Process for the continuous production of ethylene glycol monoethyl ether acetate

Ethyl acetate and an excess of ethylene glycol monoethyl ether are subjected to ester-interchange reaction at a temperature not exceeding 100.degree. C. in the presence of a strongly acidic cation exchange resin as a catalyst. The resulting reaction mixture is distilled to recover the ethylene glycol monoethyl ether acetate formed. The unchanged reactants are recycled to the reaction zone for re-use, and the by-product ethanol is recovered in a highly pure form. This method can be continuously performed easily and effectively without the need to separate the catalyst or to employ complex distilling steps and without involving other difficulties.

This invention relates to an improved process for continuously producing 
ethylene glycol monoethyl ether acetate by an ester-interchange reaction 
or transesterification between ethyl acetate and ethylene glycol monoethyl 
ether. 
Ethylene glycol monoethyl ether acetate is also called ethyl cellosolve 
acetate, and is well known as a solvent having superior properties. The 
term "cellosolve" is originally a trademark, but is used i this 
application as if it were a common noun because it is now widely accepted 
as such. This substance is extensively used as a solvent for the 
production of acrylic resin paints and also as a solvent for polyurethane 
resins, epoxy resins, nitrocellulose, etc. 
Ethyl cellosolve acetate has been conventionally produced by esterifying 
ethylene glycol monoethyl ether, or ethyl cellosolve, with acetic acid in 
the presence of an acid catalyst, and this method is still in commercial 
use. The starting ethyl cellosolve is generally produced by an addition 
reaction between ethanol and ethylene oxide, and therefore, the exxpensive 
ethanol is consumed in a considerable quantity. Furthermore, since highly 
corrosive acetic acid is used in this esterification reaction, an 
anti-corrosive apparatus is disadvantageously required in commercial 
practice. Moreover, water is formed as a by-product in this esterification 
reaction, and since an azeotrope is formed between water and ethyl 
cellosolve and between water and ethyl cellosolve acetate, the distilling 
and separating operations for the reaction mixture are very complex. 
Another defect is that at the time of neutralizing the acid catalyst and 
unreacted acetic acid prior to the distillation of the ester, the desired 
ester, which has fairly high solubility in water, dissolves in water and 
is lost in the waste water. 
Recently, methods were suggested for producing ethyl cellosolve acetate by 
an ester-interchange reaction between ethyl acetate and ethyl cellosolve 
in an attempt to eliminate the defects of the aforesaid esterification 
reaction. One of such methods is described in (1) Japanese Pat. No. 
16966/1968, and another, in (2) U.S. Pat. No. 3,700,726. 
This ester-interchange reaction is an equilibrium reaction expressed by the 
following equation (I). Ethanol formed as a by-product in this reaction 
forms an azeotrope (weight ratio at atmospheric pressure 31:69) with ethyl 
acetate present in the reaction system. 
##EQU1## 
The above substances expressed by abbreviations have the following boiling 
points at atmospheric pressure. 
______________________________________ 
EC: ethyl cellosolve 135.1.degree. C. 
EA: ethyl acetate 77.1.degree. C. 
ECA: ethyl cellosolve acetate 
156.4.degree. C. 
ET: ethanol 78.3.degree. C. 
EA/ET: ethyl acetate/ethanol azeotrope 
71.8.degree. C. 
______________________________________ 
In the following description, the above abbreviations will often be used. 
The starting EA can be commercially produced by the Tischchenko reaction of 
acetaldehyde and is available at low cost. EC, the other starting 
material, is produced by an addition reaction between ET and ethylene 
oxide as stated hereinabove. Since ET is formed as a by-product in the 
ester-interchange reaction of equation (I), the effective utilization of 
ET for the production of EC would obviate the consumption of ET as in the 
conventional esterification method described hereinabove. Since ethylene 
oxide is a very highly reactive substance, unless ET is of sufficiently 
high purity, impurities contained in ET will react with ethylene oxide to 
form by-products. Separation of these by-products is difficult, and the 
yield of the desired EC decreases. In practice, therefore, the EA/ET 
azeotrope cannot be used directly as an Et component for the production of 
EC. 
The production of ECA by the ester-interchange reaction (I) essentially 
poses trouble some problems. For example, ET occurs as a by-product, and 
forms a minimum boiling point azeotrope with EA used as a reactant. 
