bottom of the phenol separation tower T2 is pumped, part of which is transported to the tube layer of the economizer H1, heated and vaporized, and then returned to the bottom of the phenol separation tower T2 through the reboiler H2, and the other part is used as the discharge material, after passing through the filter F2, it enters the sleeve type crystallization cooler H3 and is cooled to 75°C. The DPC refining process of this embodiment is an intermittent operation, and the cooled DPC material is continuously sent to the stirring tank for storage (flow rate is 6 kg / h). After two days of continuous and stable operation (a total of 288 kg), the above process is stopped, and the material in the stirring tank is fed to a closed centrifugal filter for filtration, the mother liquor is discharged for waste treatment, and the filter cake is temporarily stored. After the filtration is completed, all the filter cakes are placed back in the stirring tank, 200 kg of pure water is added for washing again, and then the filtration operation is repeated again. The obtained filter cake is dried and crushed to obtain 67.4 kg of refined DPC;

[0078] Among them, the setting of detection points such as temperature, pressure, flow, etc. belongs to the conventional means in this field and will not be described in detail one by one (for example, temperature and pressure measurement points should be set at the top and bottom of the distillation tower, etc.). The specifications and operating conditions of each distillation tower are shown in the following table:

[0079] Table 5 Specifications and operating conditions of each distillation column

[0080] .

[0081] Embodiment 2~5, comparative example 1~3:

[0082] The operating steps are basically the same as those in Example 1, except that the main differences are the operating parameters such as the pressure of the phenol separation tower T2, the operating temperature of the separation tank F1, the by-product feed and the circulating water volume. The amount of washing water in the liquid-solid washing process is scaled proportionally according to the feed (based on 200 kg of pure water in Example 1). The operating parameters of Comparative Examples 1 to 3 are basically the same as those in Example 5.

[0083] The operating parameters and test results of the above embodiments 1 to 5 are shown in the following table:

[0084] Table 6 Operating parameters and test results of Examples 1 to 5

[0085] .

[0086] The operating parameters and test results of the above comparative examples 1 to 3 are shown in the following table:

[0087] Table 7 Operating parameters and test results of Comparative Examples 1 to 3

[0088] .

[0089] Embodiment 6~7, comparative example 4~5:

[0090] The operation steps are basically the same as those in Example 1, and the operating conditions of each distillation tower, the circulating water flow rate and other parameters are also the same. The difference is the operating temperature of the separation tank F1, so the parameters related to the refined phenol obtained are also the same and are not listed again.

[0091] Some operating parameters and test results of the above-mentioned Examples 6-7 and Comparative Examples 4-5 are shown in the following table:

[0092] Table 8 Partial operating parameters and test results of Examples 6-7 and Comparative Examples 4-5

[0093] .

[0094] The specific test steps of all embodiments and comparative examples are as follows:

[0095] (1) All conventional properties such as temperature, pressure, flow, etc. are tested by conventional industry means, such as thermometers, pressure gauges, flow meters, etc.;

[0096] (2) The energy consumption or heat of the heat exchanger is also a common property in the industry and can be calculated. Specifically, if it is sensible heat exchange, then Q=m*cp*Δt / 3600; if it is latent heat exchange, then Q=m*r / 3600;

[0097] Q: heat (kw);

[0098] m: heat exchange medium flow rate (kg / h);

[0099] cp: specific heat (kJ / (kg·℃));

[0100] Δt: inlet and outlet temperature of heat exchange medium (℃);

[0101] r: latent heat of heat transfer medium (kJ / kg);

[0102] (3) The actual power of the compressor can be calculated by using the actual current * voltage;

[0103] (4) The purity of phenol, phenol and DPC in water was determined by liquid chromatography;

[0104] (5) DPC yield: calculated as follows: DPC yield = amount of DPC recovered / amount of DPC in the feed*100%.

[0105] It can be seen from Tables 6 to 8 that:

[0106] (1) From the results of Examples 1 to 5, it can be seen that within the process parameter range of the present invention, ideal refined phenol and refined DPC are obtained; for the different pressures at the top of the phenol refining tower T3, the main difference is the energy consumption of the distillation tower and the power consumption of the compressor C1, but the difference is not very large; the energy consumption is relatively different in the heat exchange of the economizer H1, but the feature of the present invention is that the top temperature of the phenol separation tower and the tower top temperature differ by only about 2°C, and a heat pump is used to heat the bottom stream with slightly pressurized tower steam, so the economizer H1 does not increase the energy consumption;

[0107] (2) Comparative Examples 1 and 2 respectively increase the pressure of the phenol separation tower T2. It can be seen that with the increase in pressure, the heat exchange capacity and circulating water volume of the economizer H1 are reduced to varying degrees, but the energy consumption of the reboiler and the bottom temperature of the phenol separation tower T2 are significantly increased, which also leads to the decomposition of DPC and high boiling products. It can also be seen from the results that the purity of refined phenol is greatly reduced. In addition, some unknown impurities are also precipitated together with DPC in the DPC liquid-solid separation process, resulting in a slight decrease in the purity of refined DPC.

