The present invention relates to a method for producing
phenol, particularly to a method wherein the product obtained from the acid catalyzed
cleavage of cumene hydroperoxide is separated by distillation into at least a phenol-containing
fraction and an aqueous fraction comprising hydroxyacetone whereby said aqueous
fraction is treated with an oxidizing agent in presence of a base to obtain a basic
aqueous medium reduced in hydroxyacetone.
Background of the invention
The process for preparing phenol from cumene is well known.
In this process cumene is at first oxidized by air oxygen to cumene hydroperoxide.
This process step is typically called oxidation. In the second reaction step, the
so-called cleavage, the cumene hydroperoxide is cleaved to phenol and acetone using
a strong mineral acid as catalyst, for example sulfuric acid. The product from this
second reaction step, the so-called cleavage product, is then fractionated by distillation.
The purity requirements for phenol to be marketed are becoming
more and more stringent. Consequently, in order to operate a phenol production plant
economically, overall yield and selectivity to the desired end product has to be
improved and impurities formed during any of the above-described reaction steps
have to be removed as quantitatively as possible with the lowest possible loss of
the desired end product, especially phenol and acetone, at low investment and variable
costs, especially energy costs. The predominant by-products formed in the oxidation
steps are dimethylbenzyl alcohol and acetophenone. Acetophenone leaves the process
with the high-boilers from the distillation. Dimethylbenzyl alcohol is dehydrated
in the cleavage step to alpha-methylstyrene which partially forms high-boiling dimers
and cumylphenols in the acid catalyst cleavage step. The high-boilers are separated
from phenol in the distillation step. The unreacted alpha-methylstyrene is separated
and hydrogenated in order to form cumene that is recycled into the process. Depending
on the market demand, alpha-methylstyrene can also be further purified and sold
as value product. Thus, one focus in the prior art is how to operate the oxidation
step as well as the cleavage step in order to reduce the formation of these high-boilers
which can be considered as direct cumene losses. For example for the cleavage these
methods are described in
But besides these high-boilers other components are formed
in the cleavage, as for example hydroxyacetone, 2-methylbenzofuran and mesityloxide.
These so-called micro impurities are not easy to separate from phenol in the distillation.
Hydroxyacetone is the most critical component as it is nearly impossible to separate
it from phenol by simple distillation. Hydroxyacetone is typically also the micro-impurity
with the highest concentration in the product obtained from the cleavage step. The
concentration of hydroxyacetone in the cleavage product may vary between 200 and
3.000 wppm (weight parts per million).
Thus, there are great efforts in the prior art to remove
and separate hydroxyacetone from the product obtained from the cleavage step (see
). The disadvantage of all these methods is that high volume flows of cleavage
product must be processed. In addition, in
, the high volume flow of cleavage product must be treated with an oxidizing
agent that may cause enormous efforts to operate the process safely.
Prior to distillation, the cleavage product is neutralized
with a basic aqueous solution such as sodium phenate or caustic soda. The cleavage
product which is saturated with water is then worked up by distillation. A well
known method is to separate most of the hydroxyacetone with an aqueous phase which
is separated in the first distillation column while a crude phenol together with
the high-boilers is taken as the bottom product, as described in
According to the teaching of
, a portion of the aqueous phase obtained after phase separation of the
side take-off product of the first distillation column is discarded whereas the
remainder is returned to the first distillation column. Consequently, discarded
portion of the aqueous phase has to be subjected to waste water treatment according
to safety and environmental legislation. This considerably increases the costs of
running the process. Furthermore, the portion of water discarded from the system
has to be reintroduced as fresh water which is additionally a waste of resources.
Thus, it is the object of the present invention to provide a method for producing
phenol that avoids the disadvantages of the prior art discussed above and allows
for an effective reduction of hydroxyacetone from the crude phenol stream at low
investment and variable costs.
Summary of the invention
This object has been attained by a method for producing
- a) oxidizing cumene to form an oxidation product containing cumene hydroperoxide;
- b) cleaving said oxidation product using an acidic catalyst to form a cleavage
product containing phenol, acetone and impurities;
- c) neutralizing and washing said cleavage product with a basic aqueous medium
to obtain a neutralized cleavage product;
- d) separating said neutralized cleavage product by at least one distillation
step into at least a phenol containing fraction and an aqueous fraction comprising
- e) treating said aqueous fraction with an oxidizing agent in presence of a base
to obtain a basic aqueous medium reduced in hydroxyacetone;
- f) recycling at least a portion of said basic aqueous medium to the neutralizing
and washing step c); and
- g) recovering phenol from said phenol containing fraction obtained in step d).