Because separation of the azeotrope into EA and ET is not easy, these 
components are difficult to re-use. Moreover, Ea, ET and the EA/ET 
azeotrope which are involved in this reaction have boiling points close to 
one another. These problems make it difficult to produce ECA commercially 
in an economical way. The previously suggested, publications methods 
disclosed in (1) and (2) above meet with difficulty in separating or 
regenerating the catalyst, and this makes the process steps more complex. 
No entirely satisfactory method has thus been established for the 
commercial production of ECA. 
Before disclosing the invention, we shall first briefly explain the methods 
(1) and (2) suggested heretofore. 
The method (1) is a batch method which comprises subjecting EC, produced by 
the reaction of ethylene oxide with ET, and an excess of EA (about 3 moles 
of EA per mole of EC) to ester-interchange reaction in the presence of a 
catalyst preferably at the boiling point of the EA/ET azeotrope, driving 
off the EA/ET azeotrope formed with the progress of the reaction, and 
distilling the remaining reaction product to recover ECA. The Japanese 
Patent Publication states that the azeotrope which has been driven off and 
recovered is reacted with ethylene oxide as such to form EC, or the 
azeotrope is separated into EA and ET by extractive distillation with a 
suitable third component, and the separated ET is reacted with ethylene 
oxide to form EC. However, as stated hereinabove, it is impossible in 
practice to produce EC efficiently by reacting the azeotrope directly with 
ethylene oxide. Recovery of highly pure ET by the extraction and 
distillation of the azeotrope would be by no means easy. Aluminum 
alkoxides, isopropyl titanate, phosphoric acid and p-toluenesulfonic acid 
are exemplified as catalysts for use in the ester-interchange reaction. 
These catalysts homogeneously dissolve in the reaction mixture, and 
various difficulties are encountered in the separation or regeneration of 
these catalysts after the reaction. When an acid catalyst such as 
phosphoric acid and p-toluene-sulfonic acid is used, it must be 
neutralized before the purification of ECA. When an aqueous alkaline 
solution is added to the reaction mixture for neutralization, the same 
disadvantage as in the conventional esterification reaction is caused. 
Specifically, ET and EC which are completely soluble in water, and ECA and 
EA which have a fairly high solubility in water, are lost in the waste 
water. Moreover, the organic layer containing EA has a considerable amount 
of water dissolved therein to form H.sub.2 O/ET (b.p. 78.2.degree. C.), 
H.sub.2 O/EA (b.p. 70.4.degree. C.), H.sub.2 O/ET/EA (b.p. 70.2.degree. 
C.), and EA/ET (b.p. 71.8.degree. C.) azeotropes. Hence, the operation of 
recovering EA and ET is very much complicated. 
When a metal alkoxide is used as a catalyst, problems also arise in 
post-treatment and in the activity of the catalyst. Specifically, the 
metal alkoxide catalyst is normally required to be removed as a hydroxide 
by hydrolysis with water or an acid prior to the distillation of the 
reaction mixture. Furthermore, its activity often varies from lot to lot, 
or depending upon the effect of moisture in the starting material. Thus, 
its handling and regeneration are also difficult. 
The process for the continuous production of glycol ether acetate described 
in the U.S. Pat. (2) comprises the following steps: 
(a) reacting an alkyl acetate with a glycol ether (1-3 moles of alkyl 
acetate per mole of glycol ether) in the presence of a catalyst selected 
from aluminum alkoxides, titanium alkoxides and dialkyl tin oxides, at a 
temperature of 150.degree. to 225.degree. C. and an elevated pressure of 
25 to 150 psia; 
(b) directing the equilibrium of the reaction to produce the desired glycol 
ether acetate by continuous distillative removal of the by-product alcohol 
from the reaction zone; 
(c) withdrawing from the reaction zone a stream comprising catalyst and 
glycol ether acetate; 
(d) separating the glycol ether acetate from the catalyst by flash 
distillation below 225.degree. C.; and 
(e) purifying the glycol ether acetate by distillation. 
The main conditions in this process are to use the starting alkyl acetate 
in an amount equivalent to, or usually in excess of, the glycol ether, to 
use a homogeneous catalyst of the type described above, to perform the 
reaction at specified high temperatures and pressures, and to remove the 
by-product alcohol continuously and distillatively from the reaction zone 
together with the formed alkyl acetate/alcohol azeotrope. The use of 
homogeneous catalysts brings about the same disadvantage as described 
above with regard to the method (1). The U.S. Patent states that the 
reaction temperature and pressure are selected to minimize the feed 
acetate component of the overhead stream removed from the reactor. 
