TW201605883A - Method for producing sucrose fatty acid ester - Google Patents

Abstract

Provided is a method for producing sucrose fatty acid ester having suppressed coloration and a negligible amount of residual organic solvent. A method for producing sucrose fatty acid ester that acquires sucrose fatty acid ester from a reaction mixture obtained by reacting sucrose and fatty acid lower alkyl ester in the presence of an alkali catalyst in a reaction solvent, wherein the method for producing sucrose fatty acid ester includes a predetermined extraction step, purification step, and vacuum drying step.

Description

Method for producing sucrose fatty acid ester

The present invention relates to a method for producing a sucrose fatty acid ester. More specifically, it relates to a method for producing a sucrose fatty acid ester which can obtain a sucrose fatty acid ester having suppressed coloration and an extremely low residual amount of an organic solvent.

The sucrose fatty acid ester is a compound which can be used for various purposes, and is represented by additives such as auxiliaries, foods, cosmetics, pharmaceuticals, detergents, and the like, and antistatic agents such as packaging containers. Biodegradability and surface activity.

A sucrose fatty acid ester can be obtained, for example, by reacting sucrose with a fatty acid alkyl ester in the presence of a catalyst in a reaction solvent. At this time, after the completion of the reaction, a reaction mixture containing a sucrose fatty acid ester can be obtained as a method of obtaining a sucrose fatty acid ester from the reaction mixture. For example, a method is known in which the reaction mixture is subjected to an organic solvent and water. After liquid-liquid extraction, the obtained organic solvent solution is subjected to continuous counter-current extraction with water to obtain a dimethyl sulfoxide which contains almost no reaction solvent. The organic solvent solution is distilled off from the organic solvent solution (Patent Document 1).

However, the method of Patent Document 1 has the following problem: insufficiently In addition to the problem of the organic solvent; and, since the sucrose fatty acid ester is extremely easy to color, the coloring of the sucrose fatty acid ester after drying cannot be sufficiently suppressed.

[Previous Technical Literature] (Patent Literature)

Patent Document 1: Japanese Patent Laid-Open No. Hei 7-228590

The present invention is intended to provide a method for producing a sucrose fatty acid ester, which is suppressed in coloring and has an extremely low residual amount of an organic solvent.

The inventors of the present invention have intensively studied in order to solve the above problems, and as a result, have found the following facts and have further studied the present invention: by using a vacuum drying step for obtaining a target, it is necessary to use one or more decompression steps. The dryer and the internal pressure and/or the heating temperature are subjected to a series of steps of at least two stages or more, and the internal pressure and the heating temperature of the first stage and the final stage are respectively set within a specific range, The coloring is suppressed, and the solvent residual amount is extremely low.

That is, the present invention relates to a process for producing a sucrose fatty acid ester, which comprises reacting sucrose with a fatty acid lower alkyl ester in a reaction solvent in the presence of a basic catalyst to obtain a reaction mixture, and obtaining sucrose from the reaction mixture. A fatty acid ester, and the method for producing the sucrose fatty acid ester comprises the steps of: (1) an extraction step of extracting the reaction mixture with a liquid and liquid using an organic solvent and water, and sucrose fatty acid ester Extraction to organic solvent phase; (2) a purification step of using a counter current distribution extractor having a light liquid supply port and a heavy liquid supply port to compare the specific ratio of the organic solvent phase obtained in the extraction step (1) with the solvent for purification The heavy one is supplied from the heavy liquid supply port, and the lighter specific gravity is supplied from the light liquid supply port, and the organic solvent phase and the purification solvent are subjected to countercurrent contact; and (3) the reduced pressure drying step, which is a series of reduced pressure drying. In the step, the organic solvent phase obtained in the above-mentioned purification step (2) is subjected to vacuum drying by changing the pressure inside the machine and/or the heating temperature by at least two stages or more by one or more vacuum dryers; The internal pressure of the first stage of the reduced-pressure drying step is not more than atmospheric pressure, and the heating temperature thereof is 20° C. or higher and 200° C. or lower, and the internal pressure of the final stage of the reduced-pressure drying step is 10 kPa ‧ abs or less, and The heating temperature is 20 ° C or more and 200 ° C or less.

In the vacuum drying step (3), at least one of the vacuum dryers is preferably a twin screw extruder.

In the vacuum drying step (3), the series of reduced pressure drying steps preferably consist of any of the stages 2 to 9.

In the vacuum drying step (3), the content of the sucrose fatty acid ester in the organic solvent phase obtained by the purification step (2) is preferably from 10 to 80% by weight.

Further, the present invention relates to a sucrose fatty acid ester obtained by the above production method, having a Gardner color scale of 3 or less and an organic solvent residual amount of 10 ppm or less.

According to the method for producing a sucrose fatty acid ester of the present invention, a sucrose fatty acid ester having suppressed coloration and an extremely low residual amount of an organic solvent can be obtained. more detail In addition, a sucrose fatty acid ester having a Gardner color number of 3 or less and a residual solvent amount of 10 ppm or less can be obtained.

The respective elements constituting the present invention will be described below.

(Target)

In the present invention, a sucrose fatty acid ester means an ester of sucrose and a fatty acid. As the fatty acid, a saturated or unsaturated fatty acid having 6 to 30 carbon atoms is generally used. The carbon number is preferably 8 or more. On the other hand, the carbon number is preferably 22 or less. Examples of the saturated fatty acid include hexanoic acid, octanoic acid, decanoic acid, dodecanoic acid, myristic acid, palmitic acid, octadecanoic acid, and behenic acid. Examples of the unsaturated fatty acid include octadecadienoic acid, oleic acid, octadecatrienoic acid, cis 13-docosadienoic acid, and cis-12-hydroxy-9-octadecenoic acid.

The target of the present invention has a hue of 3 or less, preferably 2 or less, and an organic solvent residual amount of 10 ppm or less, preferably 5 ppm or less.

In the sucrose fatty acid ester of the present invention, the average value of the esterification ratio of the alcohol moiety of sucrose (hereinafter referred to as "average degree of esterification") is not particularly limited, and is generally considered to be easy to color the sucrose fatty acid ester. When the degree of esterification is 6.0 or less, the method of removing the solvent of the present invention can be suitably applied.

