Cobalt-catalyzed oxidation of hydrocarbons

A process for the liquid phase oxidation of hydrocarbons with a molecular oxygen-containing gas in the presence of a dissolved cobalt salt catalyst characterized in that the oxidation is carried out in the substantial absence of chromium in the reaction medium i.e. a concentration of chromium in the liquid phase of not greater than 400 ppm.

The present invention relates to the cobalt-catalysed liquid phase 
oxidation of hydrocarbons to produce oxygen-containing organic compounds. 
Processes for the liquid phase oxidation of hydrocarbons, especially 
saturated aliphatic hydrocarbons containing from 3 to 7 carbon atoms, in 
the presence of a soluble cobalt salt as catalyst to produce oxygenated 
reaction products including a substantial proportion of carboxylic acids 
are well known. However, scant attention appears to have seen paid to such 
oxidation processes wherein the recovered catalyst solution is recycled 
directly to the oxidation reaction. In the past it has been the practice 
to recover the cobalt catalyst from the unwanted oxidation by-products and 
reactor corrosion products by precipitation as an insoluble cobalt 
compound, followed by conversion to the soluble cobalt catalyst species. 
Representative of the published patent specifications describing this 
technique is U.S. Pat. No. 3,840,469 wherein cobalt (II) ions initially 
present in a first liquid medium consisting essentially of acetic acid 
derived from the liquid phase oxidation of an aliphatic hydrocarbon are 
recovered and a fresh solution of the same produced by: 
(A) INCORPORATING INTO SAID FIRST ACETIC ACID MEDIUM CONTAINING COBALT (II) 
ions and about 3 to 15 percent water by weight a source of oxalate ions in 
a quantity at least sufficient to react stoichiometrically with said 
cobalt (II) ions to form cobalt oxalate. 
(B) FORMING A SOLID PRECIPITATE OF COBALT (II) oxalate in said first acetic 
acid medium while at a temperature of about 20 to 150.degree. C., (c) 
separating said precipitate of cobalt oxalate from said first acetic acid 
medium, 
(D) CONTACTING SAID SEATED PRECIPITATE OF COBALT OXALATE WITH A SECOND 
MEDIUM CONSISTING ESSENTIALLY OF ACETIC ACID CONTAINING ABOUT 3 TO 15 
PERCENT WATER BY WEIGHT PROVIDED AT AN ELEVATED TEMPERATURE OF ABOUT 60 TO 
140.degree. C. and containing calcium ions dissolved therein in a quantity 
at least equivalent stoichiometrically to the cobalt (II) ions contained 
in said cobalt oxalate to form calcium oxalate, 
(E) FORMING A SOLID PRECIPITATE OF CALCIUM OXALATE IN SAID SECOND ACETIC 
ACID MEDIUM WHILE SAID COBALT (II) ions are solubilised, and 
(F) SEATING SAID SOLID PRECIPITATE OF CALCIUM OXALATE FROM SAID SECOND 
ACETIC ACID MEDIUM TO PRODUCE A FRESH SOLUTION OF COBALT (II) ions 
dissolved in acetic acid. 
Such a process is typical of one approach to the problem of catalyst 
recycle. It is involved and economically unattractive. Another approach, 
apparently not attempted in the past, is to examine the oxidation 
by-products and corrosion products with a view to eliminating or reducing 
to negligible proportions the formation of those components which are 
harmful to the oxidation reaction. 
Adopting this approach we have found that chromium has a marked inhibitory 
effect on the oxidation reaction rate at concentrations greater than 400 
ppm. Since chromium is a component of stainless steel, the material 
conventionally used for the fabrication of reaction vessels, corrosion 
inevitably leads to the presence of chromium in the liquid reaction medium 
in contact with stainless steel surfaces; the longer the duration of 
contact the greater being the concentration of chromium in the reaction 
medium. Consequently, we have found it impossible to maintain continuous 
oxidation in a stainless steel reactor with recycle of the catalyst 
solution over extended periods. 
Thus the present invention provides a process for the production of 
oxygen-containing organic compounds by the oxidation of a hydrocarbon 
feedstock with a molecular oxygen-containing gas in the presence of from 
0.1 to 10 percent by weight, based on the weight of reaction medium, of a 
soluble cobalt compound other than the halide as the sole catalyst, and an 
essentially inert reaction medium at a temperature in the range 70 to 
150.degree. C. and at a pressure sufficient to maintain the reactants in 
the liquid phase characterised in that the oxidation is carried out in a 
reaction vessel either fabricated in or lined, on at least that part of 
the internal surface thereof in contact with the liquid reaction medium, 
with a substantially chromium-free material when the cobalt catalyst is 
separated from the oxygenated products as a solution in the inert reaction 
medium and returned to the oxidation. 