However, such high temperatures and pressures are likely to cause 
decomposition of the product, and are economically disadvantageous. Since 
this process comprises fairly complex process steps, various difficulties 
would be experienced in its commercial performance. 
It is an object of this invention to provide a process for continuously 
producing ethylene glycol monoethyl ether acetate from ethyl acetate and 
ethylene glycol monoethyl ether easily and economically without 
encountering the aforesaid various disadvantages of the prior art. 
The process of this invention is especially characterized by the fact that 
ethylene glycol monoethyl ether in an amount exceeding that of ethyl 
acetate is reacted with the latter under mild conditions, that a strongly 
acidic cation exchange resin is used as a catalyst, and that a special 
distillation system coordinated with these conditions is used. 
Thus, the present invention provides a process for continuously producing 
ethylene glycol monoethyl ether acetate by an ester-interchange reaction 
between ethylene glycol monoethyl ether and ethyl acetate, which comprises 
subjecting ethyl acetate and a stoichiometrically excess amount of 
ethylene glycol monoethyl ether to esterinterchange reaction in a reaction 
zone in the liquid phase in the presence of a strongly acidic cation 
exchange resin as a catalyst at a temperature not exceeding 100.degree. 
C., said ethyl acetate and ethylene glycol monoethyl ether being 
continuously fed into the reaction zone; continuously sending the liquid 
reaction mixture from the reaction zone to a distillation column and 
distilling it to separate it into 
(a) an overhead fraction which is an azeotrope of the unreacted ethyl 
acetate and a part of by-product ethanol, 
(b) a liquid side stream fraction from an upper stage of the column which 
contains the remaining ethanol in a high concentration, 
(c) another liquid side stream fraction from another upper stage of the 
column which contains the unreacted ethylene glycol monoethyl ether in a 
high concentration, and 
(d) a bottom fraction comprising the resulting ethylene glycol monoethyl 
ether acetate; 
recycling the fractions (a) and (c) to the reaction zone together with a 
fresh supply of ethyl acetate and ethylene glycol monoethyl ether; 
recovering the fraction (d) as a product; and sending the fraction (b) to 
another distillation system and distilling it to recover the by-product 
ethanol.

In tank (A), the EA/ET azeotrope (5) and EC (10) recycled from the 
distillation system are mixed with a fresh supply of EA (1) and EC (2). 
The resulting starting mixture (3) is continuously fed into reactor (B) 
where it is subjected to an ester-interchange reaction between EA and EC 
in the presence of a strongly acidic cation exchange resin catalyst. 
In the starting mixture, EC must be present in a stoichiometrically excess 
amount relative to EA. The suitable amount of EC is 1.1 to 10 moles, 
preferably 2 to 5 moles, per mole of EA. 
It has been found that ET in the recycled EA/ET azeotrope exerts no adverse 
effect on the ester-interchange reaction in the presence of a strongly 
acidic cation exchange resin catalyst. 
Strongly acidic cation exchange resins used as a catalyst are well known 
and generally marketed. They are usually produced by introducing sulfonic 
groups into a styrene/divinylbenzene copolymer. Those cation exchange 
resins which are known as a porous type or macroreticular type are 
especially preferable because a high rate of reaction can be obtained with 
them. The catalyst can be used in the form of a fixed or fluidized bed. 
The fixed bed of catalyst is preferred from the standpoint of preventing 
the wear of resin particles. The stream of the reactant mixture may be 
directed downward or upward with respect to the fixed catalyst bed. 
It has been found that the strongly acidic cation exchange resin catalyst 
used in the invention eshibits good activity under the reaction conditions 
of this invention, and this activity can be maintained over long periods 
of time. This catalyst has the further advantage that it can be separated 
from the reaction mixture without any difficulty because it is a solid 
catalyst insoluble in the reaction mixture as contrasted with the 
aforesaid homogeneous catalysts soluble in the reaction mixture used in 
the prior art. 