(transesterification reaction)

In the present invention, a sucrose fatty acid ester can be produced by subjecting sucrose to a fatty acid lower alkyl ester by a transesterification reaction according to a conventional method. The transesterification reaction can be carried out under full reflux. Or, the transesterification reaction is preferably carried out by first distilling off the reaction solvent while distilling off a certain amount of the reaction solvent. It is no longer distilled and is carried out while refluxing. Thereby, the transesterification reaction can be accelerated, and the distilled reaction solvent can be recovered and reused. The amount of the reaction solvent to be distilled off is, for example, about 0.1 to 30% by weight, preferably about 1 to 20% by weight, based on the total amount of the reaction solvent.

Any commercially available product can be used for sucrose. Further, the sucrose may be a recyclable sucrose (recycled sugar) which is unreacted when the sucrose fatty acid ester is produced. The form may be a solid state or a solution state dissolved in a solvent. Further, when recycled sugar is used, it may contain a salt or the like which is produced by the neutralization step described later.

The fatty acid lower alkyl ester may, for example, be an ester of the above fatty acid and a lower aliphatic alcohol. The lower aliphatic alcohol may, for example, be a lower aliphatic alcohol having 1 to 4 carbon atoms, and specific examples thereof include methanol, ethanol, propanol, and butanol. The fatty acid lower alkyl ester may be used alone or in combination of two or more kinds in any ratio. The fatty acid lower alkyl ester is generally used in an amount of 0.1 to 15 moles, preferably 0.2 moles or more, and preferably 10 moles or less, relative to the sucrose 1 mole.

As the reaction solvent for the transesterification reaction, any reaction solvent generally used in the field can be used as long as the reaction is carried out as appropriate. Examples of the reaction solvent include trimethylamine, triethylamine, and N-. N-methylmorpholine, pyridine, quinoline, pyrazine, methylpyrazine, N,N-dimethylpiperidine, etc. Amines; indoleamines such as formamide, N,N-dimethylformamide, 2-pyrrolidone, N-methyl-2-pyrrolidone; Dimethyl sulfoxide) and the like. Among them, preferred are pyridine, N,N-dimethylformamide, dimethyl hydrazine, from thermal stability, From the viewpoint of solubility and safety of sucrose, dimethyl sulfoxide (hereinafter, abbreviated as DMSO) is preferred. The reaction solvent may be used singly or in combination of two or more.

The amount of the reaction solvent used is usually from 30 to 90% by weight, more preferably 50% by weight or more, and still more preferably 85% by weight or less in the reaction mixture. Further, the reaction solvent preferably has a water content of 0.1% by weight or less.

The transesterification reaction is carried out in the presence of a basic catalyst. If the reaction system is substantially non-aqueous, the basic catalyst is present in the reaction system in a suspended state. As the basic catalyst, for example, an alkali metal hydride (sodium hydrogencarbonate, potassium hydrogencarbonate, etc.), an alkali metal hydroxide (lithium hydroxide, sodium hydroxide, potassium hydroxide, etc.), an alkali metal salt (carbonic acid) is effective. Sodium, potassium carbonate, etc., etc., especially potassium carbonate, sodium carbonate, etc. are preferred. The basic catalyst may be used singly or in combination of two or more. The amount of the basic catalyst to be used is usually 0.1 to 2.0 parts by weight, more preferably 0.2 part by weight or more, and still more preferably 1.0 part by weight or less based on 1 part by weight of the sucrose and the fatty acid alkyl ester.

The reaction temperature of the transesterification reaction is generally preferably 80 to 120 ° C, more preferably 90 ° C or more, and still more preferably 100 ° C or less. The reaction pressure is generally 0.2 to 43 kPa, preferably 0.7 to 32 kPa. The reaction time is not particularly limited, but is usually preferably 1 to 50 hours, more preferably 2 hours or more, and still more preferably 40 hours or less.

The reaction of sucrose with a fatty acid lower alkyl ester is preferably carried out under reflux of a reaction solvent. Thereby, the alcohol of the by-product in the reaction can be easily removed to the outside of the reaction system.

a transesterification reaction by neutralizing a basic catalyst by adding an acid The reaction stopped. As the acid, generally, for example, formic acid, acetic acid, propionic acid, oxalic acid, succinic acid, benzoic acid, 2,3-dihydroxysuccinic acid, lactic acid, hydroxysuccinic acid, carbonic acid, citric acid, and cisplatin are mentioned. An organic acid such as olefinic acid; an inorganic acid such as hydrochloric acid or sulfuric acid; and among them, lactic acid, hydroxysuccinic acid, citric acid or the like is preferred from the viewpoint of the influence on the living body and the like. When the reaction is stopped, the acid is added so that the pH of the reaction mixture becomes in the range of about 5.0 to 7.5, and the pH is preferably about 5.5 to 6.5, more preferably about 6.0. The acid may be used singly or in combination of two or more.

Since sucrose has eight hydroxyl groups, the resulting sucrose fatty acid ester is theoretically a mixture of monoesters to octaesters, and the composition of the ester can be adjusted by the ratio of use of the raw materials. Therefore, the reaction mixture of the obtained sucrose and the fatty acid lower alkyl ester contains, in addition to the sucrose fatty acid ester of the target, a reaction solvent, an unreacted raw material, a basic catalyst, an acid, and the like.

(extraction step)

The extraction step in the present invention is a step of extracting the above reaction mixture into a liquid phase using an organic solvent and water, and extracting the sucrose fatty acid ester to the organic solvent phase. At this time, the reaction solvent, the unreacted raw material, the basic catalyst, the acid, the neutralized salt generated by the basic catalyst and the acid, and the like are transferred to the aqueous phase.

In the liquid-liquid extraction, when the concentration of the reaction solvent in the reaction mixture is too high, it is likely to cause insufficient transfer to the aqueous phase. Therefore, when the concentration of the reaction solvent in the reaction mixture is high, it is preferred to partially distill off a portion thereof to lower the concentration. The concentration of the preferred reaction solvent is 30% by weight or less, more preferably 20% by weight or less, in the reaction mixture. On the other hand, the lower limit of the concentration of the reaction solvent is not particularly limited as long as it is not excessively high in viscosity or hardened. It is preferably 5% by weight or more, more preferably 10% by weight or more in the reaction mixture. The distillation method is not particularly limited, and for example, batch distillation using a reduced pressure and heated stirred tank; or a thin film evaporation apparatus (also referred to as a "thin film evaporator" or a "thin film evaporator") may be used. It also includes bowl type, barrel type (vertical, horizontal, etc.), etc. The same applies to the following, and is carried out at a temperature of 60 to 150 °C under reduced pressure. The reaction solvent which is distilled off as described above can be recovered and reused in the transesterification reaction.

