Compartmented oxidation method

An apparatus and a method for continuously oxidizing an aromatic alkyl in the liquid phase and under oxidation reaction conditions, for the purpose of producing a desired aromatic carboxylic acid product, is disclosed. A seriatim arranged array of reactor compartments, each of which is adapted to accommodate a gaseous phase as well as a liquid phase is contemplated. Means for separately adjusting the reaction mixture composition of each reaction compartment is provided.

TECHNICAL FIELD OF THE INVENTION 
The present invention is generally directed to a method and to an apparatus 
for continuously oxidizing an alkylaromatic in the liquid phase and under 
oxidation reaction conditions to produce an aromatic carboxylic acid. The 
method and the apparatus each make use of a compartmented oxidation 
reactor to produce the aromatic carboxylic acid. 
BACKGROUND OF THE INVENTION 
Continuous stirred tank reactors offer distinct advantages over batch-type 
reactors for a wide variety of hydrocarbon oxidations, such as those 
involving gas-liquid reactions as well as those involving gas-liquid-solid 
reactions. For example, it is generally the case that the capital 
investment expenses as well as the operating costs are typically 
relatively lower for a conventional continuous stirred tank reactor than 
for a conventional batch-type reactor of comparable throughput. 
There are certain instances, however, where a batch-type reactor is 
preferred over a continuous stirred tank reactor. For example, in the 
manufacture of certain aromatic carboxylic acid products, it is well known 
that the presence of the product causes inhibition of the oxidation 
reaction itself, depending upon the concentration of the products within 
the reaction mixture. Thus, in many instances involving product-inhibited 
oxidation reactions the utilization of a single stirred tank reactor 
cannot provide the high yields desired because of the presence of a 
relatively higher concentration of the inhibiting product during much of 
the oxidation reaction period. While a batch-type reactor offers one 
advantage over a continuous stirred tank reactor in this regard, in that 
the inhibiting product is typically not present in sufficient 
reaction-inhibiting concentration until towards the end of the reaction, 
certain disadvantages of utilizing batch-type reactors generally favor 
utilization of a continuous stirred tank reactor, if at all possible. For 
example, it is well known that product quality generally varies from 
batch-to-batch. Occasionally, the variation in product quality renders the 
product unacceptable for its intended purpose. Also, it is generally well 
recognized that batch-type operations typically involve relatively greater 
manpower expenses and batch-type reactors are typically larger than 
continuous reactors for the same throughput. 
While it is possible to connect a large number of continuous stirred tank 
reactors in series to achieve desired reaction kinetics, to do so is not 
economically practical from the standpoint of capital-investment and 
operating cost considerations. 
Liquid phase oxidation of an alkylaromatic is exothermic. Certain 
conventional processes for oxidizing an alkylaromatic in the liquid phase 
employ a reaction mixture which includes a solvent. In such processes, a 
desired reaction temperature is achieved by maintaining the 
oxidation-reactor internal pressure at a value such that evaporation of a 
portion of a reaction mixture occurs at a desired rate. The thus-vaporized 
portion of the reaction mixture is then passed from the oxidation reactor 
to a condenser which serves to remove heat of reaction and to condense at 
least a portion of the reaction-mixture vapor supplied thereto. The 
condensate that it produced is then typically returned to the reactor as 
reflux. 
The liquid phase oxidation of aromatic alkyls to aromatic carboxylic acid 
products is currently of significant commercial importance. It is 
accordingly highly desirable to improve the yield and quality of aromatic 
carboxylic acid products. 
SUMMARY OF THE INVENTION 
One aspect or feature of the present invention is directed to a method for 
continuously oxidizing an alkylaromatic in the liquid phase and under 
oxidation reaction conditions to produce an aromatic carboxylic acid. The 
method comprises sequentially passing at least a portion of a reaction 
mixture comprising the alkylaromatic through a plurality of a 
series-arranged walled but communicating compartments of an oxidation 
reactor under predetermined oxidation reaction conditions, from an initial 
one of the series-arranged walled plural compartments to a terminal one 
thereof, optionally introducing the remainder portion of the reaction 
mixture or one or more reaction mixture components in one of the 
series-arranged walled plural compartments downstream from the initial 
compartment, and passing the introduced remainder portion sequentially to 
the terminal compartment, while agitating the contents of each of the 
walled plural compartments to produce in each of the walled compartments 
an aromatic carboxylic acid-containing liquid phase and a gaseous phase. 
At least a portion of the walls of the series-arranged compartments are 
apertured to enable passage of the reaction mixture through the oxidation 
reactor from the initial one of the series-arranged plural compartments to 
the terminal one thereof. The liquid phase temperature in each of the 
walled plural compartments is monitored and controlled, if desired. The 
gaseous phase contained within the terminal compartment is withdrawn. At 
least a portion of the thus-withdrawn gaseous phase is then condensed to 
produce a liquid stream. At least a portion of the thus-condensed liquid 
stream is then apportionably returned to preselected ones of the walled 
plural compartments in a controlled manner so as to maintain a preselected 
liquid-phase temperature differential between the terminal and initial 
compartments of the oxidation reactor. 