By substantially chromium-free material within the context of the present 
application is meant a material which on prolonged contact with liquid 
reaction medium does not lead to a concentration of chromium in that 
reaction medium of more than 400 ppm. 
Substantially chromium-free materials which may be used include glass, 
resins, porcelain, enamel, titanium, tin and high nickel content alloys 
such as Hastelloy B. The preferred material is titanium. 
The hydrocarbon to be oxidised may be a saturated aliphatic hydrocarbon, a 
cycloaliphatic or alkyl aromatic hydrocarbon. It is preferred to employ a 
saturated aliphatic hydrocarbon containing from 3 to 7 carbon atoms. An 
especially preferred hydrocarbon is butane. The butane is preferably used 
in the form of n-butane but may contain isobutane and minor amounts of 
other saturated hydrocarbons. Although isobutane may be present in greater 
proportions, it is preferred that the n-butane contain not more than 40% 
w/w isobutane. 
The molecular oxygen-containing gas may be substantially pure oxygen or may 
be any gas mixture containing molecular oxygen. It is preferred to employ 
air, because the likelihood of forming explosive mixtures in the unreacted 
gas withdrawal system can be prevented far more easily than when using 
substantially pure oxygen or gaseous mixtures containing higher 
concentrations of molecular oxygen than air. Alternatively substantially 
pure oxygen may be introduced into the reaction mixture together with 
recycled (recirculated) off-gases from the oxidation process as a means of 
reducing the likelihood of forming explosive mixtures in the unreacted 
gases. 
The oxidation is carried out in a liquid phase comprising the hydrocarbon 
to be oxidised and, in addition, an essentially inert liquid reaction 
medium. The latter may be any essentially inert liquid in which the cobalt 
compound catalyst is soluble. It is preferred that the inert reaction 
medium comprises a lower acid having from 2 to 4 carbon atoms such as 
acetic acid, propionic acid or normal butyric acids. When butane is the 
saturated aliphatic hydrocarbon, it is preferred to employ acetic acid 
since it is the desired product and separation procedures are simplified. 
Other inert reaction media such as benzene, chlorobenzene, phenyl benzoate 
or benzoic acid may be used. The amount of the inert reaction medium 
employed is not critical provided that sufficient is used to provide a 
substantially homogeneous reaction medium throughout the course of the 
oxidation reaction. The weight ratio of inert reaction medium to 
hydrocarbon may be in the range 1:10 to 100:1. 
The cobalt catalyst may be any cobalt salt soluble in the reaction medium. 
Suitable examples include cobalt sulphate, nitrate, acetate, propionate, 
butyrate, isovalerate, benzoate, toluate, naphthenate, salicylate, 
phthalocyanine or acetyl acetonate. The cobalt catalyst may be introduced 
as a cobalt salt in which the cobalt is in the +2 or +3 oxidation states 
or partially in the +2 and +3 oxidation states. Preferred salts are 
cobaltous and cobaltic acetates or a mixture thereof. Even more preferably 
not less than 5 and not more than 90 percent of the cobalt catalyst is 
introduced into the reaction vessel in the +3 oxidation state in a 
continuous process and not less than 11 and not more than 90 percent of 
the cobalt catalyst is introduced into the reaction vessel in the +3 
oxidation state in a batch process. Using these particular catalysts it is 
not necessary to employ a co-oxidant because the long induction period 
normally experienced using a cobalt catalyst exclusively in the +2 
oxidation state is substantially eliminated. However, if desired a 
co-oxidant such as methyl ethyl ketone may be employed. The amount of the 
cobalt catalyst employed, calculated as cobalt, may be in the range 0.1 to 
10% by weight and is preferably in the range from 0.2 to 5% by weight. 
The reaction temperature is preferably in the range 90 to 140.degree. C. 
and a reaction pressure in the range 10 to 100 bar absolute is usually 
found to maintain the reactants in the liquid phase. Reaction time is not 
critical, being dependent merely upon the extent of conversion required. 
Thus the reaction period may be in the range of from one minute to 20 
hours, preferably from 10 minutes to 3 hours. 