The ester-interchange reaction in reactor (B) is carried out in the liquid 
phase at a relatively low temperature not exceeding 100.degree. C., for 
example 20.degree. to 100.degree. C. Usually, temperatures of 20.degree. 
to 90.degree. C., preferably 30.degree. to 80.degree. C., are selected. At 
temperatures above 100.degree. C., the catalyst tends to be degraded, and 
the amounts of by-product ethyl ether and high-boiling impurities 
increase. The reaction pressure is not critical so long as it is 
sufficient to maintain the reaction mixture in the liquid phase. Usually 
it is normal atmospheric pressure (1 kg/cm.sup.2), but if desired, 
pressures of up to about 5 kg/cm.sup.2.absolute can be used. The contact 
time is generally 0.2 to 5 hours, preferably 1 to 3 hours. An ECA 
selectivity of as high as 98% or more can be obtained under the 
aforementioned ester-interchange reaction conditions. Since this reaction 
is an equilibrium reaction as shown by equation (I), a higher conversion 
of EA is generally obtained when the mole ratio of EC to EA is higher. The 
unchanged reactants can be utilized without any substantial loss because 
they can be recycled to the reaction zone from the distillation system. 
The catalyst-free reaction mixture (4) from reactor (B) is continuously fed 
into distillation column (C) where it is separated into the following four 
fractions 
(a) An overhead fraction (5) which is an azeotrope consisting of the 
unreacted EA and a part of by-product ET. 
(b) A liquid side stream fraction (6) containing the remaining ET in a high 
concentration. 
(c) Another liquid side stream fraction (10) containing the unreacted EC in 
a high concentration. 
(d) A bottom fraction (12) consisting almost entirely of the resulting ECA. 
The azeotrope overhead (5) is recycled to reactor (B) through tank (A). 
Since ET contained in the azeotrope does not adversely affect the reaction 
under the reaction conditions employed in this invention, it is possible 
to recycle the azeotrope directly and to re-use EA contained in it for the 
reaction. In other words, it is not necessary to separate EA from the 
azeotrope by a special means. Nor is it necessary to re-distill the 
azeotrope at a different pressure and to increase the EA concentration in 
the distillate. This is one reason why the amount of heat required for the 
process of this invention can be decreased. 
One important characteristic feature of the distilling process in this 
invention is that side stream fractions (10) and (6) are taken out in 
liquid form from the upper stages of distillation column (C). When the 
ester-interchange reaction mixture obtained by using an excess of EC is 
distilled, a liquid zone containing the unreacted EC in a high 
concentration is formed in an upper stage of the distillation column, and 
side stream fraction (10) can be withdrawn from this liquid zone. On the 
other hand, a liquid zone containing the remaining ET (the excess of ET, 
which does not form the overhead azeotrope) in a high concentration is 
formed in another upper stage which is nearer to the top of the column, 
and side stream fraction (6) can be withdrawn from this liquid zone. This 
separation of side streams is possible by distillation of the reaction 
mixture obtained by ester-interchange reaction under EC-excess conditions, 
and is impossible in a reaction under EA-excess conditions as in the prior 
art. 
The EC-rich fraction (10) withdrawn as a side stream is recycled to reactor 
(B) through tank (A) for reuse. The side stream fraction (6) containing 
the remaining ET in a high concentration, as will be described 
hereinbelow, is sent to the subsequent distillation system from which high 
purity ET can be easily recovered. 
Separation of the side stream, especially the side stream fraction (10), in 
the present invention is one reason why the amount of heat required for 
the process of this invention can be small. In contrast, if an ordinary 
distillation system of evaporation-condensation is used to recover high 
boiling EC present in a relatively large amount for re-use, a large amount 
of heat would be required accordingly. 
The bottom fraction (12) contains most of the resulting ECA and is 
recovered as the desired product. The bottom fraction sometimes contains a 
small amount of EC and traces of by-product high-boiling components. In 
this case, purified ECA can be simply obtained by using another 
distillation column. 
Another characteristic of the distilling process of this invention is that 
the ECA product is directly withdrawn from the bottoms of the first 
distillation column (C). This means that in the distilling system in 
accordance with this invention, the frequency of exposure of ECA to high 
temperatures is reduced, and therefore, there is no likelihood of its loss 
by decomposition or of the secondary inclusion of impurities. If, on the 
other hand, the ester-interchange reaction mixture is subjected to an 
ordinary distillation system and the components are successively removed 
in the order of increasing boiling points, ECA having the highest boiling 
point is carried over to the final step of distillation, and therefore, 
the frequency of its exposure to high temperatures would increase. 