Another aspect or feature of the present invention is directed to an 
apparatus for continuously oxidizing the alkylaromatic in the liquid phase 
and under oxidation-reaction conditions to produce the aromatic carboxylic 
acid product. Such an apparatus, as well as still other aspects and/or 
features of the present invention, are discussed hereinbelow.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS 
Referring to FIG. 1, there is shown a system and apparatus for continuously 
oxidizing an alkylaromatic in the liquid phase and under conditions that 
produce an aromatic carboxylic acid product. 
A generally cylindrical oxidation reactor 100A, oriented vertically along a 
central, vertical axis, includes an agitator shaft 102A that is disposed 
generally along the reactor central axis. The agitator shaft 104A is 
affixed thereto and is driven by a commerically available drive mechanism 
106A. 
Internally, reactor 100A further includes several spaced-apart, centrally 
apertured dividers or partitions 108A, each of which is disposed in a 
generally horizontal plane. The aperture in each such divider 108A is 
positioned as to enable the agitator shaft 102A to be disposed generally 
through the central portion of each such divider 108A and to allow reactor 
contents to flow between each such divider 108A and the agitator shaft 
102A. Adjacent dividers 108A are so spaced as to accommodate a set of 
impeller blades 104A to be disposed therebetween. Each such set of 
impeller blades 104A rotates generally in a horizontal plane when driven 
by drive mechanism 106A. 
The reactor 100A thus defines internally a plurality of series-arranged 
walled compartments 110A which are disposed along the central longitudinal 
axis of the reactor. 
An alkylaromatic feed stream to the reactor 100A from a source 112 is fed 
by a feed pipe 114A, and an oxidizing gas from a source 116 is fed by a 
feed pipe 118A, into the initial one of the series-arranged walled plural 
compartments in reactor 100A. In a preferred embodiment, oxidizing gas 
inlet means are added to each of the series-arranged walled compartments 
of the reactor wherein the amount of oxidizing gas added to each of said 
compartments can be independently controlled. Commercially available gas 
flow controllers may be used to achieve desired, pre-selected oxidizing 
gas flow rates for each of said compartments. For example, referring to 
the embodiment in FIG. 1, instead of running directly to the initial 
compartment as shown, line 118A would branch and run separate lines to 
each compartment shown. Each branch line, including the one to the initial 
compartment, would then have a separate commercially available gas flow 
controller for independently controlling the flow of oxidizing gas to the 
respective compartment. The alkylaromatic feed source 112 and the 
oxygen-containing gas source 116 each include separate fluid and gas 
transfer means (not shown), for passing the thus-introduced alkylaromatic 
feed stream and oxygen-containing gas to the terminal one of the plural 
compartments. 
The alkylaromatic feed and other reaction components, along with the 
oxygen-containing gas, move upwardly through the reactor 100A, passing 
from one compartment 110A to an adjacent reactor compartment 110A located 
immediately thereabove, eventually passing to a terminal one of the 
series-arranged walled plural compartments 110A. As the alkylaromatic feed 
and oxygen-containing gas pass upwardly through the plural compartments 
110A, the aromatic carboxylic acid product is produced. The internal 
pressure of the oxidation reactor 100A is preselected for producing, under 
predetermined oxidation-reaction conditions in each of the walled plural 
compartments 110A, an aromatic carboxylic acid-containing liquid phase and 
an oxygen-containing gaseous phase. Because the dividers 108A are 
apertured, adjacent compartments 110A are thus in phased communication 
with one another. 
Rotation of the impeller blades 104A about the central axis of the reactor 
100A and relative to the internal wall surface of the reactor 100A causes 
the contents of each of the walled plural compartments 110A to circulate 
about the reactor's central axis relative to the reactor inner surface. 
This, in turn, causes disentrainment of the liquid phase from the gaseous 
phase by centrifugal force within each of the walled plural compartments 
110A. However, in some or all of the reactor compartment baffles can be 
placed on the internal wall surface to suppress the rotation of the 
reactor contents and improve the mixing therein. These baffles may be of 
any size and shape. For example, while a baffle made with a circular 
cross-section could be used, this shape would be less efficient than a 
baffle with a rectangular cross-section but would allow for some rotation 
of the liquid to aid disentrainment while retaining some of the desirable 
features of baffles by providing mixing and air dispersion. 
The oxidation reaction is exothermic. To control reaction temperature, a 
portion of the product-containing reaction mixture is allowed to vaporize. 
The vaporized portion of the reaction mixture is passed from the terminal 
or downstream-most one of the series-arranged walled plural compartments 
to a condenser 120A via a discharge pipe 122A. From the terminal reactor 
compartment the product-containing reaction mixture is passed via 
discharge pipe 124A to a collection site 126 for further processing, as 
desired. 
Process vapors are condensed in condenser 120A and enter, via conduit 128A, 
separator 130A which separates non-condensable gases from the 
now-condensed process vapors. The non-condensable gases are vented via 
vent line 132A. The condensed process vapors are returned back into 
reactor 100A as reflux. 
To optimize the reaction of alkylaromatic to aromatic carboxylic acid, it 
is desirable to control the temperature of the reaction mixture internally 
along the entire length of the reactor. Accordingly, the reactor 100A 
includes a plurality of commercially available temperature transducers 
134A, each such temperature transducers 134A having a temperature sensing 
element that is disposed within the corresponding one of the plural 
reactor compartments 110A. Each one of the plural temperature transducers 
134A is operatively connected to a commercially available process 
controlled system 136A and senses the liquid-phase temperature of the 
reaction mixture within the corresponding compartment. 