The process may be carried out in any suitable manner, either batchwise 
wherein the catalyst solution separated from the oxidation products is 
re-used in a further batch operation or continuously wherein the separated 
catalyst solution is continuously recycled to the oxidation, the latter 
method of operation being preferred. 
The oxygenated products resulting from the oxidation of saturated aliphatic 
hydrocarbons comprise carboxylic acids in substantial proportions and 
minor proportions of ketones, esters and oxides of carbon, e.g. carbon 
monoxide and carbon dioxide. Thus the oxidation of butane results in a 
product comprising predominantly acetic acid, together with minor amounts 
of propionic acid and butyric acid, methyl ethyl ketone, sec-butyl 
acetate, ethyl acetate, methyl acetate, acetone, succinic acid, carbon 
monoxide, carbon dioxide and higher boiling products. Compared with other 
hydrocarbon oxidation routes to acetic acid, in the oxidation of butane by 
the process of the present invention the quantity of carbon monoxide, 
being less than 2% of the acetic acid made, is exceptionally low. 
When the process is operated batchwise, the hydrocarbon, the inert reaction 
medium and cobalt catalyst, for example n-butane, acetic acid and 
Co(+2)/Co(+3) catalyst such as cobaltous/cobaltic acetate respectively may 
be placed in a closed reactor which is pressurised to the desired reaction 
pressure with a molecular oxygen containing gas, e.g. air. The temperature 
of the mixture may then be raised to the desired reaction temperature 
accompanied by stirring. Since oxygen is consumed in the reaction, 
additional molecular oxygen containing gas, e.g. air, may be introduced 
into the reactor. The reaction may be discontinued at any time but 
preferably when no further oxygen absorption occurs. The reaction mixture 
may be brought to atmospheric pressure, withdrawn from the reaction zone 
and separated into its components, the cobalt catalyst solution being 
recovered as a solution in the inert reaction medium. 
Alternatively hydrocarbon, for example n-butane, and molecular oxygen 
containing gas, e.g. air may be fed continuously to such a reactor 
containing a cobalt catalyst e.g. Co(+2)/Co(+3) and inert reaction medium, 
e.g. acetic acid and the reaction products removed from the reactor, 
either partially or entirely, by continuously withdrawing substantially 
liquid-free gases from the top of the reactor, partially cooling said 
gases thereby providing a condensate which comprises a hydrocarbon-rich 
phase and an aqueous phase rich in acetic acid, separating the phases and 
thereafter recycling the hydrocarbon-rich phase to the reactor and 
separating the aqueous phase into its components. In this manner, as an 
additional benefit water is continuously withdrawn from the reaction 
because the ratio of water to acetic acid in the condensate from the 
reactor off-gas is higher than the corresponding ratio in the reaction 
mixture by virtue of the fact that the relative volatility of water to 
acetic acid is greater than unity. The withdrawal of the aqueous acetic 
acid-rich phase of the overhead condensate therefore leads to a lower 
standing concentration of water in the reaction mixture and results in a 
higher reaction rate and a reduced tendency for the reaction mixture to 
partition into two phases. 
It is further preferred to treat the condensate resulting from partially 
cooling the substantially liquid-free gases withdrawn from the top of the 
reactor with a metal salt having a high solubility in water and acetic 
acid. A preferred metal salt is an alkali metal salt of a carboxylic acid. 
A particularly preferred metal salt is potassium acetate. It is preferred 
to treat the condensate with a concentrated solution of the metal salt in 
a suitable solvent. Suitable solvents are acetic acid, water or mixtures 
thereof. The salt dissolves preferentially in the aqueous phase rich in 
acetic acid and thereby aids phase separation and reduces the quantities 
of water and acetic acid in the hydrocarbon-rich phase, both of which are 
desirable objectives. The products may be removed from the concentrated 
metal salt solution in a single distillation or flash evaporation step and 
the salt solution recycled to the condensate treatment. 
Alternatively, or in addition substantially gas-free liquid may be 
withdrawn from the base of the reactor and separated in a distillation 
column into a distillate fraction containing the bulk of the oxygenated 
reaction products and a base fraction comprising the inert reaction medium 
with the cobalt catalyst dissolved therein, which fraction is recycled to 
the oxidation. 
The separated fraction comprising inert reaction medium with the cobalt 
catalyst dissolved therein may contain, in addition, minor amounts of 
higher boiling compounds, otherwise known as "heavy ends". Because of 
their higher boiling point the "heavy ends" may tend to build up in the 
catalyst solution recycle stream. For this reason it may be desirable to 
remove the "heavy ends" by cooling all or part of the catalyst solution 
recycle stream and subjecting it to settling and/or filtration prior to 
recycle to the oxidation. 
