The present invention relates to a method for treating
phenol, particularly to a continuous method for removing hydroxyacetone as well
as methylbenzofuran from a crude phenol stream.
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 distillation. Hydroxyacetone is typically also the 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
. In any case the crude phenol which will be further worked up in successive
columns will still have concentrations of hydroxyacetone, 2-methylbenzofuran and
mesityloxide of some 100 wppm which are not tolerable in pure phenol or even high
DE-AS 1 668 952
discloses a method to remove carbonyl-containing components, like mesityloxide,
isomesityloxide, methylisobuylketone, hydroxyacetone and acetophenone, from crude
phenol by passing the crude phenol over acidic ion exchange resins at temperatures
between 45 and 200°C, preferably between 80 and 150°C.
This reference also discloses the possibility to use two
different ion exchange resins whereby - when using such a catalyst combination -
different temperatures can be used in order to optimize the efficiency for each
catalyst type. This reference is totally silent with respect to the removal of methylbenzofuran
nor does this reference disclose that reactors containing the ion exchange resin
are operated using a well-defined temperature profile throughout the series of reactors.
As is evidenced by the introductory part of
DE-A 199 51 373
, the process disclosed in
DE-AS 1 668 952
is not suitable to remove low activity carbonyl compounds like methylisobutylketone
and methylcyclopentenone and compounds like methylbenzofuran . Consequently, the
process described in
DE-AS 1 668 952
does not solve the problem of removing hydroxyacetone as well as methylbenzofuran
from a crude phenol stream.
a method for purification of phenol is described wherein a phenol stream
is contacted with an ion exchange resin at temperatures between 70 and 120°C.
It is emphasized in that patent that the effectiveness of the treatment with ion
exchange resin increases with increasing temperature taking into account the range
of temperature stability for the resin. But it is furthermore evident that methylbenzofuran
can only be separated from the crude phenol stream if the initial hydroxyacetone
concentration in the crude phenol stream is below 260 wppm and preferably hydroxyacetone
is completely removed in order to have hydroxyacetone present in an amount of 10
wppm in order to effectively remove methylbenzofuran when contacting the crude phenol
stream with the ion exchange resin.
Consequently, the process described in
requires an effective removal of hydroxyacetone in upstream units prior
to the purification on the ion exchange resin. This is a disadvantage because any
removal of hydroxyacetone in upstream units prior to the purification of the ion
exchange resin means an extra effort in terms of investment costs and variable costs.
Examples of such methods are described in
DE-A 1 995 373
, taking into account the disadvantages of
DE-A 1 668 952
that 2-methylbenzofuran cannot be effectively separated by the method
disclosed therein and the disadvantage of the teaching of
that methylbenzofuran can only be effectively removed from the crude phenol
if the hydroxyacetone content of the crude phenol stream treated with the ion exchange
resin is below a certain limit, suggests a two-step process of contacting the crude
phenol stream with an ion exchange resin with a thermal separation step, like a
distillation step, between the two steps of contacting the crude phenol with an
acidic ion exchange resin. This method has the tremendous disadvantage that an additional
distillation step is necessary which considerably increases the investment and energy
costs of the entire process.
Finally, the teaching of
tries to avoid these disadvantages of multi-step purification of crude
phenol in order to remove hydroxyacetone and methylbenzofuran by using a one-step
process whereby the crude phenol is contacted with an ion exchange resin at low
temperatures between 50 and 100°C. Although this one-step process is indeed
effective in reducing the hydroxyacetone as well as the methylbenzofuran , this
method need relatively high reactor volumes and/or highly activated ion exchange
resins because of the lower reaction rate at this relatively low temperatures.
In view of the prior art as discussed above, there is still
a need for an effective and economical method for purification of crude phenol to
reduce hydroxyacetone as well as methylbenzofuran and other impurities in order
to produce high purity phenol.
Summary of the invention
This object has been surprisingly attained by a continuous
method for treating a crude phenol stream comprising methylbenzofuran and hydroxyacetone
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.
Contrary to the teachings of the prior art references 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 present invention is considerably higher than for the one-step process described
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 a process for preparation of phenol. Thus, the present invention
also relates to a process for preparation of phenol comprising:
Detailed description of the present invention
- 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) separating said cleavage product into a plurality of product streams, one
of said product streams is a crude phenol stream comprising hydroxyacetone and methylbenzofuran;
- d) treating said crude phenol stream with the method as defined above to form
a treated phenol stream reduced in hydroxyacetone and methylbenzofuran; and
- e) subjecting said treated phenol stream to a distillation to obtain a purified
The crude phenol that can be effectively purified by the
process 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 400 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.
As will be explained in more detail in the examples 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, as shown in
Fig. 1. 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 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 phenol.
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 determined by Fig. 1.
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.
The key point of the present invention is to use a certain
temperature profile throughout the series of reactors containing the acidic ion
exchange resin as defined above. According to 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 invention will now be described in more detail with
reference to specific embodiments shown in the attached figures.
Figure 1 shows the function of catalyst activity and reactor
temperature dependent on catalyst utilization as determined in example 1.
Figure 2 exhibits a diagrammatic illustration of one embodiment
according to the present invention with two reactors connected in series.
Figure 3 shows an alternative embodiment of the present
invention with three groups of three parallel reactors connected in series.