Liquid-liquid extraction can be carried out by adding an organic solvent (extraction solvent) and water to the reaction mixture, and stirring the mixture. Examples of the organic solvent used for the liquid-liquid extraction include an alcohol or a ketone having 4 to 10 carbon atoms, or ethyl acetate. Examples of such an alcohol include n-butanol, isobutanol, tertiary butanol, n-pentanol, isoamyl alcohol, n-hexanol, and cyclohexanol; and examples of the ketone include methyl ethyl ketone and diethyl ketone. Methyl isobutyl ketone. Among them, ethyl acetate, methyl ethyl ketone, and isobutanol are preferred.

The amount of the organic solvent to be used in the reaction mixture is usually 0.5 to 20 parts by weight, preferably 1 to 10 parts by weight, per part by weight of the sucrose fatty acid ester in the reaction mixture. On the other hand, the amount of water used is usually 0.05 to 10 parts by weight, preferably 0.5 to 2 parts by weight, per part by weight of the organic solvent. Stirring is generally sufficient for 0.5 to 4 hours. The organic solvent may be used singly or in combination of two or more. Further, in the liquid-liquid extraction, as the water, in addition to the water which is generally used, the solvent for purification which is recovered by the subsequent purification step may be used instead of the "water".

The liquid-liquid extraction is preferably carried out by adjusting the pH of the mixture of the organic solvent and water. The reason for this is: thereby inhibiting the hydrolysis of the sucrose fatty acid ester, and Moreover, it also contributes to efficient extraction into the organic solvent phase. Specifically, the pH of the mixed solution is preferably from 5.5 to 6.5. For the adjustment of the pH, an acid used for neutralizing the above basic catalyst can be suitably used.

Further, liquid-liquid extraction is preferably carried out by adding a salting-out agent from the viewpoint of stable operation. The type of the salting-out agent can be arbitrarily selected, and is preferably a salt of an acid for pH adjustment. Among such salts, preferred examples of the salt include alkali metal carbonate (potassium, sodium, etc.), alkali metal citrate (potassium, sodium, etc.), alkali metal lactate (potassium, sodium, etc.), and alkali metal sulfate. (potassium, sodium, etc.), etc. The liquid-liquid extraction is generally carried out under the conditions of a salting-out agent concentration of 50 ppm or more, and is preferably in the range of 500 to 3,000 ppm from the viewpoint of stable workability and economic efficiency.

Further, in the liquid-liquid extraction step, it is preferred to add a bleaching agent to the mixture in a desired manner to reduce the coloring thereof. Thereby, a sucrose fatty acid ester which is a target which is more suppressed in coloration can be obtained. Examples of the bleaching agent include oxidizing agents such as hydrogen peroxide, sodium chlorite, sodium hypochlorite, and ozone; and reducing agents such as sodium sulfite, potassium metabisulfite, sodium metabisulfite, and sodium thiosulfate. Further, an adsorbent such as activated carbon can also be used. The amount of the bleaching agent to be added is preferably 0.1 to 5.0 parts by weight, more preferably 0.3 parts by weight or more, and still more preferably 3.0 parts by weight or less based on 100 parts by weight of the sucrose fatty acid ester.

The liquid-liquid extraction is carried out by stirring the mixture obtained in accordance with the above operation. The stirring of the mixture is generally carried out at a temperature of 40 to 80 ° C, more preferably 50 ° C or more, and still more preferably 70 ° C or less. Further, the stirring time is usually preferably 1 to 8 hours, more preferably 2 hours or more, and still more preferably 6 hours or less.

After the end of the agitation, when the bleach has been added, the excess bleach is decomposed or the residual bleach is removed. That is, the excess will be as follows Decomposition of bleach: When an oxidizing agent is added as a bleaching agent, a reducing agent is used. Conversely, when a reducing agent is added as a bleaching agent, an oxidizing agent is used. The amount of such reducing agent or oxidizing agent used is preferably from 0.9 to 1.5 equivalents relative to the residual amount of the bleaching agent (i.e., 90 to 150% by mass in order to decompose all the residual bleaching agent). Preferably, it is 1.0 equivalent or more or 1.2 equivalent or less. Further, when an adsorbent is used as the bleaching agent, for example, the bleach remaining in the mixture is removed by filtration or the like.

After the bleaching agent is decomposed or removed, the mixture is allowed to stand to separate into an organic solvent phase and an aqueous phase, and the two phases are separately extracted. At this time, the sucrose fatty acid ester is contained in the organic solvent phase. On the other hand, the aqueous layer contains unreacted sucrose, a reaction solvent, a salt formed in the reaction step, and the like. Further, the extracted organic solvent phase is supplied to the following purification step. Further, the extracted aqueous phase can be distilled off under reduced pressure of water and an organic solvent, and the reaction solvent and unreacted sucrose (recovered sugar) are recovered and reused in the transesterification reaction.

(refining step)

In the purification step of the present invention, the organic solvent phase obtained by the extraction step is purified by a countercurrent distribution type extraction apparatus including a light liquid supply port and a heavy liquid supply port, using the solvent for purification. Among the organic solvent phase and the solvent for purification, the heavy specific gravity is supplied from the heavy liquid supply port, and the lighter specific gravity is supplied from the light liquid supply port, and the organic solvent phase and the solvent for purification are countercurrently contacted to reduce the organic a reaction solvent or the like remaining in the solvent phase.

When the purification step is started, it is preferred to adjust the pH of the solvent for purification to 5.0 to 6.0. The reason for this is: thereby inhibiting the water of the sucrose fatty acid ester Solution and help to make countercurrent distribution easier. The pH can be appropriately adjusted by neutralizing the acid used in the above basic catalyst.

The "countercurrent distribution type extraction device having a light liquid supply port and a heavy liquid supply port" as used in the present invention refers to an extraction device designed in the following manner: a lighter liquid having a light specific gravity and a heavy liquid having a heavier specific gravity. When supplied from an individual supply port, the two liquids can be brought back in an efficient countercurrent flow. In the present invention, the reaction solvent or the like remaining in the organic solvent phase is transferred to the solvent for purification by countercurrent contact, and then the organic solvent phase and the solvent for purification are separately recovered, and the recovered organic solvent is supplied as a subsequent purification step. This operation is repeated until the concentration of the reaction solvent remaining in the organic solvent phase is 1 ppm, preferably 0.5 ppm or less, more preferably 0.2 ppm or less.