When using cobalt partially in the +3 oxidation state as catalyst it is 
preferred to separate the catalyst solution in inert reaction medium from 
the oxygenated products by the process involving a low-residence time 
distillation column as described in co-pending application Ser. No. 
706,197 filed July 16, 1976, now U.S. Pat. No. 4,086,267, and Ser. No. 
755,148 filed Dec. 29, 1976. 
Whether the reaction products be removed from the reactor by continuously 
withdrawing substantially liquid-free gases from the top of the reactor or 
withdrawing a substantially gas-free liquid from the base of the reactor 
or by a combination of both methods the crude catalyst-free product 
collected in the case of oxidation of butane comprises acetic acid, water, 
minor amounts of other carboxylic acids e.g. propionic acid and butyric 
acids, methyl ethyl ketone, sec-butyl acetate and trace amounts of 
alcohols and other volatile ketones and esters, the actual composition 
depending very much on the composition of the feedstock. 
The crude product may be separated by feeding the product to a distillation 
column wherein a fraction comprising water, methyl ethyl ketone, sec-butyl 
acetate and minor amounts of alcohols and other volatile ketones and 
esters, which products, not including water, are collectively hereinafter 
referred to as "light ends", are taken off as an overhead fraction and 
condensed thereby forming a water-rich phase and a ketone/ester-rich 
phase, said water-rich phase being separated and returned wholly or 
partially to the column at a point near the top thereof and the 
ketone/ester rich phase being removed and passing the base product 
comprising acetic acid, water and carboxylic acid impurities to a second 
distillation column wherein water is removed overhead as an azeotrope 
leaving a base product comprising substantially anhydrous carboxylic 
acids. By passing the base product to a third distillation column acetic 
acid may be separated from the other higher-boiling carboxylic acids as a 
substantially pure product. The ketone/ester-rich phase removed from the 
first column may be further separated into substantially pure methyl ethyl 
ketone and sec-butyl acetate products is so-desired or may be recycled to 
the reactor, though their presence in the reactor is not necessary to 
initiate the oxidation reaction. 
Alternatively the condensate from the first distillation column may be 
separated into a water-rich phase and a ketone/ester-rich phase, the 
water-rich phase being withdrawn and the ketone/ester-rich phase being 
returned at least in part to the column and a base product comprising 
substantially anhydrous acetic acid and minor amounts of other 
higher-boiling carboxylic acids passed to a second distillation column 
wherein substantially pure acetic acid is separated from higher-boiling 
carboxylic acids. Any ketone/ester-rich phase removed from the first 
distillation column may be separated to isolate methyl ethyl ketone and 
sec-butyl acetate or may be returned directly to the reactor.

With reference to FIG. 1 the numeral 1 denotes a reaction vessel fabricated 
in titanium of approximately 5 cm. internal diameter and 1 m in height 
having an external pipe loop 1a between the mid-point and the base; 2 is a 
gas inlet pipe; 3 is a reactor pressure-controlled gas release valve; 3a 
is a reactor off-gas line; 4 is a condensate return pipe; 5 is a butane 
inlet pipe; 6 is the liquid reaction mixture discharge line connecting to 
the distillation zone of FIG. 2; 7 is a catalyst input line; 7a is a 
catalyst solution recycle line; 8 is a water-cooled condenser and 9 is a 
refrigerated brine-cooled condenser. 
With reference to FIG. 2, 10 denotes a vessel for receipt of reaction 
mixture from the reactor through line 6 of FIG. 1; 11 is the distillation 
column feed vessel; 12 is a distillation column feed pump; 13 is a 
thermally insulated glass column of approximately 2.5 cm inner diameter 
and 80 cm. in height packed with Raschig rings (6 mm) connecting through a 
cone and socket joint (not shown) with a falling film evaporator 14, 
serving as a reboiler, consisting of a glass tube of approximately 2.5 cm 
inner diameter and 40 cm. in height. A vapour jacket 15 surrounds the 
evaporator/reboiler 14, the vapor jacket being fed with propionic acid 
vapour from a reboiler 16 which is condensed in the condenser 18; a pump 
19 controls the withdrawal of the catalyst-containing inert reaction 
mixture base product from evaporator/reboiler 14 which in turn controls 
the level of the solution within the reboiler and hence to some extent, 
the rate of vapour generation; 20 is a weir built into the column 13 below 
the level of a condenser 21; 22 is an overflow vessel from which a line 23 
in the base feeds a column reflux pump 24; 25 is a reflux preheater 
(2-pentanol B.Pt. 119.degree. C.); 26 is a pot for receipt of distillate 
from the overflow vessel 22; 27 is a catalyst feed pot and 28 a catalyst 
recycle pump. 