Condensed process vapors from separator 130A are returned as reflux back 
into reactor 100A by transfer pump 138A. 
A plurality of return lines 140A are supplied by transfer pump 138A. Each 
one of the plural return lines 140A returns as reflux a portion or all of 
the condensed process vapors, as necessary, to a corresponding one of the 
plural walled compartments 110A. Each such reflux return line 140A 
includes an automatic flow control valve 142A, operatively connected to 
the process control system 136A, for controlling flow of condensed process 
vapors that are being refluxed to a corresponding one of the plural walled 
compartments 110A. Generally, the thus-condensed liquid stream is 
apportioned and returned to the plural walled compartments 110A so as to 
maintain a preselected liquid-phase temperature differential along the 
reactor between the terminal and initial compartments. Transfer pump 138A 
in fluid communication with the reactor 100A, thus assists in passing the 
refluxed reaction mixture to the final or terminal one of the plural 
walled compartments 110A. 
The alkylaromatic feed stream, mentioned above, may include solvent or 
solvents and catalyst, or the solvent and the catalyst may be added 
separately to each compartment. 
Suitable solvents for use in the method of this invention include any 
C.sub.2 -C.sub.6 fatty acids such as acetic acid, propionic acid, 
n-butyric acid, isobutyric acid, n-valeric acid, trimethylacetic acid and 
caproic acid and water and mixtures thereof. Preferably, the solvent is a 
mixture of acetic acid and water, which more preferably contains from 0.5 
to 20 weight percent of water, as introduced into the oxidation reactor. 
Suitable catalysts for use include any catalyst system conventionally used 
for liquid phase oxidation of alkyl aromatics and preferably include the 
catalyst metals comprising cobalt and manganese, and a bromine source. 
Zirconium compounds may also be added as catalyst metals. 
Such bromine sources include elemental bromine (Br.sub.2), or ionic bromide 
(for example, HBr, NaBr, KBr, NH.sub.3 Br, etc.), or organic bromides 
which are known to provide bromide ions at the operating temperature of 
the oxidation (e.g., bromobenzenes, benzylbromide, mono-and di-bromoacetic 
acid, bromoacetyl bromide, tetrabromoethane, ethylenedibromide, etc.). 
As mentioned above, a portion or all of the alkylaromatic feed stream is 
introduced into the initial compartment of the reactor 100A via feed pipe 
114A. If only a portion of the alkylaromatic is added to the first 
compartment, the remainder of the alkylaromatic feed stream can be 
selectively fed to any one of the series-arranged compartments 110A via 
the inlet pipes 144A. 
The feed pipe 114A can include an automatic flow control valve 146A 
operatively connected to a second process control system 148A, and each of 
the inlet pipes 144A can include an automatic flow control valve 150A 
operatively connected to the second process control system 148A, for 
passing at least a portion of the reaction mixture comprising the 
alkylaromatic through the plurality of series-arranged walled compartments 
110A and for selectively introducing the remainder portion of the reaction 
mixture into another one, or into other ones, of the series-arranged 
walled plural compartments. 
A preferred embodiment of the instant invention also comprises a method and 
an apparatus for apportioning the solvent or solvents, catalyst components 
and oxidizing gas independently to each of the series-arranged walled 
compartments. The reaction mixture composition of each reactor compartment 
may, therefore, be individually adjusted to achieve optimal results for 
oxidation of the selected alkylaromatic. 
Thus the reactor means of this invention further comprises inlet means for 
separately introducing, in preselected amounts, reaction solvent, 
alkylaromatic, catalyst metals, and bromine source to each of the 
series-arranged walled but communicating compartments. This is 
accomplished, for example, by employing additional feed source means 
separately for the catalyst metals, bromine source and solvent. The 
catalyst metals, bromine source and solvent may be separately added to the 
reactor and individual reactor compartments utilizing a method or means 
equivalent to that depicted in FIG. 1 for the addition of the 
alkylaromatic, including additional feed pipes to each compartment, 
additional automatic flow control valves and process control systems for 
apportioning the solvent, catalyst metals or bromine source separately and 
in preselected amounts to each reactor compartment. 
For example, in one preferred embodiment all or essentially all of the 
reaction solvent, catalyst metals and bromine source are added to the 
initial reactor compartment and it is the alkylaromatic that is 
apportioned to some or all of the reactor compartments. In still another 
preferred embodiment the catalyst components, which may comprise a cobalt 
compound, a manganese compound, a zirconium compound and a bromine source, 
are separately apportioned in preselected amounts among the reactor 
compartments. Thus it may be desirable to separately add some of the 
bromine source, or some of the manganese or cobalt or zirconium compounds 
to one or more downstream compartments to achieve optimal results for a 
given alkylaromatic feedstock. The compartmented design of the reactor of 
this invention with the flexibility to separately add alkylaromatic, 
solvent, catalyst metals and bromine source is an advantage of the instant 
process and apparatus and it permits the adjustment of the reaction 
composition within each compartment to be optimized for a specific 
alkylaromatic. It follows that there are many variations of reactor 
compartment composition that can be achieved by utilizing the reactor 
design and method of this invention. This versatility is a substantial 
improvement over prior processes and reactors. 