The countercurrent distribution type extraction device can be of any form of vertical (tower type) and horizontal type. Further, it may be a 堰-type countercurrent distribution type extraction device, and an inside thereof may be provided for the purpose of allowing the countercurrent to proceed efficiently. Further, from the viewpoint of the efficiency of the operation, a continuous countercurrent distribution type extraction apparatus which continuously supplies the recovered organic solvent phase is preferable. As the countercurrent distribution type extraction apparatus, any countercurrent distribution type extraction apparatus generally used in the field may be used, and specific examples thereof include a perforated-plate extraction column and a ruthenium-type extraction column (for example, WINTRAY (trade name, manufactured by JGC CORPORATION), etc., baffle extraction tower, rotary disc extraction tower and other stage tower type extraction devices; spray towers, packed towers, pulsed extraction towers and other differential contact devices.

As the solvent for purification, water may be used, or a mixture of water and an organic solvent used in the above extraction step may be used. As a solvent for purification, base In the case where water is used in the above, when the purification step is continuously carried out, if a specific amount of the organic solvent used in the extraction step is mixed with water in advance, and it is used as a solvent for purification, the water in the solvent for purification may not be changed. The mixing ratio of the organic solvent is preferably carried out stably, and thus it is preferred. Further, the recovered solvent for purification can be reused as described above to replace the "water" in the extraction step.

This purification step is preferably carried out by adding a salting out agent to a solvent for purification from the viewpoint of stable operation. The type of the salting-out agent can be arbitrarily selected, and is preferably a salt of an acid for pH adjustment. Among such salts, preferred examples of the salt include alkali metal citrate (such as potassium and sodium), alkali metal lactate (such as potassium or sodium), and alkali metal sulfate (such as potassium or sodium). The purification step is generally carried out under the conditions of a salting-out agent concentration of 50 ppm or more, and is preferably in the range of 500 to 3,000 ppm from the viewpoint of stable workability and economic efficiency.

The temperature in the refining step is preferably 40 to 80 ° C from the viewpoint of suppressing decomposition and coloration of the product.

The supply amount of the heavy liquid and the light liquid in the purification step differs depending on the type of the apparatus, the type of the organic solvent, and the like. For example, when a seesaw type extraction column is used as a device and isobutanol is used as an organic solvent, the refining of the heavy liquid is performed. The ratio of the supply amount of the solvent to the isobutanol phase of the light liquid is preferably in the range of 5/1 to 1/5 (capacity ratio).

An organic solvent phase comprising a sucrose fatty acid ester can be obtained via this purification step. The organic solvent phase is composed of a sucrose fatty acid ester and an organic solvent and water.

(decompression drying step)

In the present invention, the vacuum drying step is a step of drying the organic solvent phase obtained in the above-mentioned purification step under reduced pressure, distilling off the organic solvent and water contained therein, and separately separating the sucrose fatty acid ester of the target. .

The vacuum drying step of the present invention comprises the following series of steps: the pressure in the machine and/or the heating temperature are changed by at least two stages or more by one or more vacuum dryers, and the pressure is obtained through the purification step. An organic solvent phase comprising a sucrose fatty acid ester and comprising an organic solvent and water, and drying under reduced pressure; the internal pressure of the first-stage vacuum drying step is below atmospheric pressure, and the heating temperature is 20 ° C or higher and 200 ° C Hereinafter, the internal pressure of the vacuum drying step in the final stage is 10 kPa ‧ abs or less, and the heating temperature thereof is 20 ° C or more and 200 ° C or less. Here, the "in-machine pressure" refers to the pressure of the gas phase in the vacuum dryer, and the internal pressure can be carried out according to a conventional method, for example, by a pressure gauge or the like provided in a manner of being connected to the gas phase. Determination. The pressure gauge, for example, is not limited as long as it is connected to the gas phase in the vacuum dryer, and may be, for example, in an exhaust pipe from a vacuum dryer. The "heating temperature" refers to the temperature of the following two used to heat the mixture: a heating medium (warm water, steam, oil, etc.) which flows through a jacket and coil of a vacuum dryer (coil) ); or electric heaters, etc.

The concentration of the sucrose fatty acid ester in the organic solvent phase used as the drying step is preferably 10% by weight or more, more preferably 30% by mass or more. When the concentration is less than 10% by weight, it is industrially inefficient. On the other hand, the concentration is preferably 80% by weight or less, more preferably 50% by weight or less. If the concentration exceeds 80% by weight, the viscosity will be very high, and there is a tendency that it is difficult to apply to a vacuum dryer. In addition, sucrose fatty acid is used in the drying under reduced pressure at each stage. The concentration of the sucrose fatty acid ester is adjusted by adding water as needed in such a manner that the concentration of the ester is within the above range.

In the present invention, the vacuum dryer is not particularly limited as long as it can be dried under reduced pressure while maintaining the internal pressure and the heating temperature at a specific pressure and a specific temperature, and any machine or apparatus having such a function can be used. Either one can be used for the purpose of the present invention. Specific examples of the vacuum dryer of the present invention include a batch type vacuum dryer such as a stirring tank, or a flash evaporator, a thin film evaporation device, a drum dryer, and an extruder. A continuous type vacuum dryer such as a belt dryer. The various vacuum dryers described above include various types. For example, the extruder includes a single-axis extruder, a twin-shaft extruder, and the like, and such extruders have various screw diameters or screw shapes.

As the vacuum dryer, at least one of them is preferably an extruder, and particularly preferably a twin-screw extruder. When a biaxial extruder is used, it is advantageous in suppressing the coloration of sucrose fatty acid ester and reducing the amount of residual organic solvent.

In the present specification, the "one-stage reduced-pressure drying" means drying under reduced pressure by maintaining the internal pressure and the heating temperature at a specific pressure and a specific temperature by a vacuum dryer. Here, the "one-pressure drying machine" is not limited to a single pressure-reducing dryer, for example, as in an extruder, it can be divided into two or more sections by the vacuum drying section in each section. When the pressure and temperature are individually set, the interval is representative of the "one-pressure drying machine" of the present invention.

Therefore, it is possible to carry out a series of decompression drying of two or more stages by using one or two or more apparatuses which perform at least one stage or more of decompression drying, for example, by performing the following means: Internal pressure And / or heating temperature for more than 2 stages. In this case, the internal pressure and/or the heating temperature may be changed by at least two stages or more, and as long as the condition is reached, the internal pressure and the heating temperature may not be changed in two or more consecutive stages. Here, "a series of steps" means a step (for example, a refining step or the like) other than the incorporation of reduced-pressure drying during the drying under reduced pressure.

In the drying under reduced pressure, "changing the pressure inside the machine and / or heating the temperature" means changing the "pressure" and / or "temperature" to "discontinuous change".