With reference to FIG. 3, 29 is a reaction vessel fabricated in 
unstabilised 316-type (FMB) stainless steel having an internal diameter of 
approximately 100 mm and a height of approximately 1.5m; 30 is a heating 
oil jacket; 31 is a `draught tube` of segmental cross-section, its purpose 
being to promote circulation of the reactor contents; 32 is an air 
distributor; 33 is a level-controlled liquid release valve; 35 is a 
water-cooled condenser; 35 is a refrigerated brine-cooled condenser; 36 is 
an air inlet pipe; 37 is a butane inlet pipe; 38 is a condensate return 
pipe and 39 is a catalyst inlet pipe also functioning as a "light ends" 
return pipe when operating under "light ends" recycle conditions. 
EXAMPLE 1 
Preparation of Catalyst Solution 
Approximately 5l of acetic acid was charged to a glass vessel equipped with 
a gas inlet pipe and stirrer. Cobaltous acetate was dissolved in the 
acetic acid to provide a solution containing about 1% cobalt by weight. 
Oxygen at a rate of about 7l/h was passed via an OZONO (Registered Trade 
Mark) air conditioning unit into the stirred vessel at ambient temperature 
for at least 6h, thereby converting at least 80% of the cobalt acetate in 
the solution into the Co(+3) oxidation state. Ozone in the effluent gas 
was destroyed by passage through vessels containing an aqueous solution of 
potassium iodide. 
Start-up for Continuous Operation 
With reference to FIG. 1 approximately 1.3l of catalyst solution 
(containing about 1% w/w cobalt of which ca. 80% was in the +3 oxidation 
state) was introduced into the reaction vessel via line 7 and the vessel 
was resealed. A stream of nitrogen was introduced slowly via pipe 2 and 
the pressure within the reaction vessel was allowed to increase to ca. 35 
bar absolute by means of the control system which regulated the gas 
release valve 3. About 250g butane was pumped rapidly into the reactor via 
pipe 5. About 350l/h, as measured at S.T.P., of air was then introduced 
via pipe 2 and the reactor was heated rapidly to about 110.degree. C. by 
means of external electrical heating elements. Butane was introduced at 
about 20g/h to compensate for uncondensed butane lost in the off-gas. 
By monitoring the oxygen content of the off-gases it was evident that 
significant oxygen absorption commenced with 15 minutes of the attainment 
of reaction temperature. The butane feed-rate was then increased to about 
100g/h and the withdrawal of liquid reaction mixture, containing reaction 
products, was commenced at this time. Within 1 hour the oxygen content of 
the off-gases has become reasonably stable at 8 to 9% by volume. 
Recovery of catalyst solution from its mixture with the organic compounds 
resulting from the oxidation described above 
With reference to FIG. 2 liquid reaction mixture was continuously withdrawn 
from the reactor via pipe 6 into the reactor product pot 10 held at or 
near atmospheric pressure, from which the greater part of the butane 
present in the mixture was removed as gas. The remaining liquid product 
(approximately 300g/h) was fed through the feed vessel 11 and the feed 
pump 12 to the top of column 13 where it was heated by vapours passing up 
the column and thereby itself separated into a vapour and liquid, the 
liquid passing quickly down the column into the falling-film evaporator 14 
wherein both separation and passage downwards of the liquid portion were 
accelerated. Withdrawal of the catalyst-rich base product was controlled 
by the pump 19. The distillate passing up the column was condensed by the 
condenser 21, the condensate falling into the weir 20, from where it 
flowed into the overflow vessel 22, the overflow passing into the head pot 
26 and the remainder being recycled through the line 23, the pump 24 and 
the preheater 25 to a point in column 13 approximately 10 cm. below the 
feedpoint, thereby providing the required level of reflux with an 
additional heat load on the reboiler 16. 
The recovered cobalt catalyst solution in acetic acid contained 
approximately 0.7% by weight of cobalt of which about 40% was present in 
the +3 oxidation state. The solution was returned to the reactor via the 
catalyst feed pot 27, the recycle pump 28 and the line 7a. The residence 
time of the catalyst solution with the distillation zone was approximately 
2 minutes. 