In FIG. 2, a novel reactor configuration that has proven itself 
particularly useful for certain gas-liquid or gas-liquid-solid reactions, 
especially hydrocarbon oxidation reactions, is disclosed. 
As mentioned above, a continuous stirred tank reactor generally offers 
distinct advantages over a batch-type reactor when effecting hydrocarbon 
oxidation reactions. Initial cost of a continuous stirred tank reactor is 
typically lower than that for a batch-type reactor of comparable 
throughput. Continuous stirred tank reactor volume is generally utilized 
more efficiently than is the reactor volume of a batch-type reactor. 
Variation of product quality, which is typically undesirable, occurs much 
more frequently in batch-type reactors than in a continuous stirred tank 
reactor. Also, operational expenses are generally higher for a batch-type 
reactor than for a continuous stirred tank reactor at a given throughput 
rate. 
However, in the oxidation of certain alkylaromatics such as the oxidation 
of pseudocumene to trimellitic anhydride (or acid), where certain 
product-inhibited reactions are experienced, batch reactors are generally 
utilized. In the oxidation of pseudocumene to trimellitic acid using heavy 
metal catalysts, the trimellitic acid product complexes to and thereby 
deactivates the metal catalysts and inhibits the reaction. The reason that 
batch reactors are used to effect such hydrocarbon oxidation reactions is 
that the desired oxidation reaction, at least up to the advent of our 
invention, has generally not been able to be carried out to the desired 
degree of conversion in a single continuous stirred tank reactor (CSTR) 
because of the presence of the reaction-inhibiting product. The use of a 
pair of CSTR's has been similarly unavailing. That is, the concentration 
of the reaction-inhibiting product, which can be either a desired product 
or an undesired by-product of the desired hydrocarbon oxidation reaction, 
is typically so great as to inhibit production of the desired product. 
With the present invention, on the other hand, the desired degree of 
conversion of product-inhibited oxidation reactions can be readily 
effected in continuous stirred tank reactors. It has been demonstrated, 
for example, that pseudocumene can be continuously oxidized to desirable, 
high yields while utilizing continuous stirred tank reactors 100B and 100C 
compartmentalized by apertured, substantially horizontally-arranged 
dividers 108B and 108C. The oxidation reactors 100B and 100C are connected 
to one another in series. Each reactor 100B and 100C is provided with 
apertured, horizontal dividers 108B and 108C to provide three walled 
compartments 110B and 110C for each of the reactors 100B and 100C. 
Each reactor 100B and 100C is also provided with a centrally located and 
vertical disposed agitator shaft, 102B and 102C, that is driven by a 
respective, commerically available drive mechanism 106B and 106C, 
respectively. Impeller blades 104B and 104C, affixed to each of the 
agitator shafts 102B and 102C, respectively, are arranged along the length 
of the agitator shaft 102B and 102C and disposed generally radially 
outwardly therefrom, so as to provide means for circulating the contents 
of each of the walled compartments relative to the inner surface of the 
reactor 100B and 100C. In the reactor 100B and 100C, a gap or aperture is 
provided between the agitator shaft 102B and 102C, and each of the 
horizontal dividers 108B and 108C surrounding such agitator shaft 102B and 
102C. Such a gap or aperture allows upward flow of contained liquid and 
gas (or vapor) within each reactor. FIG. 2 thus schematically depicts a 
six-stage, sequentially-arranged, compartmented continuous stirred tank 
reactor apparatus for effecting desired hydrocarbon oxidation. 
In the oxidation of pseudocumene to produce trimellitic anhydride (or acid) 
it is desirable that that the liquid feed comprise pseudocumene as 
alkylaromatic hydrocarbon reactant, water and acetic acid as solvent, and 
a catalyst. In the oxidation of pseudocumene to trimellitic acid, such a 
catalyst typically comprises cobalt, manganese, zirconium and bromine. 
In one preferred method of operation, all of such liquid feed from the 
alkylaromatic feed source 112 is introduced into the bottom compartment of 
the first reactor 100B via feed pipe 114B. An oxygen-containing gas from 
the oxygen-containing gas source 116 is concurrently introduced into the 
bottom of both reactors 100B and 100C via respective feed pipes 118B and 
118C. In each reactor 100B and 100C vapor is separated from liquid. The 
thus-separated vapor is next passed from the uppermost or downstream-most 
compartment via respective discharge pipes 122B and 122C to a respective 
condenser 120B and 120C, respectively. 
In the first reactor 100B, the partially-reacted liquid reaction mixture, 
comprising the trimellitic anhydride product, is discharged from the 
uppermost compartment via discharge pipe 152, and is passed via transfer 
pump 154 and via inlet pipe 156 into the bottom compartment of the second 
reactor 100C. 
Because this oxidation of pseudocumene is an exothermic reaction, heat of 
reaction is controlled by vaporizing a portion of the reaction mixture. 
Off gas and reaction mixture vapors from reactors 100B and 120C are 
respectively passed to condensers 120B and 120C, each of which produce 
non-condensables as well as condensed process vapors. 
