A method for stepwise separation of fatty acids

Abstract

The application belongs to the technical field of oil by-product resource utilization, and specifically discloses a method for gradient separation of fatty acids, which does not need traditional four-decontamination pretreatment, can prepare high-purity C8-C10, C12-C14, C16 saturated, C18 monounsaturated and EPA / DHA products in gradient mode, the purity fluctuation of the products is less than or equal to 1%, the separation efficiency is high, the energy consumption is low, there is no three-waste emission, it is green and environment-friendly, is suitable for industrialized continuous green separation of bulk oil soapstock by-products, and has remarkable industrial application value.

Description

Technical Field

[0001] This invention belongs to the field of resource separation technology for oil and fat chemical by-products, and specifically relates to a method for the step-by-step separation of fatty acids. Background Technology

[0002] Fatty acids are the core raw materials of oleochemicals. Fatty acids with different carbon chain lengths and saturations have significantly different application values. Soap residue, a by-product of bulk oilseed processing, can be hydrolyzed by biological enzymatic methods to obtain mixed fatty acids with an acid value of 160-185 mgKOH / g, which are low-cost, high-value-added raw materials.

[0003] Existing fatty acid separation processes not only have narrow feedstock compatibility but also cannot achieve the stepwise separation of high-purity C8-C14 fatty acids, C16-C18 saturated / monounsaturated fatty acids, and EPA / DHA polyunsaturated fatty acids. For example, the continuous fractional distillation method for fatty acid production disclosed in CN101812373A is only applicable to crude fatty acids from hydrolyzed hardened oil and requires pretreatment with degassing and dehydration at -0.06~-0.08MPa (moisture content <3%). It is only for the separation of C16 / C18 stearic acid carbon chains, has narrow feedstock compatibility, and produces products with low purity. The method for preparing oleic acid by cryogenic separation of high-carbon fatty acids disclosed in CN111574361A is only used for the separation of C16~C18 oleic acid and produces products with low purity. Summary of the Invention

[0004] The purpose of this invention is to overcome the above-mentioned technical defects and provide a method that has wide raw material adaptability, low cost, and can achieve high-purity stepwise separation of fatty acids with different carbon chains and different saturation levels. To achieve the above objectives, the technical solution adopted by the present invention is a method for the stepwise separation of fatty acids, comprising the following steps: (1) Raw material pretreatment: The mixture of soapstock fatty acids with an acid value of 160-185 mgKOH / g is filtered through a 500-800 mesh filter membrane at 0.1-0.3 MPa. The residual oil is ≤0.5%, and no pretreatment of degumming / deacidification / decolorization / deodorization / degassing and dehydration is required. (2) Pump-driven distillation separation in the first-stage column: The filtered raw material is fed into a high-vacuum distillation column. The rectification section and the stripping section of the high-vacuum distillation column are each equipped with at least one Roots vacuum pump to control the rectification temperature at 240-260℃, the vacuum degree at 1-5kPa, the reflux ratio at 3:1-5:1, and the feed flow rate at 1-5mL / min, so as to separate C8-C10 light fatty acid components and C12+ heavy fatty acid components. (3) Pump-driven distillation separation in a two-stage column: C12+ heavy fatty acid components are separated from C16+ heavy fatty acid components by passing them through a high-vacuum distillation column with the same structure, with the temperature controlled at 270-290℃, the vacuum degree at 0.5-2kPa, the reflux ratio at 5:1-8:1, and the feed flow rate at 1-4mL / min. (4) Precise gradient freezing crystallization: C16+ heavy components are cooled at a rate of 1-5℃ / h, and a three-stage temperature control program of 30℃→10℃→-20℃ + timed heat preservation is implemented. The stirring speed is 50-150rpm, and fatty acids with different saturation are separated step by step. (5) High-speed countercurrent chromatography purification: using a solvent system of n-hexane-ethyl acetate-methanol-water (volume ratio 5:5:7:3), controlling the rotation speed at 800-1200 rpm and the mobile phase flow rate at 2-5 mL / min, and using methanol gradient elution to achieve high-purity separation of EPA / DHA; (6) Recycling: The elution solvent is recovered by distillation and recycled, and the crystallization mother liquor is returned to the crystallization process.

[0005] Preferably, in step (4), the three-stage gradient crystallization parameters are: 30±2℃ for 2-3h to separate C16 saturated fatty acids; 10±2℃ for 3-4h to separate C18 monounsaturated fatty acids; and -20±2℃ for 5-6h to obtain an EPA / DHA enrichment solution.