[0108] (3) In Comparative Example 3, the pressure of the phenol separation tower T2 was reduced to 0.01 MPaG, and the operating temperature of the separation tank F1 was set to 40°C. At this time, x1 was less than x2, and the separation tank F1 was no longer stratified. Therefore, the phenol refining tower T2 no longer had the conditions for refining phenol, and the entire process was in an unstable state. In Comparative Example 3, "-" means that the relevant conditions cannot be obtained, and "*" means the relevant parameters at a certain time, which is for reference only;

[0109] (4) Different liquid-solid separation temperatures are set in Examples 6-7 and Comparative Examples 4-5. It can be seen that within the process parameter range of the present invention, the purity of the refined DPC obtained can be greater than 99.8%. If the separation temperature is too low, some other substances will also precipitate, resulting in a decrease in the purity of DPC. If the separation temperature is not greater than the solidification temperature of DPC, as shown in Comparative Example 5, DPC cannot be precipitated as a solid at 85°C.

Claims

1. A polycarbonate byproduct recovery process, characterized in that The process includes: Process 1: Light removal process: The by-product mixture from the polycarbonate production unit is mixed with circulating water and then enters the light component removal tower T1 to remove light components; Step 2: Phenol removal step: The bottom stream of the light-removal tower T1 enters the phenol separation tower T2. The vapor phase at the top of the phenol separation tower T2 is pressurized by the compressor C1 and enters the economizer H1 to heat the stream from the bottom of the phenol separation tower T2. After heat exchange and condensation, a part of the stream is returned to the top of the phenol separation tower T2 as reflux liquid, and the other part enters the phenol refining process. A part of the bottom stream of the phenol separation tower T2 enters the economizer H1 and then enters the reboiler H2. After being heated, it returns to the bottom of the phenol separation tower T2, and the other part enters the DPC refining process. Step 3: Phenol refining step: Part of the distillate from the top of the phenol separation tower T2 enters the cooler H4 for cooling, and then enters the separation tank F1 for liquid-liquid separation operation. The upper part of the separation tank F1 is a phenol-containing water phase, which is returned to the process 1 as circulating water. The lower part of the separation tank F1 is a water-containing phenol phase, which enters the phenol refining tower T3. The top of the phenol refining tower T3 is a phenol-containing water phase, which is returned to the process 1 as circulating water. The bottom of the phenol refining tower T3 produces refined phenol. Process 4: DPC refining process: Part of the material from process 2 enters the crystallization cooler H3 through the filter F2. DPC is precipitated as a solid and enters the liquid-solid separation equipment X1 for liquid-solid separation operation. The liquid phase output by the liquid-solid separation equipment X1 is the mother liquor wastewater. The solid phase output by the liquid-solid separation equipment X1 is the crude DPC product, which enters the liquid-solid washing equipment X2. At the same time, the supplemented pure water is added to the liquid-solid washing equipment X2, and the liquid-solid separation operation is performed again. The solid phase output by the liquid-solid washing equipment X2 is the refined DPC. The mass content of phenol in the overhead stream of the phenol separation tower T2 in step 2 is recorded as x1, and the mass content of phenol in the upper layer liquid in the separation tank F1 in step 3 is recorded as x2. The relationship between x1 and x2 satisfies: x1>x2; The top operating pressure of the phenol separation tower T2 in step 2 is 0.1-0.6 MPaG.

2. The polycarbonate byproduct recovery process according to claim 1, characterized in that: The relationship between x1 and x2 satisfies: 2.5*x2≥x1≥1.5*x2.

3. The polycarbonate byproduct recovery process according to claim 1, characterized in that: In the step 1, the lightness removal tower T1 is operated under positive pressure.

4. The polycarbonate byproduct recovery process according to claim 1, characterized in that: In the step 1, the mass ratio of water to DPC in the bottom stream of the light removal tower T1 is not less than 2.

5. The polycarbonate byproduct recovery process according to claim 1, characterized in that: The cooling temperature of the cooler H4 in the step 3 is -10~60°C.

6. The polycarbonate byproduct recovery process according to claim 1, characterized in that: The operating temperature of the separation tank F1 in step 3 is -10~60°C.

7. The polycarbonate byproduct recovery process according to claim 1, characterized in that: In the step 3, the phenol refining tower T3 is operated at positive pressure or slightly positive pressure.

8. The polycarbonate byproduct recovery process according to claim 1, characterized in that: The cooling temperature of the crystallization cooler H3 in step 4 is 70-78°C.

9. The polycarbonate byproduct recovery process according to any one of claims 1 to 8, characterized in that: Part of the liquid phase produced by the liquid-solid separation device X1 in step 4 is discharged as wastewater, and the other part is returned to step 1 as circulating water; The liquid phase produced by the liquid-solid washing equipment X2 in step 4 is returned to step 1 as circulating water.

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