The non-condensables exiting the first condenser 120B are separated from 
process condensate by separator 130B which vents a non-condensable off gas 
via vent line 132B, and passes the process condensate to return pump 138B 
which, in turn, returns the process condensate to the first reactor 100B 
as reflux. Similarly, exiting the second condenser 120C, the 
non-condensables are separated from the process condensate by separator 
130C which vents the non-condensable off gas via vent line 132C and passes 
the process condensate to return pump 138C which returns the process 
condensate to the second reactor 100C, as reflux, as will be described in 
greater detail hereinbelow. 
On the one hand, it is desirable to maintain the oxygen concentration in 
the vent gas at, or below, a predetermined concentration for certain 
safety reasons. On the other hand, it is desirable in oxidizing an 
alkylaromatic to maintain a relatively high oxygen concentration 
throughout the reaction mixture for purposes of optimizing product 
conversion and yields. High oxygen partial pressure also reduces the 
formation of undesirable colored by-products by suppressing coupling 
reactions. In a conventional hydrocarbon oxidation reactor, the vent-gas 
oxygen-concentration safety consideration, typically limits the particular 
oxygen concentration level that can be achieved within the conventional 
purposes. In practicing the present invention, however, it has been 
observed that a relatively highest oxygen concentration can be maintained 
in, for example, the lower two compartments of each of the reactors 100B 
and 100C, relative to a conventional reactor, while a particular off-gas 
oxygen concentration is maintained for safety reasons. 
It has also been found that desirable reaction controls can be effected by 
apportioning the returned process condensate among the walled compartments 
of the reactors 100B and 100C. That is, the process condensate is 
controllably split and returned as reflux to the respective walled 
compartments 110B and 100C of each reactor 100B and 100C as dictated by 
process considerations. Such reflux split, for example, can be based upon 
certain desired temperature differences between adjacent compartments. 
Referring to the first reactor 100B in the embodiment of FIG. 2, a first 
three-way valve 158B is operatively connected to the process control 
system 136B and controllably returns process condensate as reflux, either 
to the uppermost compartment or to a second three-way valve 160B. The 
second three-way valve 160B, also operatively connected to the process 
control system 136B, in turn, controllably returns process condensate, as 
reflux, either to the intermediate compartment or to the lowermost 
compartment. Each of the compartments of the reactor 100B is provided with 
a liquid-phase temperature-sensor 134B operatively connected to the 
process-control system 136B. In operation, that fraction of process 
condensate that is refluxed to any one compartment of the reactor 100B is 
controlled to maintain preselected liquid-phase temperature differentials 
between adjacent compartments, thereby also maintaining a desired 
liquid-phase temperature differential between the lowermost and uppermost 
reactor compartments. Accordingly, the thus-condensed liquid stream is 
apportioned among and returned to the plural, walled compartments so as to 
maintain a corresponding plurality of preselected liquid-phase temperature 
differentials having a predetermined magnitude. 
The second reactor 100C is similarly equipped with a transfer pump 138C, 
first and second three-way valves 158C and 160C, temperature-transducers 
134C, and a process-control system 136C, for maintaining a desired 
liquid-phase temperature differential between the lowermost and uppermost 
compartments. 
Yet another aspect or feature of the present invention provides means for 
effecting desired disentrainment within the three-compartment continuous 
stirred tank reactors shown, for example, in FIG. 2. As further 
illustrated in FIG. 3, a three-compartment continuous stirred tank reactor 
100D can be provided with substantially vertical, i.e., upstanding, 
baffles in some of the compartments. The apertures in partitions 108D are 
in substantial registry so as to accommodate impeller shaft 102D 
therethrough. In particular, it has been observed that a desired liquid 
hold-up and aeration can be achieved in each of the compartments 110D, 
110E and 110F by incorporating upstanding baffles 162 in at least one of 
the downstream compartments, usually in the uppermost or downstream-most 
compartment 110D. The absence of baffles from the intermediate compartment 
110E and lowermost or upstream-most compartment 110F enables impeller 
blades 104E and 104F affixed to the agitator shaft 102D, to circulate the 
contents of the intermediate and lowermost compartments 110E and 110F 
relative to the inner surface of the reactor 100D. Such circulation of the 
contents, in turn, generates centrifugal force fields within the 
intermediate and lowermost compartments 110E and 110F. 
Thus, in operation, rotation of the impeller blades 104E and 104F relative 
to the reactor 100D causes the liquid portion of the reactor contents 
(generally, a gas-in-liquid dispersion) to be urged radially outwardly 
from agitator shaft 102D while the gas (or vapor) portions of the reactor 
contents are free to pass upwardly from one reactor compartment to the 
next reactor compartment, thereby effecting disentrainment of the liquid 
portion of the reactor contents from the vapor (or gaseous) portion within 
the intermediate and lowermost compartments 110E and 110F. This is highly 
desirable from a processing standpoint. Such an arrangement, it has been 
observed, tends to concentrate the gaseous (or vapor) portion of the 
reactor contents (a liquid-in-gas-dispersion) located generally near the 
central axis of the reactor compartment to the next because of the 
provided gap or aperture between the agitator shaft 102D and the 
horizontal partitions or dividers 108D. At the same time, and because the 
liquid-in-gas dispersion is passing through the rotating impeller blades 
of the agitator shaft, the excess entrained liquid is effectively knocked 
back by centrifugal fields, impellers 104E in the intermediate compartment 
110E are preferably of the axial-downflow type. Impellers 104F in the 
lower-most compartment 110F are of the so-called "straight-blade" type, 
which generally cause the gas-in-liquid dispersion in the lowermost 
compartment 110F to be urged radially outwardly from agitator shaft 102D. 