The number of stages is preferably any number of stages from 2 to 9 from the viewpoint of the quality of the target and the efficiency of obtaining the target.

<pressure>

In the present invention, the internal pressure in the first-stage vacuum drying step is not more than atmospheric pressure, preferably 80 kPa ‧ abs or less, more preferably 60 kPa ‧ abs or less. When the pressure exceeds atmospheric pressure, the heating temperature required to remove the solvent is too high, so that the sucrose fatty acid ester tends to be easily colored. On the other hand, the pressure is preferably 3 kPa ‧ abs or more, more preferably 4 kPa ‧ abs or more. When it is less than 3 kPa ‧ abs, there is a tendency that clogging caused by flash and liquid entrainment is likely to occur. Further, the internal pressure is preferably higher than the internal pressure in the final stage.

The internal pressure in the vacuum drying step at the final stage is 10 kPa ‧ abs or less, preferably 2 kPa ‧ abs or less, more preferably 0.6 kPa ‧ abs or less. When it is larger than 10 kPa ‧ abs, it is difficult to sufficiently reduce the residual solvent, or it is easy to color because the temperature of the sucrose fatty acid ester must be raised. On the other hand, the lower limit of the pressure is not particularly limited, and the lower the possible range, the better.

Further, in the present invention, the internal pressure of the final stage of the vacuum drying step is preferably lower than the pressure in the first stage.

<temperature>

In the present invention, the heating temperature in the first-stage vacuum drying step is 20 to 200 °C. When the temperature is more than 200 ° C, the coloring tends not to be suppressed. On the other hand, when the temperature is less than 20 ° C, it is difficult to sufficiently reduce the residual solvent. To solve this problem, for example, it is necessary to significantly make the internal pressure high. There is a tendency to increase the equipment cost. The upper limit of the heating temperature is preferably 160 ° C or lower, more preferably 140 ° C or lower. The lower limit of the temperature is preferably 60 ° C or higher, more preferably 100 ° C or higher.

On the other hand, the heating temperature in the vacuum drying step in the final stage is 20 to 200 °C. When it is more than 200 ° C, coloring tends not to be suppressed. On the other hand, when it is less than 20 ° C, it tends to be difficult to sufficiently reduce the residual solvent. The upper limit of the heating temperature is preferably 160 ° C or lower, more preferably 140 ° C or lower. The lower limit of the temperature is preferably 60 ° C or higher, more preferably 80 ° C or higher, and still more preferably 100 ° C or higher.

The internal pressure and the heating temperature in the intermediate stage between the first stage and the last stage can be appropriately set within the above pressure and temperature range. However, not only in this intermediate stage, but also in the vacuum drying step, the higher the temperature, the more the coloring tendency is increased. Therefore, the pressure and temperature used in the first stage, and the pressure and temperature to be used in the final stage are also considered. And make the appropriate decision.

<processing time, processing flow>

The treatment time and the treatment flow rate are explained for the one-stage decompression drying. The processing time when using the batch type vacuum dryer is not particularly limited, and generally The words are 0.5 to 20 hours. If it exceeds 20 hours, there will be doubts about the coloration of sucrose fatty acid esters. On the other hand, the treatment flow rate in the case of using the continuous type decompression dryer is not particularly limited depending on the size of the apparatus, and is generally 0.5 to 10,000 kg/hour.

<Other>

When an extruder (preferably a twin-screw extruder) is used as the vacuum dryer, the number of rotations of the screw may be appropriately set in accordance with other conditions, and is generally preferably 100 to 1,500 rpm.

Through the above-described vacuum drying step, the sucrose fatty acid ester of the target can be isolated.

(shredding step)

When the sucrose fatty acid ester which is isolated as described above is in the form of a block, the sucrose fatty acid ester can be made into a powder by performing a pulverization step. This pulverization can be carried out in accordance with a conventional method. For example, the sucrose fatty acid ester obtained as described above can be micronized by combining the following pulverization methods: a hammer mill, a cage mill, and an axial flow type. Grinding by a high-speed rotary grinder such as a grinder, an annular mill or a shear mill; and a pulverizer such as a supersonic jet nozzle method, a counter nozzle method or a vortex air jet mill One or more kinds of pulverization.

The sucrose fatty acid ester obtained in this manner is suppressed in coloring and the residual amount of the organic solvent is extremely low. Therefore, it can be suitably used as an additive for various uses, that is, an auxiliary agent such as a chemical reaction, a detergent, etc.; a packaging container, etc. Antistatic agents, etc.; and additives for feed, cosmetics, food, pharmaceuticals, and the like.

Further, although it is not intended to be bound by theory, it is considered that, in the present invention, when the solvent is dried under reduced pressure, in the first stage in which the concentration of the sucrose fatty acid ester is low, while maintaining a high internal pressure, Preventing the mixture containing the sucrose fatty acid ester and the solvent from boiling out, and preventing the coloration of the sucrose fatty acid by making the heating temperature within a specific range, and further, in the final stage of the higher concentration of the sucrose fatty acid, The internal pressure setting is lowered to improve the reduced-pressure drying efficiency, whereby the organic solvent can be removed while achieving both coloring inhibition and efficiency.

In this specification, the term "parts" is used to mean "parts by weight". The Gardner color number is determined in accordance with JIS K 0071 using a tetrahydrofuran solution (concentration: 200 g/L) of a sucrose fatty acid ester. The amount of "organic solvent" (ppm) in the sucrose fatty acid ester is a value obtained by gas chromatography (GC). When the "average degree of esterification" is referred to for the sucrose fatty acid ester, the value is a value obtained by the following formula.

<Method for obtaining average esterification degree>

In the formula of the average degree of esterification = X, the following OHV, AV, MwSug, and MwFa are substituted to calculate X.

(Number of OH groups in one molecule of sucrose - X) = (moles of OH groups in sample 1 g) / (moles of sucrose fatty acid ester in sample 1 g) = ((OHV-AV) / ( 1000×56.11))/{1/(MwSug+(MwFa-18)X)}

OHV: the valence of sucrose fatty acid esters

AV: acid value of sucrose fatty acid ester

MwSug: Molecular Weight of Sucrose

MwFa: the average molecular weight of the constituent fatty acids

[Examples]

The present invention will be described in detail based on the examples, but the present invention is not limited to the examples.