The results presented in Table 1 indicate that the desired rate of removal 
of acetic acid, "light ends" and a considerable proportion of the water 
from the catalyst solution was effected. At the same time, a significant 
proportion of the cobalt was maintained in the +3 oxidation state. 
TABLE 1 
______________________________________ 
Short Residence Time Distillation for Separation of 
Oxidation Products from Catalyst Solution 
Feed to Base 
Distill. Zone 
Product Distillate 
(Reaction 
(Catalyst 
(Net Reaction 
Mixture) Solution) 
Products) 
______________________________________ 
Feedrate (g/h) 
300 -- 
Take-off rate (g/h) 
-- 210 90 
Compositions (% w/w) 
Water 8.2 2.6 21 
Acetic Acid 79.0 83.4 69 
Propionic Acid 
2.0 2.6 0.8 
Butyric Acid 3.7 4.5 0.5 
"Light Ends"* 
2.6 trace 8.7 
"Involatils"** 
4.4 6.6 -- 
Cobalt concentration 
(% w/w) 0.47 0.7 -- 
Proportion as Co(+3) 
(%) ca. 60 ca. 40 -- 
______________________________________ 
*Composition of the "light ends" is given in Table 2. 
**Involatiles comprise compounds of cobalt, together with "heavy ends" 
which consist in the main of succinic acid. 
Recovery of Oxidation Products with no recycle of "light ends" 
The distillate from the column heads pot 26 was passed to a distillation 
column (not shown) of 2.5 cm inner diameter and 140 cm in height, packed 
with Raschig rings (6 mm) and provided with a conventional reboiler. The 
distillate mixture was introduced to the column near its mid-point. The 
fraction taken overhead from the column separated into two phases, and the 
whole of the lower (aqueous phase) was returned to the top of the column 
while the upper (ketone/ester) phase was withdrawn. The base product was 
withdrawn from the reboiler under liquid level control. This base product 
comprised all the carboxylic acids and the greater part of the water, with 
no ketones or esters detectable by gas-liquid chromatography. The 
ketone/ester phase withdrawn comprised the "light ends" of composition 
given in Table 2, and additionally contained about 6% water in solution. 
The results obtained over a period of 48h continuous operation are given in 
Table 2. Continuous oxidation was maintained for a period of at least 240 
hours without any significant reduction in the oxygen absorption rate. 
TABLE 2 
______________________________________ 
Example 2 Example 3 
"light-ends" 
"light-ends" 
recycle Withdrawal 
______________________________________ 
Reaction Temperature (.degree. C) 
110 110 
Reaction Pressure (bar) 
35 35 
Butane Feedrate (g/h) 
105-110 105-110 
Air Feedrate (l/h S.T.P.) 
ca. 350 ca. 350 
Cobalt concentration in 
catalyst solution returned 
ca. 1.0 ca. 0.7 
through pipe 12 (% w/w) 
Porportion as Co(+3) (%) 
30 40 
Oxygen consumption (g/h) 
70 64 
Acetic Acid production (g/h) 
76 63 
Weight Selectivities 
(g/100g butane consumed) 
Acetic Acid 175 150 
Propionic Acid 2 &lt;2 
Butyric Acid 1 &lt;2 
"Light-ends" -- 19 
"Heavy-ends" &lt;3 &lt;3 
Carbon Dioxide 31 27 
Carbon monoxide 2 2 
Composition of "light ends" 
withdrawn (approx. % w/w) 
Methyl ethyl ketone 70 
Sec-butyl acetate 20 
Ethyl acetate 5 
Methyl aceate 2.5 
Acetone 1.5 
______________________________________ 
EXAMPLE 2 
Example 1 was repeated except that the "light ends" of composition given in 
Table 2 were recycled to the reactor 1 through the line 7a. 
The results obtained over a period of about 80 hours of continuous stable 
operation are presented in Table 2. Continuous oxidation was maintained 
for a period of at least 240 hours without any significant reduction in 
the oxygen absorption rate. 
Examples 1 and 2 illustrate that no reaction-inhibiting species (whether 
inorganic or organic) accumulate rapidly within the reaction/catalyst 
recycle system when the reactor is fabricated from titanium i.e. in a 
substantially chromium-free reactor. 