[0006] Preferably, the allocation coefficients of EPA and DHA in step (5) are stable at 0.8-1.5.

[0007] The beneficial effects achieved by this invention are as follows: This invention eliminates the need for traditional four-stage pretreatment, has broad raw material adaptability, and can achieve high-purity stepwise separation of fatty acids with different carbon chains and saturation levels. Specifically, the purity of C8-C10 light fatty acids is ≥95%, C12-C14 fatty acids is ≥95%, C16 saturated fatty acids are ≥97%, C18 monounsaturated fatty acids are ≥96%, and the total purity of EPA / DHA is ≥98%. The separation efficiency is improved by more than 30% compared to traditional distillation structures, with product purity fluctuations ≤1% and significant separation stability. Energy consumption and costs are greatly reduced. Combined with a solvent-crystallization mother liquor dual-circulation process, total energy consumption is reduced by 40%-41% compared to existing mechanical combined processes. Solvent recovery rate is ≥95%, crystallization mother liquor recovery rate is ≥85%, and overall production costs are reduced by 25%-40%. There is no pretreatment of acidic or alkaline wastewater or discharge of crystallization waste liquid, achieving industrial-scale production without waste, making it green and environmentally friendly. It is suitable for the industrial continuous green separation of bulk oilseed soapstock by-products and has significant industrial application value. Detailed Implementation

[0008] The present invention will be further described below with reference to specific embodiments. In the following embodiments, the raw material is a mixture of fatty acids from soapstock, a by-product of oil extraction from bulk oilseeds; the detection method is as follows: the purity of fatty acids is determined by gas chromatography, referring to GB / T 17376-2008 "Preparation of Fatty Acid Methyl Esters from Animal and Vegetable Oils" and GB / T 17377-2008 "Gas Chromatographic Analysis of Fatty Acid Methyl Esters from Animal and Vegetable Oils"; the product yield and solvent recovery rate are calculated by gravimetric method, and the process energy consumption is detected by calorimeter.

[0009] (1) Raw material preparation: Take 100 kg of fatty acid mixture from the by-product soapstock of soybean oil processing, which is obtained by enzymatic hydrolysis and separation purification of organic acids. The acid value of the fatty acid mixture is 165 mg KOH / g (within the range of 160-185 mg KOH / g). It is then filtered through a 500-mesh filter membrane (filtration pressure 0.2 MPa) to remove a small amount of residual oil. After filtration, the residual oil content in the raw material is 0.3%, and qualified raw material is obtained. (2) Pump-assisted distillation separation in the first-stage column: Qualified raw materials are fed into a high-vacuum distillation column. The rectification section and stripping section of the high-vacuum distillation column are each equipped with a Roots vacuum pump to control the distillation temperature at 250℃, the vacuum degree at 3kPa, the reflux ratio at 4:1, and the feed flow rate at 3mL / min. 3.2kg of C8-C10 light fatty acid components are collected from the top of the column with a purity of 96.5% and a yield of 88.9%. (3) Pump-assisted distillation separation in a secondary column: The C12+ heavy fatty acid component at the bottom of the column is fed into a high-vacuum distillation column (with the same structure as in step (2)). The distillation temperature is controlled at 280℃, the vacuum degree is 1kPa, the reflux ratio is 6:1, and the feed flow rate is 2mL / min. 12.8kg of C12-C14 fatty acid component is collected from the top of the column with a purity of 95.8% and a yield of 86.7%. (4) Gradient freeze crystallization separation: The C16+ heavy fatty acid component from the bottom of the tower was fed into a freeze crystallization device, and 1.0 kg of fatty acid seed crystals (accounting for 1.0% of the raw material mass) was added. The stirring speed was controlled at 100 rpm and the cooling rate was 2℃ / h. After holding at 30℃ for 2.5 h, 38.5 kg of C16 saturated fatty acid was obtained by centrifugation, with a purity of 97.2% and a yield of 87.5%. The temperature was further reduced to 10℃ and held for 4 h. After centrifugation, 32.8 kg of C18 monounsaturated fatty acid was obtained, with a purity of 96.1% and a yield of 85.8%. The temperature was further reduced to -20℃ and held for 6 h to obtain 10.2 kg of C18 polyunsaturated fatty acid enrichment solution, with a total purity of 90.5% for EPA and DHA. (5) High-speed countercurrent chromatography purification and separation: The C18 polyunsaturated fatty acid enrichment solution was fed into a high-speed countercurrent chromatograph. A two-phase solvent system of n-hexane-ethyl acetate-methanol-water (volume ratio 5:5:7:3) was used. The rotation speed was controlled at 1000 rpm and the mobile phase flow rate was 3 mL / min. The elution program was used to gradually increase the methanol ratio from 30% to 50%. After elution and purification, 0.8 kg of EPA product (purity 98.3%) and 0.5 kg of DHA product (purity 98.5%) were obtained. The total yield of EPA+DHA was 81.2%. (6) Recycling index: The elution solvent is recovered and purified by distillation and then recycled to the elution process. The solvent recovery rate is 96.2% and the number of recycling times can reach 10 times. The crystallization mother liquor recovery rate is 85%, and it is directly refluxed to the gradient freeze crystallization process after recovery. (7) Comprehensive process indicators: The total energy consumption of this embodiment is 820kWh, which is 40% lower than the traditional multi-stage distillation + gradient refrigeration simple combined process. There is no waste pollutant emission related to degumming, deacidification, decolorization and deodorization pretreatment, and the comprehensive production cost is reduced by 32%.