[Example 1] (transesterification reaction)

345.7 parts by weight of DMSO was placed in the reactor, and the total reflux was carried out until the moisture content of DMSO became 0.06 wt% under the conditions of 90 ° C and 3.7 kPa. Here, 0.26 parts by weight of potassium carbonate, 67.2 parts by weight of sucrose, and 21.4 parts by weight of methyl octadecanoate were added, and the reaction was carried out while distilling DMSO at 95 ° C and 3.7 kPa, and the amount of DMSO distilled was reached. After 42.9 parts by weight, the reaction was carried out under full reflux so that DMSO was not distilled off. After a total of 4 hours of reaction, a 90% by weight aqueous lactic acid solution was added at a pH of 6.0 to stop the reaction. The distilled DMSO is recovered and reused.

(extraction step)

In the stirring tank (hereinafter, when the stirring tank is used, it is only described as "stirring" in the table), DMSO is distilled off under reduced pressure, and the solid content of the reaction mixture is concentrated to 84.0% by weight. The distilled DMSO is recovered and reused.

Here, 240 parts by weight of a 0.1% by weight potassium sulfate aqueous solution and 240 parts by weight of isobutyl alcohol (IBA) were added, and further, a 5% by weight aqueous potassium carbonate solution was added to adjust to pH 7.8, and further, a 1.5% by weight aqueous solution of 35% by weight of hydrogen peroxide was added. Serve and stir for 4 hours. Then, a 10% by weight aqueous solution of potassium metabisulfite was used, and the hydrogen peroxide was decomposed by adding it so that the residual amount of hydrogen peroxide was 1.2 equivalents of potassium metabisulfite. The solution was adjusted to pH 2 with a 5% potassium carbonate aqueous solution, and then allowed to stand to separate into 2 layers. The upper layer (IBA phase) and the lower layer (aqueous phase) are extracted. The IBA phase is transferred to a subsequent refining step. On the other hand, the aqueous phase is provided as Example 2 in order to reuse DMSO and unreacted sucrose contained therein.

(refining step)

A porous plate extraction column (hereinafter, when a perforated plate extraction column is used, only "porous" in the table) is used, and the IBA phase obtained above is purified. In other words, the IBA phase is supplied from the light liquid supply port and the IBA phase is obtained from the heavy liquid supply port so that the temperature in the column is 60° C. and the supply amount is the IBA phase/purification solvent=2/1 (capacity ratio). Supplying a solvent for purification (water/isobutanol/potassium sulfate = 92.9/7/0.1 (weight ratio), pH=5.5), continuously distributing it in a countercurrent, and repeating the refining step so that the resulting IBA phase is superposed The medium DMSO concentration was 0.2 ppm.

(decompression drying step)

The specific amount of the IBA phase (12% by weight of sucrose octadecanoate (SE), 70% by weight of isobutanol (IBA), and 18% by weight of water) obtained above was reduced by the conditions described in Table 1. The pressure is dried, and the solvent water and isobutanol are removed. After the second stage of drying under reduced pressure, at the beginning of the process, water is appropriately added to adjust the SE concentration. (In each table, after the second stage is dried under reduced pressure, the so-called "SE concentration" is the SE concentration before the start of the vacuum drying step at this stage. ). Further, isobutanol and water which are distilled off by this step are collected for recycling (for example, reuse in an extraction step).

<machine used for drying under reduced pressure>

Stirring tank (stirring): Stirring tank with a capacity of 5 m ³ ; drum dryer (roller) with anchor blade: double drum dryer, heat transfer area 0.2 m ² ; drum diameter 200 mm flash machine ( Flash): flash tank capacity 200L, heating tube electric heating area 0.8m ² (for preheating); twin shaft extruder (double shaft): screw diameter 25mm; thin film evaporation device (film): heat transfer area 0.5m ² .

(shredding step)

The sucrose octadecanoate after drying treatment was coarsely pulverized by a high-speed rotary grinder, and further pulverized by a jet mill to prepare a sucrose octadecanoate powder.

For the sucrose octadecanoate obtained above, the amount of residual IBA was measured, and the hue was evaluated.

(Measurement of residual amount of organic solvent (IBA))

According to the method described in JECFA (Joint FAO/WHO Expert Committee on Food Additives) Determination.

(hue)

The hue of each sample of the sucrose octadecanoate obtained above was evaluated in accordance with JIS K 0071 using a tetrahydrofuran solution (concentration: 200 g/L) of a sucrose fatty acid ester according to the Gardner color number. The results are shown in Table 1.

[Embodiment 2]

(reprocessing treatment)

To 263 parts by weight of the aqueous phase of the lower layer obtained in the extraction step of Example 1, 70.5 parts by weight of DMSO was added. Heat this solution to 90 ° C at 2.7 kPa Under the conditions, water and isobutanol were removed, whereby a 35 wt% sucrose DMSO solution was obtained.

(Transesterification reaction - pulverization step)

A mixture of 280.7 parts by weight of DMSO and 100 parts by weight of the above-mentioned 35% by weight sucrose DMSO solution was used instead of 345.7 parts by weight of DMSO, and further, the amount of sucrose used was 32.2 parts by weight, and the same operation as in Example 1 was carried out, and A sucrose octadecanate powder was obtained.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 1.

[Example 3]

The amount of sucrose used was 40.0 parts by weight, the amount of methyl octadecanoate used was 172.0 parts by weight, the amount of DMSO used was set to 265.4 parts by weight, and potassium carbonate was set to 0.56 parts by weight, and the initial stage of the reaction was The DMSO was distilled off to 49.8 parts by weight, and the same operation as in Example 1 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 1.

[Example 4]

21.4 parts by weight of methyl octadecanoate were respectively changed to even fatty acid methyl esters having a carbon number of 8 to 18 (fatty acid methyl esters having carbon numbers of 8 and 10: 5% by weight, methyl dodecanoate: 70% by weight, carbon) Fatty acid methyl esters of 14, 14 and 18: 25% by weight of 18.7 parts by weight, the amount of DMSO used was changed to 333.9 parts by weight, and the amount of potassium carbonate used was changed to 0.29 parts by weight, and the same operation as in Example 1 was carried out. And obtaining a specific sucrose fatty acid ester powder.

The specific sucrose fatty acid ester obtained was subjected to the same manner as described above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 1.

[Table 1]

[Example 5]

In the extraction step, a thin film evaporation apparatus (hereinafter, when a thin film evaporation apparatus is used, only a "film" in the table) is used instead of the stirring tank, and the same operation as in the first embodiment is carried out to obtain sucrose octadecane. Acid powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 2.