EXAMPLE A 
The reactor shown in FIG. 3, the catalyst recovery system shown in FIG. 2 
and the "light ends" recovery system described in Example 1 were 
started-up and subsequently operated in a similar manner to the 
corresponding equipment in Example 1, except that all quantities and 
hourly feed rates were approximately 6-fold greater. 
Within about 8 hours from the commencement of the experiment the oxygen 
absorption rate was acceptably stable at 16 .+-. 2 moles per hour, with 
recycle of both cobalt catalyst solution and "light ends", at a reaction 
temperature of ca. 110.degree. C. and a total pressure of ca. 35 bar. 
However, after about 20 hours of continuous operation the rate of oxygen 
absorption began to fall, and the butane oxidation reaction ceased within 
24 hours. At the time when a significant reduction in the rate of oxygen 
absorption occurred the concentration of iron in the liquid reaction 
mixture withdrawn had risen to approximately 2000 ppm (by weight). It was 
only possible to restart the butane oxidation reaction after the reactor 
and catalyst recycle system had been drained and a fresh solution of the 
cobalt in acetic acid catalyst solution introduced. This behaviour was 
taken as an indication that a reaction-inhibiting species, probably a 
stainless steel corrosion product, had accumulated within the 
reactor/catalyst recycle system. 
The approximately composition of stainless steel used in the fabrication of 
the reactor is presented in Table 3, together with anticipated 
concentrations of other metal ions which might be present in a solution 
containing 2000 ppm of iron, resulting from the corrosion of such a 
stainless steel. However, insofar as certain components of the stainless 
steel may be dissolved preferentially, and some components may form 
compounds of limited solubility, the ratios of concentrations of metals in 
solution may differ somewhat from those in the attacked metal. 
Table 3 
______________________________________ 
Anticipated 
Concentration 
Approximate Composition 
of Metal in Solu- 
of Stainless Steel 
tion Containing 
Reactor Used 200 ppm Iron 
Metal (Example A) (% by weight) 
(ppm by weight) 
______________________________________ 
Iron 70 (2000) 
Chromium 18 510 
Nickel 8 230 
Manganese 
2 60 
Molybdenum 
2 60 
______________________________________ 
EXAMPLE B 
The titanium reactor illustrated in FIG. 1 and the catalyst separation 
apparatus illustrated in FIG. 2 were used in the method described in 
Example 2 i.e. with "light ends" recycle. Under stable operating 
conditions an oxygen consumption rate of approximately 2 moles/hour was 
obtained at 110.degree. C. and 35 bar absolute pressure to give an oxygen 
concentration in the off-gas of 6-8% by volume at a constant air flow rate 
(ca. 300 liters per hour, referred to S.T.P.). Thereafter only the oxygen 
concentration in the off-gas was monitored to assess the effect of 
introducing concentrated solutions in acetic acid of the soluble compounds 
of metals which are components of common stainless steels. Aliquots of the 
solutions were introduced into the cobalt catalyst recycle stream between 
the catalyst feed pot 27 and the recycle pump 28 (FIG. 2) to ensure rapid 
and complete transfer into the reaction vessel. The maximum concentration 
of added metal was estimated by assuming a feed-pipe reactor and reactor 
side-arm content of ca. 1.7 kg. The final distributed concentration of 
metal was estimated, assuming the additional contents of the liquid 
take-off pipe, catalyst recovery column, its (cold) feed vessel and the 
(cold) recycle catalyst solution holding/metering vessel to be ca. 1 kg. 
Preparation of metal solutions 
(i) Iron 
Ferric formate was dissolved in acetic acid. 
(ii) Nickel 
A nickel (II) - containing solution was prepared by dissolving nickel 
acetate in acetic acid containing a little water (ca. 10% w/w). 
(iii) Manganese 
A manganese (II)-containing solution was made up in the same manner as the 
nickel solution. 
(iv) Chromium 
Chromium (VI) - containing solutions were prepared by dissolving chromium 
trioxide in the minimum volume of water, followed by dilution of the 
resultant solution with acetic acid. 
All solutions were freshly prepared for each test and contained at least 1% 
w/w metal. 
Tests 
(a) Iron 
Several aliquots of a solution containing cobalt catalyst and stainless 
steel corrosion products, obtained at the end of Example A, were 
introduced to provide a measured iron content of ca. 900 ppm, together 
with unmeasured amounts of other metal ions, in the withdrawn liquid 
reaction mixture. No adverse effect was observed. Aliquots of the 
concentrated solution of iron were then introduced to raise the measured 
iron concentration of the liquid reaction mixture to over 2000 ppm. A high 
iron concentration (ca 1900 to over 2000 ppm) was then maintained over a 
12 hour period without any reduction of the oxygen absorption rate. 