A feature of the present invention will now be discussed with reference to 
FIG. 2. In the production of other aromatic carboxylic acids, such as 
liquid-phase oxidation of meta-xylene to produce isophthalic acid and 
liquid-phase oxidation of para-xylene to produce terephthalic acid, it has 
been discovered that it is desirable to add the alkylaromatic feed not 
only to the lowermost compartment via feed pipe 114B but also to the 
intermediate compartment via feed pipe 164B, or to the uppermost 
compartment via feed pipe 166B, or to both such compartments as desired. 
To that end, the several feed pipes 114B, 164B and 166B are each provided 
with a separate automatic flow control valve 168B, 170B and 172B, each of 
which is operatively connected to a flow-control system 148B. 
As mentioned above, the partially-reacted liquid reaction mixture that is 
discharged from the uppermost compartment of the first reactor 100B is 
passed via discharge pipe 152 to transfer pump 154 which, in turn, 
supplies the partially-reacted liquid reaction mixture into the lowermost 
compartment of the second reactor 100C via the inlet line 156. It is 
desirable to add the partially-reacted liquid reaction mixture not only to 
the lowermost compartment via inlet line 156 but also to the intermediate 
compartment via feed pipe 164C or to the uppermost compartment via feed 
pipe 166C, or to both such compartments as desired. To that end, the 
several feed pipes 156C, 164C and 166C are each provided with a separate 
automatic flow control valve 168C, 170C and 172C each of which is 
operatively connected to a separate flow-control system 148C. 
From the terminal compartment of the second reactor 100C, the 
product-containing reaction mixture is passed via discharge pipe 174 to 
the product collection site 126, for further processing, as desired. 
The following examples are provided to illustrate some of the preferred 
embodiments of the present invention but are not intended to limit the 
scope of the present invention. 
EXAMPLE 1 
The continuous oxidation of p-xylene to terephthalic acid was carried out 
in a reactor having a single horizontal divider plate forming two 
compartments. The volume of the lower compartment was about 4.8 liters. 
The liquid level in the upper compartment was adjusted to give an 
effective volume of about 4.8 liters. A liquid feed stream consisting of 
3320 g/hr acetic acid, 1470 g/hr p-xylene, 66.3 g/hr water, 7.25 g/hr 
cobalt acetate tetrahydrate, 18.5 g/hr manganese acetate tetrahydrate and 
8.83 g/hr of a 48 percent HBr solution in water was continuously fed to 
the bottom compartment. Air was continuously fed to the bottom compartment 
at a rate of 180 SCFH. The liquid product was continuously withdrawn from 
the top compartment. The off gas passed through a downflow condenser and 
through a back pressure control valve. The top compartment was maintained 
at the desired temperature by adjusting the reactor pressure. The bottom 
compartment was maintained at the desired temperature by apportioning the 
liquid reflux between the two compartments. Run conditions and 
experimental results are shown in Table 1. 
EXAMPLE 2 
p-Xylene was oxidized as in Example 1 except that only 70 percent of the 
total p-xylene was added to the bottom compartment, with the balance being 
added to the top compartment. Compared to Example 1, the conversion 
increased, as evidenced by the lower contents of 4-carboxybenzaldehyde and 
p-toluic acid in the product slurry. There was no increase in carbon 
oxides production, which is normally seen when the conversion is increased 
by such means as increasing temperature or catalyst concentration. 
EXAMPLE 3 
p-Xylene was oxidized as in Example 1 except that only 50 percent of the 
total bromine was added to the bottom compartment, with the balance being 
added to the top compartment. Compared to Example 1, the production of 
benzoic acid was reduced. Benzoic acid is undesirable because it builds up 
when mother liquor is recycled to the reactor, and it contributes to the 
formation of colored and fluorescent by-products. 
TABLE 1 
______________________________________ 
Oxidation of Para-xylene 
Example 1 2 3 
______________________________________ 
Temperatures, .degree.F.: 
Top Compartment 383 383 385 
Bottom Compartment 338 338 340 
Reactor Pressure, psig 
222 223 222 
Oxygen Concentration in Vent Gas, % 
2.94 2.93 2.96 
Moles CO/(Mole pX) 0.027 0.028 0.027 
Moles CO.sub.2 /(Mole pX) 
0.118 0.120 0.109 
Product Slurry Analyses: 
Optical Density (340 nm) 
12.5 10.4 10.3 
Benzoic Acid, Weight % 
0.096 0.016 0.069 
4-Carboxybenzaldehyde, Weight % 
1.10 0.86 1.08 
p-Toluic Acid, Weight % 
2.58 1.84 2.65 
______________________________________ 
EXAMPLE 4 
The continuous oxidation of m-xylene to isophthalic acid was carried out in 
the same reactor as used in Example 1. A liquid feed stream consisting of 
5710 g/hr acetic acid, 1265 g/hr m-xylene, 530 g/hr water, 5.18 g cobalt 
acetate tetrahydrate, 15.6 g manganese acetate tetrahydrate and 5.16 g of 
a 48 percent solution of HBr in water was continuously fed to the bottom 
compartment. Air was continuously fed to the bottom compartment at a rate 
of 166 SCFH. The liquid product was continuously withdrawn from the top 
compartment. The off gas passed through a downflow condenser and through a 
back pressure control valve. The top compartment was maintained at the 
desired temperature by adjusting the reactor pressure. The bottom 
compartment was maintained at the desired temperature by apportioning the 
liquid reflux between the two compartments. Run conditions and 
experimental results are shown in Table 2. 