[Embodiment 6]

(reprocessing treatment)

To 263 parts by weight of the aqueous phase of the lower layer obtained in the extraction step of Example 5, 70.5 parts by weight of DMSO was added. This solution was heated to 90 ° C, and water and isobutanol were removed under conditions of 2.7 kPa, whereby a 35 wt% sucrose DMSO solution was obtained.

(Transesterification reaction - pulverization step)

A mixture of 280.7 parts by weight of DMSO and 100 parts by weight of the above-mentioned 35% by weight sucrose DMSO solution was used instead of DMSO 345.7 parts by weight, and further, the amount of sucrose used was 32.2 parts by weight, and the same operation as in Example 5 was carried out, and A sucrose octadecanate powder was obtained.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 2.

[Embodiment 7]

The amount of sucrose used was 40.0 parts by weight, the amount of methyl octadecanoate used was 172.0 parts by weight, the amount of DMSO used was set to 265.4 parts by weight, and potassium carbonate was set to 0.56 parts by weight, and the initial stage of the reaction was Distilled DMSO was set to 49.8 parts by weight, and in addition, the same operation as in Example 5 was carried out to obtain cane. Sugar octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 2.

[Embodiment 8]

21.4 parts by weight of methyl octadecanoate was changed to methyl dodecanoate (purity: 70% by weight, fatty acid methyl ester having 8 to 10 carbon atoms: 5% by weight, fatty acid methyl ester having 14 to 18 carbon atoms: 25 wt%) 18.7 parts by weight, the amount of DMSO used was changed to 333.9 parts by weight, and the amount of potassium carbonate used was changed to 0.29 parts by weight. Further, the same operation as in Example 5 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 2.

[Table 2]

[Embodiment 9]

In the purification step, a seesaw type extraction column (trade name: WINTRAY, manufactured by JGC CORPORATION. Hereinafter, when a seesaw type extraction column is used, only a "plate" in the table) is used instead of the porous plate extraction column. Then, the same operation as in Example 1 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 3.

[Embodiment 10]

(reprocessing treatment)

To 263 parts by weight of the aqueous phase of the lower layer obtained in the extraction step of Example 9, 70.5 parts by weight of DMSO was added. This solution was heated to 90 ° C, and water and isobutanol were removed under conditions of 2.7 kPa, whereby a 35 wt% sucrose DMSO solution was obtained.

(Transesterification reaction - pulverization step)

A mixture of 280.7 parts by weight of DMSO and 100 parts by weight of the above-mentioned 35% by weight sucrose DMSO solution was used instead of 345.7 parts by weight of DMSO, and further, the amount of sucrose used was 32.2 parts by weight, and the same operation as in Example 9 was carried out, and A sucrose octadecanate powder was obtained.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 3.

[Example 11]

The amount of sucrose used was 40.0 parts by weight, the amount of methyl octadecanoate used was 172.0 parts by weight, the amount of DMSO used was set to 265.4 parts by weight, and potassium carbonate was set to 0.56 parts by weight, and the initial stage of the reaction was Distillation DMSO set 49.8 parts by weight, in addition, the same operation as in Example 9 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 3.

[Embodiment 12]

21.4 parts by weight of methyl octadecanoate was changed to methyl dodecanoate (purity: 70% by weight, fatty acid methyl ester having 8 to 10 carbon atoms: 5% by weight, fatty acid methyl ester having 14 to 18 carbon atoms: 25 wt%) 18.7 parts by weight, the amount of DMSO used was changed to 333.9 parts by weight, and the amount of potassium carbonate used was changed to 0.29 parts by weight. Further, the same operation as in Example 9 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 3.

[table 3]

[Example 13]

In the extraction step, a thin film evaporation apparatus was used instead of the stirring tank, and further, the same operation as in Example 9 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 4.

[Embodiment 14]

(reprocessing treatment)

To 263 parts by weight of the aqueous phase of the lower layer obtained in the extraction step of Example 13, 70.5 parts by weight of DMSO was added. This solution was heated to 90 ° C, and water and isobutanol were removed under conditions of 2.7 kPa, whereby a 35 wt% sucrose DMSO solution was obtained.

(Transesterification reaction - pulverization step)

A mixture of 280.7 parts by weight of DMSO and 100 parts by weight of the above-mentioned 35% by weight sucrose DMSO solution was used instead of 345.7 parts by weight of DMSO, and further, the amount of sucrose used was 32.2 parts by weight, and the same operation as in Example 13 was carried out, and A sucrose octadecanate powder was obtained.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 4.

[Example 15]

The amount of sucrose used was 40.0 parts by weight, the amount of methyl octadecanoate used was 172.0 parts by weight, the amount of DMSO used was set to 265.4 parts by weight, and potassium carbonate was set to 0.56 parts by weight, and the initial stage of the reaction was The DMSO was distilled off to 49.8 parts by weight, and the same operation as in Example 13 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 4.

[Example 16]

21.4 parts by weight of methyl octadecanoate was changed to methyl dodecanoate (purity: 70% by weight, fatty acid methyl ester having 8 to 10 carbon atoms: 5% by weight, fatty acid methyl ester having 14 to 18 carbon atoms: 25 wt%) 18.7 parts by weight, the amount of DMSO used was changed to 333.9 parts by weight, and the amount of potassium carbonate used was changed to 0.29 parts by weight. Further, the same operation as in Example 13 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 4.

[Table 4]

[Example 17]

In the polishing step, a baffle extraction tower (hereinafter, when the baffle extraction tower is used, only a "baffle" in the table) is used instead of the perforated plate extraction column, and the same operation as in the first embodiment is performed. A sucrose octadecanate powder was obtained.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 5.

[Embodiment 18]

(reprocessing treatment)

To 263 parts by weight of the aqueous phase of the lower layer obtained in the extraction step of Example 17, 70.5 parts by weight of DMSO was added. This solution was heated to 90 ° C, and water and isobutanol were removed under conditions of 2.7 kPa, whereby a 35 wt% sucrose DMSO solution was obtained.

(Transesterification reaction - pulverization step)

A mixture of 280.7 parts by weight of DMSO and 100 parts by weight of the above-mentioned 35% by weight sucrose DMSO solution was used instead of 345.7 parts by weight of DMSO, and further, the amount of sucrose used was set to 32.2 parts by weight, and the same operation as in Example 17 was carried out, and A sucrose octadecanate powder was obtained.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 5.

[Embodiment 19]

The amount of sucrose used was 40.0 parts by weight, the amount of methyl octadecanoate used was 172.0 parts by weight, the amount of DMSO used was set to 265.4 parts by weight, and potassium carbonate was set to 0.56 parts by weight, and the initial stage of the reaction was Distilled DMSO was set to 49.8 parts by weight, and in addition, the same operation as in Example 17 was carried out to obtain a cane. Sugar octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 5.