The ET-rich side stream (6) contains a major amount of the remaining ET and 
small amounts of remaining EA and EC. If it is properly treated, it is 
easy to recycle EA and EC and recover remaining ET in pure form. The 
suitable treatment can be performed, for example, by using distillation 
columns (D) and (E) as shown in the accompanying drawing. 
The side stream (6) is sent to distillation column (D), and from its top, a 
small amount of the EA/ET azeotrope (9) consisting of ET and all EA 
present is recovered. The azeotrope can be combined with the overhead 
azeotrope (5) of distillation column (C) and recycled to reactor (B). The 
bottom fraction (7) of column (D) which no longer contains EA but contains 
a major amount of ET and a small amount of remaining EC is sent to 
distillation column (E) from the top of which an ET fraction (8) is 
recovered and from the bottom of which an EC fraction (11) is recovered. 
The EC fraction (11) can be combined with the side stream (10) of column 
(C) and recycled to reactor (B). The overhead fraction (8) consists of 
high purity ET because the bottom fraction (7) of the column (D) no longer 
contains EA which has a boiling point close to ET and azeotropes with ET. 
Such a highly pure ET has the advantage that it can be directly reacted 
with ethylene oxide without prior extractive distillation or other 
troublesome purifying means, and can be utilized in the production of EC 
which is a starting material in the ester-interchange reaction. 
As stated in detail hereinabove, the process for producing ECA in 
accordance with this invention is continuously performed. The main 
characteristic features of the process of this invention are that the 
ester-interchange reaction between EC and EA is carried out under mild 
conditions by using EC in molar excess, that a strongly acidic cation 
exchange resin, a solid catalyst, is used, and that the recovery of the 
product, the recycling of the unreacted materials and the recovery of 
by-products are effected by a unique distillation system coordinated with 
these reaction conditions. The advantages obtained by such a process are 
that the ester-interchange reaction can be continued stably and 
efficiently over long periods of time without involving troublesome 
problems such as the separation or regeneration of homogeneous catalyst, 
that the recovery of the reaction product and the recovery and recycling 
of the unreacted materials and by-products can be performed effectively 
and easily, and that the amount of heat required, i.e. the amount of steam 
consumed, in the distillation system can be considerably reduced. Thus, 
the process of this invention is very superior as a commercial process for 
ECA production. 
The following Examples illustrate the present invention more specifically. 
EXAMPLE 1 
ECA was continuously produced in accordance with the flow diagram shown in 
the accompanying drawing. 
A fixed bed-type reactor (B) was provided in which 200 liters of a 
sufficiently dried porous strongly acidic cation exchange resin having a 
matrix composed of a styrene/divinylbenzene copolymer [DIAION PK 228 
(H-form), a registered trademark; the degree of crosslinking 14; the 
surface area 0.15 to 0.20 m.sup.2 /g; porosity 10%] was packed as a 
catalyst in the state swollen with EC. A starting mixture obtained by 
combining 8.2 kg/hr of fresh EC and 8.0 kg/hr of fresh EA with 71.6 kg/hr 
of a mixture of EC, EA and ET (the mole ratio of EA:EC=1:4) recycled from 
the subsequent distillation system in a material feed tank (A) was fed 
into the bottom of the reactor (B) at 65.degree. C. under atmospheric 
pressure and reacted. The reaction mixture withdrawn from the top of the 
reactor was found to contain 14% of ECA (weight basis; the same basis will 
apply to other percentages), 11% of ET, 65% of EC, 9% of EA and 1% of 
other substances. Thus, the conversion of EA was 50.0%, and the ECA 
selectivity based on EC was 99.3%. 