Compared to the disclosure of
, only a low volume aqueous stream, comprising hydroxyacetone obtained
from the distillative separation of the cleavage product, has to be treated with
an oxidizing agent. Furthermore, compared to the experimental data presented in
, the residual amount of hydroxyacetone in the crude phenol stream is considerably
reduced when using the process of the present invention. Furthermore, compared to
the teaching of
, no hydroxyacetone containing aqueous stream obtained from the distillation
of the cleavage products that has to be subjected to waste water treatment is discarded
without compromising the quality of the crude phenol in terms of residual hydroxyacetone.
Detailed description of the present invention
According to the present invention, most of the hydroxyacetone
present in the cleavage product obtained from the acid catalyzed cleavage of cumene
hydroperoxide, preferably more than 90 percent, is removed with an aqueous fraction
obtained by separating the neutralized cleavage product by at least one distillation
step. The aqueous phase comprising the hydroxyacetone removed from the cleavage
product is treated with an oxidizing agent in presence of a base. Thereby, hydroxyacetone
is converted into neutralized oxidation products, for example salts of the corresponding
carboxyl-functional material, resulting in a basic aqueous medium having a reduced
hydroxyacetone content. At least a portion of said basic aqueous medium is used
for neutralizing the cleavage product.
According to a preferred embodiment, the aqueous fraction
treated with an oxidizing agent in presence of a base is completely recycled to
the step of neutralizing and washing the cleavage product. Thereby any additional
waste water stream obtained when separating the neutralized cleavage product by
distillation is avoided.
According to a preferred embodiment, in the neutralization
and washing step c) the mixture of cleavage product and aqueous medium is heterogeneous
and after the neutralizing washing step c) and prior to the separation step d) the
heterogeneous mixture is phase separated into an aqueous phase containing at least
a part of the neutralized oxidation products of hydroxyacetone and any additional
salts from the neutralization of the cleavage product and a water-saturated organic
phase that is fed to the separation step d).
Consequently, any salt material including the neutralized
oxidation products from hydroxyacetone is discharged from the system in a single
waste water stream after neutralization of the cleavage product.
According to a preferred embodiment, the base is added
to the aqueous fraction obtained from the work-up of the neutralized cleavage product
prior to the oxidizing treatment. Preferably, the base is added in an amount to
adjust the pH to be greater than 8, preferably to be between 10 and 12. Any water-soluble
base can be used according to the present invention, but is preferred if the base
is selected from aqueous NaOH and aqueous phenoxide solutions. It is particularly
preferred to use an aqueous sodium phenate solution whereby the concentration of
sodium phenate of the sodium phenate solution is preferably 5 to 50, more preferred
30 to 45 and most preferred 40 to 45 percent by weight. Such sodium phenate solutions
are generally obtained as process stream in a standard phenol plant as result of
one or a plurality of subsequent work-up steps. The use of a process stream anyway
obtained in a phenol plant has the advantage that neither fresh water nor fresh
caustic has to be introduced into the process step of the present invention. Furthermore,
phenol lost as sodium phenate in the work-up steps of a phenol plant is thereby
recycled into the process so that loss of valuable material is minimized.
The oxidizing agent to be used according to the process
of the present invention can be any suitable oxidizing agent that is capable of
converting hydroperoxide into oxidation products. Preferably, the oxidizing agent
is selected from hydrogen peroxide, oxygen, air or any mixture of oxygen and nitrogen,
whereby air is most preferred. The oxygen-containing gas, for example air, may be
either dissolved in the aqueous fraction obtained from the separating step d), as
explained above, by keeping the system under a sufficiently high pressure, or the
oxygen-containing gas is dispersed in the liquid in a gas liquid reactor. Both alternatives
of reacting an aqueous phase with a gaseous oxidizing agent are well-known in the
The temperature for treating said aqueous fraction with
an oxidizing agent may vary between 20 and 150°C, preferably between 18 and
120°C, and more preferred between 90 and 110°C. Preferably, the temperature
and residence time in the treating step is adjusted in order to convert at least
90 percent of the hydroxyacetone present in the aqueous phase obtained from the
separation step into neutralized oxidation products of hydroxyacetone.