[0010] (1) Raw material preparation: Take 100 kg of fatty acid mixture from rapeseed oil processing by-product soap foot after organic acid hydrolysis and separation purification by biological enzyme method. The acid value of the fatty acid mixture is 172 mg KOH / g (within the range of 160-185 mg KOH / g). It is simply filtered through a 600 mesh filter membrane (filtration pressure 0.25 MPa) to remove a small amount of residual oil. After filtration, the residual oil content in the raw material is 0.2%, and qualified raw material is obtained. (2) Pump-assisted distillation separation in the first-stage column: Qualified raw materials are fed into a high-vacuum distillation column. The rectification section and stripping section of the high-vacuum distillation column are each equipped with a Roots vacuum pump to control the distillation temperature at 245℃, the vacuum degree at 2kPa, the reflux ratio at 3.5:1, and the feed flow rate at 2.5mL / min. 2.8kg of C8-C10 light fatty acid components are collected from the top of the column, with a purity of 96.2% and a yield of 87.3%. (3) Pump-assisted distillation separation in a secondary column: The C12+ heavy fatty acid components at the bottom of the column are fed into a high-vacuum distillation column (with the same structure as in step (2)). The distillation temperature is controlled at 275℃, the vacuum degree is 0.8kPa, the reflux ratio is 5.5:1, and the feed flow rate is 1.8mL / min. 10.5kg of C12-C14 fatty acid components are collected from the top of the column with a purity of 95.5% and a yield of 85.2%. (4) Gradient freeze-crystallization separation: The C16+ heavy fatty acid component from the bottom of the tower was fed into a freeze-crystallization device, and 0.8 kg of fatty acid seed crystals (accounting for 0.8% of the raw material mass) were added. The stirring speed was controlled at 90 rpm and the cooling rate was 1.5℃ / h. After holding at 30℃ for 2.5 h, centrifugation was performed to obtain 35.2 kg of C16 saturated fatty acids with a purity of 97.0% and a yield of 88.1%. The temperature was further reduced to 10℃ and held for 4 h. After centrifugation, 30.2 kg of C18 monounsaturated fatty acids with a purity of 96.0% and a yield of 86.5% was obtained. The temperature was further reduced to -20℃ and held for 6 h to obtain 18.8 kg of C18 polyunsaturated fatty acid enrichment solution with a total purity of 91.2% for EPA and DHA. (5) High-speed countercurrent chromatography purification and separation: The C18 polyunsaturated fatty acid enrichment solution was fed into a high-speed countercurrent chromatograph. A two-phase solvent system of n-hexane-ethyl acetate-methanol-water (volume ratio 5:5:7:3) was used. The rotation speed was controlled at 1000 rpm and the mobile phase flow rate was 3 mL / min. The elution program was used to gradually increase the methanol ratio from 30% to 50%. After elution and purification, 4.2 kg of EPA product (purity 98.1%) and 3.8 kg of DHA product (purity 98.4%) were obtained. The total yield of EPA+DHA was 82.6%. (6) Recycling index: The elution solvent is recovered and purified by distillation and then recycled to the elution process. The solvent recovery rate is 95.8% and the number of recycling times can reach 9. The crystallization mother liquor recovery rate is 86%, and it is directly refluxed to the gradient freeze crystallization process after recovery. (7) Comprehensive process indicators: The total energy consumption of this embodiment is 805kWh, which is 41% lower than the traditional multi-stage distillation + gradient refrigeration simple combined process. There is no waste pollutant emission related to degumming, deacidification, decolorization and deodorization pretreatment, and the comprehensive production cost is reduced by 35%.