The reaction mixture was introduced into an approximately intermediate 
stage of distillation column (C) (with 90 trays), and distilled under 
atmospheric pressure. An EA/ET azeotrope (temperature 72.degree. C.) was 
withdrawn from the top of the distillation column at a rate of 11.0 kg/hr, 
and recycled to the material feed tank (A). A liquid (temperature 
89.degree. C.) mainly containing ET was withdrawn from a portion near the 
top of the column (from the 18th tray), introduced into an intermediate 
stage of distillation column (D) (90 trays), and distilled under 
atmospheric pressure. An EA/ET azeotrope (temperature 72.degree. C.) was 
recovered from the top of the column (D) at a rate of 1.5 kg/hr, and 
similarly to the above, recycled to the material feed tank (A). The bottom 
fraction (temperature 97.degree. C.) of column (D) was fed into an 
intermediate stage of distillation column (E) (70 trays) operated under 
atmospheric pressure. ET having a purity of 100% was recovered from the 
top of the distillation column (E) at a rate of 4.2 kg/hr. From the bottom 
of distillation column (E), EC was withdrawn at a rate of 3.6 kg/hr. This 
EC was combined with a liquid (temperature 134.degree. C.) containing EC 
and small amounts of EA, ET and ECA which was withdrawn from another upper 
stage (from the 30th tray) of distillation column (C) at a rate of 51.3 
kg/hr, and recycled to the material feed tank (A). From the bottom of the 
distillation column (C), a liquid (temperature 159.degree. C.) containing 
ECA, a small amount of EC and traces of high-boiling by-products was 
withdrawn. This bottom liquid was purified to afford the ECA product at a 
rate of 11.9 kg/hr. 
The total amount of steam required for heating distillation columns (C), 
(D), and (E) was 46.7 kg/hr (1,900 Kcal as the amount of heat consumed per 
kilogram of ECA). After continuous operation for three months, there was 
hardly any appreciable degradation of the catalyst. 
EXAMPLE 2 
EA and EC were continuously reacted using the same reactor as used in 
Example 1 and 200 liters of a sufficiently dried porous strongly acidic 
cation exchange resin having a styrene/divinylbenzene copolymer matrix 
[DIAION PK 216 (H form), registered trademark; the degree of crosslinking 
8; the surface area 0.15 to 0.20 m.sup.2 /g; porosity 10%] as a catalyst. 
The reaction mixture was continuously separated into the individual 
components by a combination of reduced pressure distillation column (C) 
and atmospheric pressure distillation columns (D) and (E). 
A starting mixture (EA:EC mole ratio=1:3) obtained by combining fresh EC 
(9.8 kg/hr) and EA (9.6 kg/hr) in material feed tank (A) with a mixture 
(68.3 kg/hr) of EC, EA and ET recycled from the subsequent distillation 
system was fed into the reactor at 65.degree. C. under atmospheric 
pressure, and reacted. The reaction mixture withdrawn from the top of the 
reactor was found to contain 18% of ECA, 11% of ET, 12% of EA, 58% of EC 
and 1% of other substances. Thus, the conversion of EA was 48.0%, and the 
ECA selectivity based on EC was 98.8%. 
The reaction mixture was introduced into an approximately intermediate 
stage of distillation column (C) (with 90 trays), and distilled under 
reduced pressure (210 mmHg ab.). An EA-ET azeotrope (temperature 
40.degree. C.) was withdrawn from the top of the column (C) at a rate of 
13.4 kg/hr, and recycled to the material feed tank (A). A liquid 
(temperature 56.degree. C.) containing remaining ET was withdrawn from a 
portion (from the 18th tray) near the top of the column, and introduced 
into that tray of distillation column (D) (with 30 trays) which was 
located at a position corresponding to about one-third of the total height 
of the column from its bottom, and distilled under atmospheric pressure. 
From the top of the column (D) an EA-ET azeotrope (temperature 72.degree. 
C.) was recovered at a rate of 0.6 kg/hr, and recycled to the material 
feed tank (A) in the same manner as above. A bottom fraction (temperature 
97.degree. C.) from column (D) was fed into an intermediate stage of 
distillation column (E) (with 70 trays) operated under atmospheric 
pressure, and ET having a purity of 100% was recovered from its top at a 
rate of 5.0 kg/hr. From the bottom of the column (E), EC was withdrawn at 
a rate of 1.7 kg/hr. This EC was combined with 47.7 kg/hr of a liquid 
(temperature 102.degree. C.) containing EC and small amounts of EA, ET and 
ECA which was withdrawn from an intermediate stage (from the 30th tray) of 
the distillation column (C), and recycled to the material feed tank (A). 
From the bottom of the distillation column (C), a liquid (temperature 
124.degree. C.) containing ECA, a small amount of EC, and traces of 
high-boiling by-products was withdrawn. This liquid was purified to obtain 
ECA product at a rate of 14.3 kg/hr. 