[Example 20]

21.4 parts by weight of methyl octadecanoate was changed to methyl dodecanoate (purity: 70% by weight, fatty acid methyl ester having 8 to 10 carbon atoms: 5% by weight, fatty acid methyl ester having 14 to 18 carbon atoms: 25 wt%) 18.7 parts by weight, the amount of DMSO used was changed to 333.9 parts by weight, and the amount of potassium carbonate used was changed to 0.29 parts by weight. Further, the same operation as in Example 17 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 5.

[table 5]

[Example 21]

In the extraction step, a thin film evaporation apparatus was used instead of the stirring tank, and further, the same operation as in Example 17 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 6.

[Example 22]

(reprocessing treatment)

To 263 parts by weight of the aqueous phase of the lower layer obtained in the extraction step of Example 21, 70.5 parts by weight of DMSO was added. This solution was heated to 90 ° C, and water and isobutanol were removed under conditions of 2.7 kPa, whereby a 35 wt% sucrose DMSO solution was obtained.

(Transesterification reaction - pulverization step)

A mixture of 280.7 parts by weight of DMSO and 100 parts by weight of the above-mentioned 35% by weight sucrose DMSO solution was used instead of 345.7 parts by weight of DMSO, and further, the amount of sucrose used was 32.2 parts by weight, and the same operation as in Example 21 was carried out, and A sucrose octadecanate powder was obtained.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 6.

[Example 23]

The amount of sucrose used was 40.0 parts by weight, the amount of methyl octadecanoate used was 172.0 parts by weight, the amount of DMSO used was set to 265.4 parts by weight, and potassium carbonate was set to 0.56 parts by weight, and the initial stage of the reaction was The DMSO was distilled off to 49.8 parts by weight, and the same operation as in Example 21 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 6.

[Example 24]

21.4 parts by weight of methyl octadecanoate was changed to methyl dodecanoate (purity: 70% by weight, fatty acid methyl ester having 8 to 10 carbon atoms: 5% by weight, fatty acid methyl ester having 14 to 18 carbon atoms: 25 wt%) 18.7 parts by weight, the amount of DMSO used was changed to 333.9 parts by weight, and the amount of potassium carbonate used was changed to 0.29 parts by weight. Further, the same operation as in Example 21 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 6.

[Table 6]

[Comparative Example 1]

In the vacuum drying step, the solvent was dried under reduced pressure to remove water and isobutanol in the solvent, and the same operation as in Example 1 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 7.

[Comparative Example 2]

In the vacuum drying step, the solvent and the isobutanol were removed under the conditions described in Table 7, and the same procedure as in Example 5 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 7.

[Comparative Example 3]

In the vacuum drying step, the solvent was dried under reduced pressure to remove water and isobutanol in the solvent, and the same operation as in Example 9 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 7.

[Comparative Example 4]

In the vacuum drying step, the solvent was dried under reduced pressure to remove water and isobutanol in the solvent, and the same operation as in Example 13 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 7.

[Comparative Example 5]

In the vacuum drying step, the solvent was dried under reduced pressure to remove water and isobutanol from the solvent, and the same operation as in Example 17 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 7.

[Comparative Example 6]

In the vacuum drying step, the solvent was dried under reduced pressure to remove water and isobutanol in the solvent, and the same operation as in Example 21 was carried out to obtain a sucrose octadecanoate powder.

The obtained sucrose octadecanoate was subjected to the same manner as above, and the amount of residual IBA was measured, and the hue was evaluated. The results are shown in Table 7.

[Industrial availability]

According to the present invention, a method for producing a sucrose fatty acid ester can be provided, in which coloring is suppressed and the amount of residual organic solvent is extremely low.

The present invention has been described in detail with reference to the specific embodiments of the present invention, and it is obvious to those of ordinary skill in the art of the invention that various changes or modifications can be made without departing from the spirit and scope of the invention.

The present application is based on Japanese Patent Application No. 2013-166980, filed on Jan.

Claims (6)

A method for producing a sucrose fatty acid ester, which comprises reacting sucrose with a fatty acid lower alkyl ester in a reaction solvent in the presence of a basic catalyst to obtain a reaction mixture, and obtaining a sucrose fatty acid ester from the reaction mixture, and The method for producing a sucrose fatty acid ester comprises the steps of: (1) an extraction step of extracting the reaction mixture into a liquid phase using an organic solvent and water, and extracting the sucrose fatty acid ester to the organic solvent phase; (2) a purification step in which the organic solvent phase obtained by the extraction step (1) and the solvent for purification are used in a countercurrent distribution type extraction apparatus including a light liquid supply port and a heavy liquid supply port, and the heavy specific gravity is heavy The liquid supply port is supplied, and the lighter specific gravity is supplied from the light liquid supply port, and the organic solvent phase and the purification solvent are subjected to countercurrent contact; and (3) the reduced pressure drying step, which is a series of reduced pressure drying steps, The organic solvent phase obtained in the above-mentioned purification step (2) is decompressed by changing the internal pressure and/or the heating temperature by at least two stages or more by one or more vacuum dryers. Drying; wherein, the internal pressure of the first stage of the vacuum drying step is below atmospheric pressure, and the heating temperature is 20 ° C or more and 200 ° C or less, and the internal pressure of the final stage of the reduced pressure drying step is 10 kPa. Below abs, and the heating temperature is 20 ° C or more and 200 ° C or less. The method for producing a sucrose fatty acid ester according to claim 1, wherein in the vacuum drying step (3), at least one of the vacuum dryers is a twin-screw extruder. The method for producing a sucrose fatty acid ester according to claim 1, wherein in the vacuum drying step (3), the series of reduced-pressure drying steps are constituted by any of the stages from the second to the ninth stages. The method for producing a sucrose fatty acid ester according to claim 2, wherein in the vacuum drying step (3), the series of reduced pressure drying steps are constituted by any of the stages from 2 to 9 stages. The method for producing a sucrose fatty acid ester according to any one of claims 1 to 4, wherein, in the vacuum drying step (3), the content of the sucrose fatty acid ester in the organic solvent phase obtained in the step (2) is refined. It is 10 to 80% by weight. A sucrose fatty acid ester obtained by the method for producing a sucrose fatty acid ester according to any one of claims 1 to 5, wherein the Gardner color number is 3 or less, and the residual amount of the organic solvent is 10 ppm or less.