The aqueous phase obtained from the separation step d)
of the present invention may contain residual amounts of acetone and phenol, but
it was surprisingly found that none of these components react during the oxidation
of hydroxyacetone resulting in undesired side products. Due to the recycle of the
aqueous phase in to the phenol process, any unwanted loss of valuable products,
like acetone and phenol, is avoided despite the oxidation step. According to a preferred
embodiment of the present invention, the acetone content is less than 0.1 wt.-%
, preferably in the wppm range. This can be achieved by sufficiently high separation
efficiency in the separating step d) or by treating the aqueous phase prior to oxidation
in an acetone stripper. With such a low acetone content no problems arise when contacting
the aqueous phase with any kind of oxidizing agent because any gaseous phase will
be non-explosive. This is one of the most important advantages compared to the methods
Vasileva, I.I et al. 2000, Neftepereab. Neftekhim. Moscow, Russ. Fed., 12:34-38
Another important advantage of the present invention is
that compared to the methods described in
only a relatively small volume flow of an aqueous phase must be treated,
thus keeping the reactor volume necessary for the treating step e) small resulting
in low investment costs.
Typically, depending on the hydroxyacetone content in the
cleavage product, the crude phenol obtained from the separation step d) of the present
invention contains between 50 and 400 wppm hydroxyacetone. It is preferred to further
treat the crude phenol stream obtained from the separation step d) to further reduce
the content of hydroxyacetone and other impurities.
Thus according to a preferred embodiment of the present
invention a crude phenol stream comprising methylbenzofuran and hydroxyacetone is
treated in a continuous method by passing the crude phenol stream through at least
two reactors connected in series the reactors containing an acidic ion exchange
resin, whereby the temperature in successive reactors decreases in flow direction
of the phenol stream so that the temperature in the first reactor in flow direction
of the phenol stream is between 100°C and 200°C and the temperature in
the last reactor in flow direction of the phenol stream is between 50°C and
90°C without a thermal separation step between any of two successive reactors.
The present inventors have realized that by using a plurality
of reactors containing the acidic ion exchange resin in series and importantly adjusting
a temperature profile throughout the series of reactors as defined above a crude
phenol stream can be purified to a low content of hydroxyacetone as well as methylbenzofuran
without initially removing hydroxyacetone prior to contact with the acidic ion exchange
resin and without an energy-consuming distillation step between two reactors comprising
the acidic ion exchange resin. Furthermore surprisingly, although at least two reactors
have to be used, the overall weight hourly space velocity of the process according
to the this preferred embodiment is considerably higher than for the one-step process
with the effect that the total reactor volume required according to the
present invention is even lower than for the one-step process, as disclosed in
The process for treating a crude phenol stream can be easily
integrated into the process of the resent invention.
The crude phenol that can be effectively purified by the
treating step of the preferred embodiment of the present invention contains as impurities
predominantly hydroxyacetone as well as methylbenzofuran. The concentration of hydroxyacetone
can be up to 1,000 wppm and the concentration of methylbenzofuran can be up to 200
ppm. One advantage of the present invention is that hydroxyacetone as well as methylbenzofuran
can be effectively removed even if the hydroxyacetone concentration is more than
260 wppm. Thus, a crude phenol stream comprising up to 1,000 wppm, preferably more
than 260 wppm to 1,000 wppm hydroxyacetone and up to 200 wppm, preferably 50 to
200 wppm methylbenzofuran can be successfully purified.
In addition to hydroxyacetone and methylbenzofuran further
impurities may be present:
- Mesityloxide up to 1,000 wppm,
- 2-phenylpropionaldehyde up to 500 wppm,
- methylisobutylketone up to 500 wppm,
- acetophenone up to 500 wppm,
- 3-methylcyclohexanone up to 500 wppm,
- alpha-methylstyrene up to 2,000 wppm,
- phenylbutenes up to 1,000 wppm.
These concentration ranges cover the relevant concentrations
of these components in crude phenol which is separated from acetone, cumene and
alpha-methylstyrene, water and high-boilers by distillation prior to the purification
on an ion exchange resin.
When contacting the crude phenol stream with the acidic
ion exchange resin hydroxyacetone and methylbenzofuran react to high-boilers. Mesityloxide
reacts with phenol to high-boilers and water. In the presence of water, which is
also formed by the reaction between hydroxyacetone and phenol, parts of the mesityloxide
may decompose to acetone on the acidic ion exchange resin. Acetone may further react
with phenol to Bisphenol A. Besides hydroxyacetone and mesityloxide there are other
carbonylic components which may still be present in the phenol in small amounts,
like phenylpropionaldehyde, methylisobutylketone, acetophenone and 3-methylcyclohexanone.