EXAMPLE 5 
m-Xylene was oxidized as in Example 4 except that only half of the total 
bromine was added to the bottom compartment, with the balance being fed to 
the top compartment. Compared to Example 4, the optical density, carbon 
oxides and benzoic acid were reduced. Benzoic acid is undesirable because 
it builds up when mother liquor is recycled to the reactor, and it 
contributes to the formation of colored and fluorescent by-products. Lower 
values for optical density, a measure of colored impurities, are 
preferred. 
EXAMPLE 6 
m-Xylene was oxidized as in Example 4 except that only 62 percent of the 
total m-xylene was fed to the bottom compartment, with the balance being 
fed to the top compartment. Compared to Example 4, both the yield of 
carbon oxides and the optical density were reduced. 
TABLE 2 
______________________________________ 
Oxidation of Meta-xylene 
Example 4 5 6 
______________________________________ 
Temperatures, .degree.F.: 
Top Compartment 417 418 417 
Bottom Compartment 368 378 376 
Reactor Pressure, psig 
332 325 350 
Oxygen concentration in Vent Gas, % 
2.20 2.38 2.33 
Yields, Mole % Based on Xylene Feed: 
m-Toluic Acid 6.90 4.30 5.48 
3-Carboxybenzaldehyde 
1.92 1.30 1.61 
Carbon Oxides 1.92 1.75 1.67 
Benzoic Acid 0.67 0.53 0.72 
Slurry Optical Density (340 nm) 
15.0 11.0 10.3 
______________________________________ 
EXAMPLE 7 
The same reactor was used as in Example 1 except that no divider plate was 
used, i.e., there was only a single compartment. Vortex-breaking baffles 
were used. A liquid feed was continuously added at the rate of 2350 g/hr 
pseudocumene, 4230 g/hr acetic acid, 220 g/hr water, 19.0 g/hr cobalt (as 
the acetate tetrahydrate), 5.7 g/hr manganese (as the acetate 
tetrahydrate), 29.9 g/hr bromine (as a 48 percent solution of HBr in 
water) and 0.54 g/hr zirconium (as zirconium acetate). Air was added at 
the rate of 300 SCFH. The liquid product was continuously withdrawn from 
the reactor. The off gas passed through a downflow condenser and a back 
pressure valve. The liquid condensate was returned to the reactor. The 
reactor temperature was controlled by adjusting the pressure. The results 
are shown in Table 3. The production of carbon oxides was very high. 
Moreover, it was found impossible to increase the conversion of 
pseudocumene significantly beyond about 70 percent, even with drastic 
increases in temperature. 
EXAMPLE 8 
The same two-compartment reactor as in Example 1 was used. Vortex-breaking 
baffles were used in both compartments. The liquid and air feed rates were 
the same as in Example 7. The results are shown in Table 3. Although the 
overall pseudocumene conversion was slightly higher than in Example 7, the 
carbon oxides production was much lower. 
EXAMPLE 9 
The same reactor as in Example 1 was used except that two horizontal 
divider plates were installed to give three compartments. Vortex-breaking 
baffles were used in all three compartments. The volume of the bottom 
compartment was approximately 2.4 liters. The volumes of the middle and 
upper compartments were approximately 3.6 liters each. The liquid feed 
rates were the same as in Example 7. The air feed rate was 170 SCFH to the 
lower compartment and 130 SCFH to the middle compartment. The condensate 
reflux rates to the middle and lower compartments were apportioned such as 
to hold the temperatures of these compartments at the desired level. The 
results are shown in Table 3. The carbon oxides production was even lower 
than with the two-compartment reactor of Example 8. 
EXAMPLE 10 
The same reactor as in Example 9 was used except that no vortex-breaking 
baffles were used in the middle and bottom compartments. The upper 
compartment did contain vortex-breaking baffles. Thus the liquid in the 
lower compartments was allowed to rotate. Model experiments in a Plexiglas 
reactor using air and water showed that the centrifugal forces generated 
by the rotation reduce entrainment of liquid into the next compartment, 
improving the liquid hold-up in the compartment. A liquid feed was 
continuously added to the bottom compartment at the rate of 2350 g/hr 
pseudocumene, 4230 g/hr acetic acid, 220 g/hr water, 19.0 g/hr cobalt (as 
the acetate tetrahydrate), 4.1 g/hr manganese (as the acetate 
tetrahydrate), 29.9 g/hr bromine (as a 48 percent solution of HBr in 
water) and 0.54 g/hr zirconium (as zirconium acetate). Air was added to 
the bottom compartment at the rate of 300 SCFH. The liquid product was 
continuously withdrawn from the reactor. The off gas passed through a 
downflow condenser and a back pressure valve. The liquid condensate was 
returned to the reactor. The temperature of the upper-compartments was 
controlled by adjusting the pressure. The temperatures of the middle and 
lower compartments were controlled by adjusting the amount of condensate 
reflux returned to these compartments. The results are shown in Table 3. 