The total amount of steam required for the heating of distillation columns 
(C), (D) and (E) was 47.4 kg/hr (1,660 Kcal as the amount of heat consumed 
per kilogram of ECA). After continuous operation for 3 months, there was 
hardly any appreciable degradation of the catalyst. 
EXAMPLE 3 
A continuous operation was performed using a small-sized apparatus. 
A fixed bed-type tubular reactor equipped with a stainless steel jacket was 
used in which 100 ml of predried EC-swollen porous strongly acidic cation 
exchange resin DIANION PK 216 was packed as a catalyst. A starting mixture 
(EA:EC mole ratio=1:3.9) was prepared in a mixer by mixing fresh EC (10.0 
g/hr) and EA (9.8 g/hr) with an EA/ET azeotrope (11.1 g/hr; ET 31% by 
weight) and EC (60.9 g/hr) recovered from the subsequent separating steps. 
The starting mixture was fed upwardly at a rate of 100 ml/hr into the 
reactor maintained at 65.degree. C. and atmospheric pressure. The reaction 
mixture discharged from the reactor was analyzed by gas chromatography in 
a conventional manner, and was found to contain 15.8% of ECA, 9.3% of ET, 
8.3% of EA, 66.3% of Ec and 0.3% of other substances. Hence, the 
conversion of EA was 55.9%, and the ECA selectivity based on EC was 98.8%. 
After continuous operation for three months, there was hardly any 
appreciable degradation of the catalyst. 
EXAMPLE 4 
A continuous operation was performed under somewhat varied conditions using 
the same type of reactor as used in Example 3 and a porous strongly acidic 
cation exchange resin having a styrene/divinylbenzene copolymer matrix 
[Amberlyst 15, registered trademark; the degree of crosslinking 16; the 
surface area 4.3 m.sup.2 /g; porosity 32%] as a catalyst. 
Fresh EC (12.0 g/hr) and EA (11.6 g/hr) were mixed in a mixer with an EA-ET 
azeotrope (17.4 g/hr; ET 31% by weight) and EC (48.4 g/hr) recovered from 
the PG,22 subsequent separating steps to provide a starting mixture 
having an EA:EC molar ratio of 1:2.5. The starting mixture was fed at a 
rate of 100 ml/hr upwardly into the reactor maintained at 70.degree. C. 
and 2 kg/cm.sup.2. The reaction mixture was found to contain 19.4% of ECA, 
12.8% of ET, 13.4% of EA, 54.1% of EC and 0.3% of other substances. Hence, 
the conversion of EA was 49.3%, and the ECA selectivity based on EC was 
98.2%. 
COMATIVE EXAMPLE 1 
In this Comparative Example, EC was used in excess of SA as in Examples 1 
to 4. However, separation of the resulting reaction mixture into the 
individual components was performed by a different distillation system 
from that used in Examples 1 to 4. This distillation system was an 
ordinary distillation system adapted to distill the components 
successively in the order of increasing boiling points and thus to recover 
ECA having the highest boiling point in the final step of distillation. It 
was found that when such a distillation system was used, the amount of 
heat consumed for distillation was larger than that in the distillation 
system of the present invention. 
The same reactor and catalyst as described in Example 1 were used. A 
starting mixture (EA:EC mole ratio=1:3) obtained by combining fresh EC 
(9.4 kg/hr) and EA (9.2 kg/hr) in the material supply tank (A) with a 
mixture of EA, EC and ET (68.0 kg/hr) recycled from the subsequent 
distillation system was fed into the bottom of the fixed bed-type reactor 
(B) at 65.degree. C. and atmospheric pressure, and reacted. The reaction 
mixture withdrawn from the top of the reactor contained 16% of ECA, 10% of 
ET, 60% of EC, 13% of EA and 1% of other substances. Thus, the conversion 
of EA was 45.2%, and the ECA selectivity based on EC was 99.0%. 