In addition, the phenol may have final traces of unsaturated hydrocarbons, like
alpha-methylstyrene and phenolbutenes which are undesirable components in purified
phenol. Like the carbonyl-containing components, the unsaturated hydrocarbons form
high-boilers with phenol when in contact with acidic ion exchange resins. It was
found that, even if these other impurities are present in impure phenol, the conversion
of hydroxyacetone and methylbenzofuran is not adversely effected. Furthermore, the
conversion of these additional impurity components to high-boilers is always completed
when the conversion of hydroxyacetone and methylbenzofuran is completed. Consequently,
the process of the present invention allows for the conversion of all the undesired
impurities in crude phenol to high-boilers that can be easily removed from the purified
phenol in a final distillation step after the crude phenol has been contacted with
the acidic ion exchange resin according to the process according to the present
After contact of the crude phenol with the acidic ion exchange
resin, final concentration of hydroxyacetone of less than 1 wppm and concentrations
of methylbenzofuran of less than 20 wppm, preferably less than 10 wppm, can be obtained.
As mentioned above, all other impurities are quantitatively converted to high-boilers.
Therefore, the process according to the present invention is well suited to prepare
high purity phenol. The number of reactors containing the acidic ion exchange resin
connected in series and, thus, the number of different temperature levels according
to the present invention is not particularly restricted, but taking into account
economic considerations in terms of investment costs and variable costs, a number
of two to four reactors connected in series is preferred whereby two reactors connected
in series are most preferred. Thus, according to this most preferred embodiment,
the process is conducted at two distinguished temperature levels.
Furthermore, the present inventors have found that the
deactivation of commercial ion exchange resin correlates very well with the degree
of utilization. The degree of utilization is defined as the total amount of treated
phenol which was contacted with the ion exchange resin during a certain period of
time. For a continuous plug flow reactor this is the total amount of treated phenol
per cross-sectional area of the reactor.
After a high degree of utilization the activity of the
catalyst is only some percent of that of the fresh catalyst. Surprisingly the temperature,
that is necessary to compensate the deactivation at a constant weight hourly space
velocity (WHSV), increases proportional to the degree of utilization. From practical
considerations the maximal temperature is 200°C in order to avoid any thermal
degradation of commercial ion exchange resins.
On the other hand it was found that for a phenol stream
comprising methylbenzofuran as well as considerably amounts of hydroxyacetone e.g.
up to 200 wppm methylbenzofuran and up to 1000 wppm hydroxyacetone a temperature
in the last reactor below 90°C is necessary to obtain a residual amount of
methylbenzofuran below 20 wppm, preferably below 70°C to obtain a residual
amount of methylbenzofuran below 10 wppm. From practical considerations the temperature
should not be below 50°C in order to avoid a too high reactor volume even with
One advantage of having a plurality of distinct temperature
levels for the contact of crude phenol with the acidic ion exchange resin is that
used or partly used acidic ion exchange resin can be contacted at relatively high
temperatures that for example favor the conversion of hydroxyacetone, but not the
conversion of methylbenzofuran , with the result that even with a used or partly
used catalyst due to the high temperatures a high activity of the already spent
catalyst can be maintained. On the other hand, at the low temperature level fresh
or only partly used catalyst can be employed at low temperatures favoring the conversion
of methylbenzofuran and since the catalyst is still relatively fresh, high catalyst
activity can be obtained even at low temperatures. Consequently, an optimum balance
of selectivity of the contact with the acidic ion exchange resin can be obtained
while at the same time assuring optimum activity of the catalyst resulting in comparatively
high weight hourly space velocity thereby reducing the necessary catalyst volume
for treatment of a specific phenol stream.
This synergistic effect of optimization of catalyst selectivity
with respect to hydroxyacetone and methylbenzofuran and catalyst activity depending
on the grade of deactivation of the catalyst by using the claimed temperature profile
was neither known nor derivable from the prior art.