Referring to Figure 2 impure phenol is passed through a
fixed bed reactor 2 which is filled with an acidic ion exchange resin. The temperature
can be up to 200°C. After leaving the first reactor the phenol is cooled down
in the heat exchanger 3 with colder phenol 8 from the outlet of the second reactor
7. The phenol is finally cooled down to the desired low temperature in heat exchanger
4 using a suitable cooling stream 5, as for example water or any other process stream
which allows to cool down the phenol, but avoid any crystallization of the phenol
in the pipe system. The final temperature of phenol 6 is at least 50°C. Phenol
6 is then passed over the second reactor 7 to complete conversion of all relevant
impurities. The phenol stream 9 is sent to a final distillation column to separate
the high-boilers which are formed during the treatment on the acidic ion exchange
resin in reactors 2 and 7. Because of the flexibility in temperature relatively
high amounts of hydroxyacetone in addition to methylbenzofuran are tolerable in
the feed even when using spent resins with low activity in the first reactor. This
will be illustrated in the examples below in more detail.
Referring to Figure 3 impure phenol 1 is passed through
three reactors a7 to a9 in parallel at a high temperature. The phenol from these
reactors is recombined to the stream 2 and cooled down to an intermediate temperature
in the heat exchanger 3. The phenol 4 is again treated in three parallel reactors
a4 to a6. The phenol from these reactors is recombined to the stream 5 and cooled
to the final temperature in the heat exchanger 6. Phenol 7 is finally treated at
the lowest temperature in the three parallel reactors a1 to a3. Phenol stream 8
is sent to a final distillation column to separate the high-boilers which are formed
during the treatment of the acidic ion exchange resin in reactors a1 to a9. The
resins in reactors a1 to a3 have the highest activity. The resins in the reactors
a4 to a6 have a medium activity and the resins in the resins in the reactors a7
to a9 have the lowest activity. Consequently, after a certain period of time of
utilization of the resin, the position of the reactors a1 to a3 will be successively
changed to the middle row and the position of the reactors a4 to a6 will be successively
changed to the left row, as indicated by the broken line. The waste resin 10 from
the reactors a7 to a8 will be regenerated or discharged. The reactors a7 to a9 containing
regenerated or refilled fresh resin 11 are positioned to the right row as shown
by the broken line. These changes in reactor sequences can easily be done by using
an interconnecting piping network between all reactors including spare reactors.
As shown in Figure 2, any heat integration between colder and warmer phenol streams
is of course possible in reactor sequences as shown in Figure 3.
In order to demonstrate the beneficial effects of the present
invention the following examples are given.
An impure phenol stream of 1 kg/h with 200 wppm hydroxyacetone
is treated on an ion exchange resin type K2431 from Lanxess. The plug flow reactor
has a diameter of 0.025 m and a height of 1.2 m. Thus the weight hourly space velocity
WHSV (tons of phenol per hour and per m3) in the plug flow reactor was
1.7. The temperature was adjusted to get a final concentration of about 10 wppm
hydroxyacetone after the treatment. Fresh catalyst as well as catalyst samples from
a production unit with different degrees of utilization were used. Figure 1 shows
the temperature which is necessary to compensate the deactivation with regard to
the conversion of hydroxyacetone. The activity of the fresh catalyst maybe defined
as "1.0". A temperature of 70°C is sufficient to convert the hydroxyacetone
from 200 to 10 wppm on this fresh ion exchange resin. The temperature must be increased
to get the same conversion on ion exchange resins which are partly used. For example
the temperature must be increased to 123°C to get the same conversion of hydroxyacetone
on an ion exchange resin which has a cumulated amount of phenol of around 85000
Comparative Example 1:
Impure phenol containing 1,000 wppm hydroxyacetone and
200 wppm methylbenzofuran is continuously treated on a fresh fixed bed ion exchange
resin CT 151 from Purolite at 120°C. The weight hourly space velocity WHSV
is 3.3. After the purification step the concentration of hydroxyacetone is <
1 wppm, but the concentration of methylbenzofuran is 78 wppm and so too high. This
concentration of methylbenzofuran cannot be further reduced at this temperature
by adding a successive reactor, what means decreasing the WHSV.
Comparative Example 2
The impure phenol according to comparative example 1 is
continuously treated on a fresh fixed bed ion exchange resin CT 151 from Purolite
at 70°C. The weight hourly space velocity WHSV must be 0.7 to get final concentrations
of hydroxyacetone to < 1 wppm and for methylbenzofuran to < 10 wppm.
The impure phenol according to comparative example 1 using
an apparatus according to Fig. 2 is at first continuously treated on a fresh fixed
bed ion exchange resin CT 151 from Purolite at 120°C with a weight hourly space
velocity WHSV of 10. After the reactor the phenol is cooled down to 70°C and
continuously treated on a fresh fixed bed ion exchange resin CT 151 from Purolite
at 70°C with a weight hourly space velocity WHSV of 2. The final concentration
of hydroxyacetone is < 1 wppm and for methylbenzofuran < 10 wppm. In comparison
with comparative example 2, the same specification with a methylbenzofuran concentration
< 10 wppm is achieved with less than half of the reaction volume in total.
Similarly to example 2 the impure phenol according to comparative
example 1 is at first continuously treated on a fixed bed ion exchange resin CT
151 from Purolite at 120°C. The activity is only 20 % compared to a fresh ion
exchange resin. The weight hourly space velocity WHSV is 2. Then the phenol is cooled
down to 70°C and continuously treated on a fresh fix bed ion exchange resin
at 70°C with a weight hourly space velocity WHSV of 2. The final concentration
of hydroxyacetone is < 1 wppm and for methylbenzofuran 10 wppm. In comparison
with comparative example 2, the same specification with a methylbenzofuran concentration
< 10 wppm is achieved with only 70 % of the total reaction volume. In addition,
again comparing to comparative example 2, half of the total reactor volume is filled
with a resin having a low activity of only 20 % compared to a fresh resin, thus
allowing a much longer and therefore much higher economical usage of the ion exchange