This reaction mixture was introduced into that tray of a first distillation 
column (with 70 trays) which was located at a position corresponding to 
about one-third of the total height of the column from its bottom, and 
distilled under reduced pressure (210 mmHg ab.). An EA/ET azeotrope 
(temperature 40.degree. C.) was withdrawn from the top of the column at a 
rate of 15.7 kg/hr, and recycled to the material feed tank (A). The bottom 
fraction (temperature 97.degree. C.) was fed into that tray of a second 
distillation column (with 50 trays) which was located at a position 
corresponding to about one-third of the total height of the column from 
its bottom, and distilled under atmospheric pressure. From the top of the 
second column, ET having a fairly high purity was distilled out at a rate 
of 4.8 kg/hr and recovered. In the meantime, the bottom fraction 
(temperature 149.degree. C.) from the second distillation column was 
introduced into that tray of a third distillation column (with 70 trays) 
which was located at a position corresponding to about one-third of the 
total height of the column from its bottom, and distilled at atmospheric 
pressure. A distillate (temperature 135.degree. C.) containing EC as a 
main component was withdrawn from the top of the third column at a rate of 
53.3 kg/hr, and recycled to the material feed tank (A). From the bottom of 
the third distillation column, a liquid (temperature 168.degree. C.) 
containing ECA and small amounts of EC and high-boiling by-products was 
withdrawn. It was purified to afford 13.7 kg/hr of ECA product. 
In operating the distillation columns, the overhead steam of the third 
distillation column was used to heat the first distillation column. The 
total amount of steam required to heat the three distillation columns was 
61.6 kg/hr (2,250 Kcal as the amount of heat consumed per kilogram of 
ECA). 
It will be appreciated that the amounts of heat consumed in Examples 1 and 
2 were substantially smaller than that consumed in this Comparative 
Example. 
COMATIVE EXAMPLE 2 
This example shows a reaction which was performed by using EA in excess of 
EC. The reaction mixture obtained by this reaction cannot be treated by 
the distillation system used in the present invention. Accordingly, it was 
treated by using the following three distillation columns. In this 
example, the amount of steam required to heat the distillation system was 
larger than that required in Comparative Example 1. 
The reactor and catalyst used were the same as those used in Example 1, and 
70 liters of the catalyst was packed in the reactor. 
In material feed tank (A), fresh EC (8.2 kg/hr) and EA (8.0 kg/hr) were 
combined with 47.5 kg/hr of a liquid containing EA and EC which was 
recycled from a third distillation column. The resulting mixture (EA:EC 
mole ratio=3:1) was fed into reactor (B) at 60.degree. C. and atmospheric 
pressure, and reacted. The reaction mixture withdrawn from the top of the 
reactor was found to contain 19% of ECA, 7% of ET, 59% of EA, 12% of EC 
and 3% of other substances. Hence, the conversion of EC was 52%, and the 
ECA selectivity based on EC was 99.4%. 
The reaction mixture was introduced into an intermediate stage of a first 
distillation column (with 160 trays) operated at 4260 mmHg ab.. An EA/ET 
azeotrope (temperature 126.degree. C.) distilled out from the top of the 
column at a rate of 11.6 kg/hr was subsequently introduced into an 
intermediate stage of a second distillation column (with 140 trays). An 
EA/ET azeotrope (temperature 72.degree. C.) was withdrawn from the top of 
the second distillation column at a rate of 7.4 kg/hr. This azeotrope was 
recycled to the feed section of the first distillation column. From the 
bottom of the second distillation column, excessive ET (4.2 kg/hr; 
temperature 86.degree. C.) formed due to a difference in the composition 
of azeotrope caused by a difference in pressure was separated. In the 
meantime, 59.5 kg/hr of the liquid (temperature 154.degree. C.) withdrawn 
from the bottom of the first distillation column was fed into the third 
distillation column (with 120 trays), and distilled under atmospheric 
pressure. A mixture (temperature 82.degree. C.) containing 37.9 kg/hr of 
EA and 7.5 kg/hr of EC was distilled out from the top of the third column, 
and recycled to the material feed tank (A). 
Thus, a liquid (temperature 163.degree. C.) containing ECA and small 
amounts of EC and high-boiling by-products was withdrawn from the bottom 
of the third distillation column, and purified to form ECA product at a 
rate of 11.8 kg/hr. 
The total amount of steam required for heating the three distillation 
columns was 55.2 kg/hr (2,320 Kcal as the amount of heat consumed per 
kilogram of ECA). 
Separately, a small-scale test was performed in this reaction system for 
the deterioration of the activity of the catalyst. It was found that after 
three months continuous operation, the catalyst was deteriorated.