A further advantage of the preferred embodiment of the
present invention is that if several reactors are connected in series, including
at least one spare reactor, in a continuous process completely spent catalyst can
be easily removed from the process line. The reactor with the most spent catalyst
which is at the highest temperature level and, thus, at the upstream end can be
disconnected from the line, and the reactor with fresh catalyst will enter the line
at the lowest temperature level, thus at the downstream end of the line. In the
reactor that is disconnected from the line, the spent catalyst will be either substituted
by fresh catalyst or regenerated in a separate process step in order to retain the
initial activity of the fresh catalyst. This reactivated reactor can then enter
the line at the lowest temperature level as soon as the reactor at the highest temperature
level, wherein the catalyst has been deactivated to an undesirable level, is removed
from the line. This allows for a continuous process wherein the efficiency of the
purification is approximately constant over the time resulting in a product of almost
constant specification which is extremely important for a high volume product as
It is preferred to use reactors of the same size. Thus,
at each position in the line, the WHSV for a certain phenol stream is the same and
does not change while changing the positions of the reactors in the line. The necessary
temperatures in the reactors with ion exchange resins of different activities can
easily be determined.
Furthermore, a plurality of reactors connected in parallel
can be used for every temperature level. Thus, it is very easy to adapt the treating
process to a changing throughput. Again it is preferred to use reactors of the same
size and the same number of reactors at each temperature level.
Additionally, it is possible to use a heat integration
of the phenol stream going through the reactors in order to minimize energy consumption.
For example, the phenol stream can be passed through a heat exchanger between a
first reactor and a successive second reactor using a colder phenol effluent from
a reactor located downstream from the first reactor as coolant in the heat exchanger.
This embodiment allows to cooling down the phenol stream between two successive
reactors whereas at the same time the phenol stream leaving the last reactor at
the lowest temperature level, when used as a coolant in the heat exchanger, is heated
up so that the energy consumption in the subsequent distillation step to remove
the high-boilers is reduced.
Furthermore, additional heat exchangers can be used between
two successive reactors employing conventional coolants like cooling water to adjust
the temperature of the phenol stream to the desired level.
According to one embodiment of the present invention, elongated
vessels are used as reactors whereby the vessels are preferably arranged in a vertical
orientation whereby the phenol flows from the top to the bottom of the reactor.
But it is also possible to use an upstream flow in vertical vessels or to use horizontal
According to a preferred embodiment of the present invention,
the reactors contain the acidic ion exchange resin in a fixed bed. Preferably, the
superficial liquid velocity in the fixed bed of the ion exchange resin is 0.5 to
5 mm/sec, preferable 1.0 to 3.0 mm/sec and more preferred 1.5 to 2 mm/sec.
Any acidic ion exchange resin can be used as the catalyst
according to the present invention. As used herein, the term "acidic ion exchange
resin" refers to a cation exchange resin in the hydrogen form wherein the hydrogen
ions are bound to the active sides which can be removed either by dissociation in
solution or by replacement with other positive ions. The active sides of the resins
have different attractive strengths for different ions and this selective attraction
serves as means for ion exchange. Non-limiting examples of suitable acidic ion exchange
resins include the series of sulfonated divinylbenzene crosslinked styrene copolymers,
such as for example Amberlyst 16, commercially available from Rohm & Haas, K2431,
commercially available from Lanxess, CT-151, commercially available from Purolite.
Other suitable resins can be commercially obtained from
producers such as Lanxess, Rohm and Haas Chemical Company and Dow Chemical Company.
According to the preferred embodiment of the present invention,
the temperature in the first reactor in flow direction of the phenol stream is at
least 100°C and temperature of the last reactor in flow direction of the phenol
stream is less than 90 °C, preferably less than 70°C.
The temperature in the first reactor in flow direction
of the phenol stream is 200°C at most, preferably 150°C at most, and most
preferred 120°C at most. The temperature in the last reactor in flow direction
of the phenol stream is at least 50°C.
The present invention will now be further illustrated with
reference to a specific embodiment and examples.
According to the embodiment shown in Fig. 1, the cleavage
product obtained from the acid catalyzed cleavage of cumene hydroperoxide is fed
via line 1 to a settler drum 4. Prior to entry of the cleavage product into the
settler drum the cleavage product is mixed with an aqueous phase containing the
salts from the oxidation products from hydroxyacetone and excess base, preferably
sodium phenate, which is fed to the system via line 2. If necessary, sulfuric acid
may be added via line 3 in order to adjust the pH typically to a range between 4
to 8. The resulting mixture is heterogeneous and is separated in the settler drum
4 into an aqueous phase 5 containing salts from the oxidation products of hydroxyacetone
as well as sodium sulfate, sodium formiate and salts from other organic acids, and
into an organic phase saturated with water. The aqueous phase, comprising in addition
to the above-mentioned salts small amounts of hydroxyacetone, is withdrawn for further
treatment via line 5. The water-saturated organic phase is fed to distillation column
7 via line 6. In distillation column 7 the organic phase is separated into a crude
phenol fraction removed from the bottom of the distillation column 7 via line 9,
a crude acetone fraction removed from the head of the column 7 via line 8, and a
fraction comprising water, cumene, alpha-methylstyrene and most of the hydroxyacetone
removed from the distillation column at a side take-off. Said aqueous phase is fed
via line 10 to a settler drum 11 whereby the fraction is separated into an aqueous
phase comprising hydroxyacetone and an organic phase comprising cumene and alpha-methylstyrene.