[0011] (1) Raw material preparation: Take 100 kg of the same soybean oil soap foot fatty acid mixture as in Example 1, and perform four pretreatment steps of degumming, deacidification, decolorization and deodorization according to the traditional process, generating 12.5 kg of acid and alkaline wastewater, with a pretreatment energy consumption of 215 kWh; (2) Traditional multi-stage distillation separation: A conventional distillation structure is adopted, with a first-stage distillation temperature of 250℃, a vacuum degree of 3kPa, and a reflux ratio of 4:1; a second-stage distillation temperature of 280℃, a vacuum degree of 1kPa, and a reflux ratio of 6:1. The purity of the C8-C10 light components and the C12-C14 components collected from the top of the column is 92.1% and 91.3%, respectively, with a product purity fluctuation of ≥5%. (3) Traditional gradient freezing crystallization: A simple gradient cooling process was adopted (without three-stage precise heat preservation and without stirring rate control). The cooling rate was 3℃ / h. Centrifugation was used to obtain C16 saturated fatty acids with a purity of 94.8% and C18 monounsaturated fatty acids with a purity of 93.2%. The total purity of EPA+DHA in the C18 polyunsaturated fatty acid enrichment solution was ≤89.5%. (4) Purification and separation: conventional solvent elution is used, with no gradient elution program. The purity of EPA+DHA after purification is ≤90%. (5) Recycling indicators: Solvent recovery rate 82%, no crystallization mother liquor recycling, 9.2 kg of crystallization waste liquid generated; (6) Overall process indicators: Total energy consumption is 1070kWh, which is 25% higher than that of Example 1 of the present invention; the overall production cost is 32% higher than that of Example 1 of the present invention, and continuous and stable production cannot be achieved.

[0012] It should be understood that the above description of the embodiments is quite detailed and should not be construed as a limitation on the scope of protection of this invention. Those skilled in the art, under the guidance of this invention, can make various modifications such as substitutions and simple combinations without departing from the scope of protection of the claims of this invention. The scope of protection of this invention should be determined by the appended claims.

Claims

1. A method for the stepwise separation of fatty acids, characterized in that, Includes the following steps: (1) Raw material pretreatment: A mixture of soapstock fatty acids with an acid value of 160-185 mg KOH / g is precisely filtered through a 500-800 mesh filter membrane at 0.1-0.3 MPa to control the residual oil content to ≤0.5%; (2) Pump-driven distillation separation in a first-stage column: The filtered raw material is fed into a high-vacuum distillation column. The rectification section and the stripping section of the high-vacuum distillation column are each equipped with at least one Roots vacuum pump to control the rectification temperature at 240-260℃, the vacuum degree at 1-5kPa, the reflux ratio at 3:1-5:1, and the feed flow rate at 1-5mL / min, so as to separate C8-C10 light fatty acid components and C12+ heavy fatty acid components. (3) Pump-driven distillation separation in a two-stage column: C12+ heavy fatty acid components are separated from C12-C14 components and C16+ heavy fatty acid components by passing them through a high-vacuum distillation column with the same structure, with the temperature controlled at 270-290℃, vacuum degree at 0.5-2kPa, reflux ratio at 5:1-8:1, and feed flow rate at 1-4mL / min. (4) Precise gradient freezing crystallization: C16+ heavy components are cooled at a rate of 1-5℃ / h, and a three-stage temperature control program of 30℃→10℃→-20℃ + timed heat preservation is performed. The stirring speed is 50-150rpm, and fatty acids with different saturation are separated step by step. (5) High-speed countercurrent chromatography purification: using a hexane-ethyl acetate-methanol-water solvent system, controlling the rotation speed at 800-1200 rpm and the mobile phase flow rate at 2-5 mL / min, and using methanol gradient elution to achieve high-purity separation of EPA / DHA; (6) Recycling: The elution solvent is recovered by distillation and recycled, and the crystallization mother liquor is returned to the crystallization process.

2. The method according to claim 1, characterized in that, In step (4), the three-stage gradient crystallization parameters are: 30±2℃ for 2-3h to separate C16 saturated fatty acids; 10±2℃ for 3-4h to separate C18 monounsaturated fatty acids; and -20±2℃ for 5-6h to obtain EPA / DHA enrichment solution.

3. The method according to claim 1, characterized in that, In step (5), the allocation coefficients of EPA and DHA remained stable at 0.8-1.

5.

4. The method according to claim 1, characterized in that, In step (5), the volume ratio of n-hexane-ethyl acetate-methanol-water is 5:5:7:3.