The production of carbon oxides was even lower than in Example 9. 
TABLE 3 
______________________________________ 
Oxidation of Pseudocumene 
Example 7 8 9 10 
______________________________________ 
No. of Compartments 1 2 3 3 
Baffles in Lower Compartments 
Yes Yes Yes No 
Temperature, .degree.F.: 
Top 314 327 354 
Middle 342 307 315 
Bottom 340 276 296 
Pseudocumene Conversion, Percent: 
Overall 62.6 65.6 65.9 66.3 
To Carbon Oxides 4.38 2.91 2.79 
2.63 
______________________________________ 
EXAMPLE 11 
Four reaction stages were obtained by connecting two two-compartment 
reactors as described in Example 1 in series. A positive displacement pump 
was used to transport the effluent from the top compartment of the first 
reactor to the bottom compartment of the second reactor. The pumping rate 
was adjusted to keep the liquid in the top compartment of the first 
reactor at the desired level. All of the liquid feed was added to the 
bottom compartment of the first reactor. The air feed was split between 
the bottom of the first reactor and the bottom of the second reactor. Run 
conditions and experimental results are shown in Table 4. The low content 
of methyl di-acids (4-methyl isophthalic acid, 2-methyl terephthalic acid 
and 4-methyl orthophthalic acid, which are reaction intermediates) shows 
that the reaction went essentially to completion. The product analyses are 
similar to those obtained from batch oxidations. 
EXAMPLE 12 
Pseudocumene was oxidized in a manner similar to that of Example 11 except 
that additional (tailout) catalyst was added. The tailout catalyst metals 
were equally split between the second, third and fourth compartment. The 
tailout bromine was evenly split between the second and third 
compartments. Run conditions and experimental results are shown in Table 
4. The methyl di-acids were even lower than in Example 11. 
EXAMPLE 13 
Six reaction stages were obtained by connecting two three-stage reactors as 
described in Example 10 in series. A positive displacement pump was used 
to transport the effluent from the top compartment of the first reactor to 
the bottom compartment of the second reactor. The pumping rate was 
adjusted to keep the liquid in the top compartment of the first reactor at 
the desired level. Run conditions and experimental results are shown in 
Table 4. Note that the pseudocumene feed was split evenly between the 
first two compartments. The oxidation intermediates (methyl di-acids) in 
the product cake were very low. 
TABLE 4 
______________________________________ 
Oxidation of Pseudocumene in Four and Six-Stage Reactors 
Example 11 12 13 
______________________________________ 
Feed Rates, g/hr: 
Pseudocumene to First Compartment 
5440 5440 2720 
Pseudocumene to Second Compartment 
0 0 2720 
Acetic Acid 9770 9770 9770 
Water 515 515 515 
Initial Catalyst, Weight % Based on 
Pseudocumene: 
Co 0.28 0.28 0.28 
Mn 0.060 0.088 0.088 
Br 0.44 0.28 0.263 
Zr 0.008 0.010 0.010 
Tailout Catalyst, Weight % Based on 
Pseudocumene: 
Mn 0 0.011 0.023 
Br 0 0.13 0.263 
Zr 0 0.006 0.011 
First Reactor: 
Residence Time, min 38.5 46.4 48.0 
Air Feed Rate, SCFH 267 264 250 
Bottom Temperature, .degree.F. 
316 316 305 
Middle Temperature, .degree.F. 338 
Top Temperature, .degree.F. 
336 336 368 
Pseudocumene Conversion, Percent 
71.5 70.4 72.0 
Carbon Oxides, Mole % of 
4.16 4.26 5.29 
Pseudocumene Fed 
Second Reactor: 
Residence Time, min 46.2 56.1 45.7 
Air Feed Rate, SCFH 124 127 NA 
Bottom Temperature, .degree.F. 
371 365 362 
Top Temperature, .degree.F. 
385 381 384 
Pseudocumene Conversion, Percent 
26.6 26.8 392 
Carbon Oxides, Mole % of 
4.26 8.60 4.51 
Pseudocumene Fed 
Dry Cake Analysis, Weight %: 
Methyl Di-Acids 0.644 0.138 0.093 
Di-Acids 2.176 1.952 1.808 
Trimellitic Acid 91.59 93.66 91.16 
Normalized Trimellitic Acid Yield 
86.3 87.4 85.8 
______________________________________ 
These examples demonstrate the advantages that are obtained by using the 
method and apparatus of the present invention for the oxidation of the 
alkylaromatic to the corresponding carboxylic or polycarboxylic acids. 
What has been illustrated and described herein is a method and an apparatus 
for continuously oxidizing an aromatic alkyl in the liquid phase and under 
oxidation-reaction conditions to produce an aromatic carboxylic acid 
product. While the method and apparatus of the present invention have been 
illustrated and described with reference to several preferred embodiments, 
the present invention is not limited thereto. Alternatives, changes and 
modifications are possible and will become apparent to those skilled in 
the art upon reference to the foregoing description and the drawings. 
Accordingly, such alternatives, changes and modifications form a part of 
the invention insofar as they fall within the spirit and scope of the 
appended claims.