The organic phase is fed via line 12 to subsequent work-up steps. The aqueous phase
comprises, besides hydroxyacetone, small amounts of acetone, preferably less than
0.1 wt.-%, some phenol as well as some organic acids like formic acid or acetic
acid and is mixed with a base introduced via line 14 in order to increase the pH
to be above 8, preferably between 10 and 12. According to a preferred embodiment,
an aqueous sodium phenate solution within the concentration ranges discussed above
is used. Preferably, the mixing ratio of aqueous phase obtained from the distillation
column 7 and the aqueous sodium phenate solution is in the range of 1:0.05 to 1:1,
preferably between 1:0.1 to 1:0.3. Thereby a homogeneous mixture of aqueous phase
and sodium phenate solution is obtained that is fed to the reactor 15. An oxidizing
agent, preferably air, is introduced into the reactor via line 16 and temperature
and residence time in the reactor are adjusted in order to convert at least 90 percent
of the hydroxyacetone into the corresponding neutralized oxidation products. The
hydroxyacetone content in the aqueous phase obtained from the distillation step
is typically between 0.5 and 2 wt.-% . The effluent from the reactor 15 containing
the salts from the oxidation products from hydroxyacetone, excess sodium phenate
and residual amounts of hydroxyacetone is then, as discussed above, used to neutralize
the cleavage product.
Alternatively to the process described with reference to
Fig. 1, it is also possible to operate the first column without any side take-off
to get an overhead product comprising acetone, cumene, alpha-methylstyrene, water
and most of the hydroxyacetone. This overhead product can then be separated in a
subsequent pure acetone column, as shown in
, to obtain a bottom product comprising cumene, water and hydroxyacetone.
This bottom product can then be separated in a settler drum and treated, as discussed
above with reference to Fig. 1.
A cleavage product contains 42 wt.-% phenol, 26 wt.-% acetone,
25 wt.-% cumene, 3.1 wt.-% alpha-methylstyrene, 200 wppm dissolved sulfuric acid
and 1500 wppm hydroxyacetone besides other organic components. All concentrations
are related to the total amount of organic components (water-free). In addition,
1 wt.-% dissolved water is present. 10 wt.-% additional fresh water, related to
the total amount of cleavage product, must be added to the cleavage product to saturate
the cleavage product with water and to form an aqueous phase containing salts. The
salts are coming from neutralization by adding a sufficient amount of sodium sulphate
to adjust the pH in the aqueous phase to around 6. In the first distillation column
water is taken as a side draw containing 90 % of the hydroxyacetone resulting in
a concentration of about 1.23 wt.-% of hydroxyacetone in the water. The crude phenol
as the bottom product contains 340 wppm of hydroxyacetone. The water is withdrawn
from the process and sent to further treatment, amongst others towards biological
The cleavage product of the comparative example is mixed
with water from the oxidation reactor of hydroxyacetone. The water contains 0.1
wt.-% of residual hydroxyacetone and sodium phenate. The resulting concentration
of hydroxyacetone in the cleavage product is around 1590 wppm. Some sulfuric acid
is added to adjust the pH to around 6. In the first distillation column water is
taken as a side draw containing 90 % of the hydroxyacetone resulting in a concentration
of about 1.30 wt.-% of hydroxyacetone in the water. The crude phenol as the bottom
product contains 360 wppm of hydroxyacetone. The water is mixed with a 40 wt.-%
aqueous sodium phenate solution in a ratio of 1:0.1 and contacted with pure oxygen
at 95°C. The conversion rate of hydroxyacetone is 92 % thus resulting in a
residual concentration of hydroxyacetone of 0.1 wt.-%. The water is completely recycled
to the neutralization of the cleavage product.
From the comparison of the comparative example and the
example of the present invention it is evident that an additional waste water stream
is avoided without compromising the quality of the crude phenol stream in terms
of hydroxyacetone concentration.