Iron concentrations were determined by dissolving the sample in an excess 
of an aqueous solution of ammonia and ammonium chloride. The resulting 
precipitate of ferric hydroxide was separated by centrifuging followed by 
repeated water washings and digestion with 10% hydrochloric acid. The iron 
present was reduced to the ferrous state by hydroxylamine and was measured 
colorimetrically as the 1:10-phenanthroline complex. 
This method was also used to determine iron concentrations in Example A. 
(b) Chromium 
The introduction of approximately 1.2 g chromium in solution (to give a 
maximum Cr concentration of ca. 700 ppm) to the continuous reaction 
containing added iron led to an almost immediate cessation of oxygen 
absorption. 
The reactor and catalyst recycle apparatus were then drained, recharged 
with a fresh solution of cobalt catalyst in acetic acid and butane 
oxidation was re-started as before. When the oxygen concentration in the 
off-gas was acceptably stable, aliquots of chromium (VI) - containing 
solution, each containing approximately 0.6g chromium, were introduced at 
intervals of 6 to 9 hours. The resulting changes in the oxygen 
concentration of the off-gas are shown in FIG. 4, together with 
approximate values for the concentration of chromium, as measured by 
Atomic Absorption Spectroscopy, in the liquid reaction mixture withdrawn. 
(c) Nickel and manganese 
In a similar experiment to those described in (a) and (b) above it was 
further demonstrated that the simultaneous presence of nickel (2.1g total) 
at approximately 750 ppm and manganese (1.1g total) at approximately 400 
ppm did not lead to a permanent, significant reduction in the rate of 
oxygen absorption. Such concentrations of nickel and manganese are far in 
excess of those which might be present in solutions containing 400 ppm 
chromium resulting from the corrosion of the stainless steel used in the 
oxidation described in Example A. 
Examples A and B are not examples according to the invention and are 
included for the purpose of comparison. Example A clearly shows that it is 
not possible to maintain a continuous reaction in a stainless steel 
reactor with recycle of both cobalt catalyst solution and "light ends" for 
a significant period of time. Example B shows that the reason for the 
inability to maintain a continuous reaction is the presence of chromium in 
the corrosion products. Examples 1 and 2 show that it is possible to 
maintain a continuous reaction for lengthy periods both with and without 
"light ends" recycle in the substantial absence of chromium. 
EXAMPLE 3 
With reference to FIG. 3, the stainless steel reaction vessel 29, and 
immediately accociated sections of piping, were replaced by essentially 
identical items fabricated from titanium. The titanium reaction vessel was 
charged with 8 liters of glacial acetic acid and was pressurized to about 
35 bar (absolute). A small nitrogen flow was then established through the 
pipe 36. Approximately 1 kg. of butane was charged rapidly to the reactor 
through the line 37, and the butane feed rate was then adjusted to ca. 0.6 
kg/h. The oil-heating system, 30 was switched on, and a cobalt catalyst 
solution, comprising 2.2% w/w cobalt, 68% as Co(+3), in acetic acid 
containing ca. 4% w/w water, was introduced through line 39 at a rate of 
ca. 1.2 l/h. 
When the reactor temperature reached 60.degree. C. air was substituted for 
the nitrogen stream. The oxidation reaction commenced within 15 minutes of 
attaining a reaction temperature of 120.degree. C. The reaction 
temperature was stabilised at 125.degree. C. and the air and butane feed 
rates were adjusted to provide an oxygen content in the off-gas of ca. 4% 
v/v and a (net) acetic acid production rate of approximately 450 g/h. 
Liquid reaction mixture withdrawn from the reactor through the valve 33 was 
passed to a distillation column (approximately 3.5 mm in diameter and 1.8 
m in height, packed with 6 mm ceramic Raschig rings), fitted with a 
specially designed low residence time reboiler fabricated in titanium, to 
provide a recovered catalyst solution containing cobalt. Recycle of this 
catalyst solution through line 39 commenced approximately 3 hours after 
start-up. 
"Light ends" were separated from the reaction products in a 20-plate 
Oldershaw column of approximately 50 mm diameter in a manner similar to 
that described in Example 1. The "light ends" were thereafter recycled to 
the reaction vessel. 
The continuous oxidation process was operated for a period of approximately 
16 days with occasional interruptions, during which period no significant 
interruption of the oxidation reaction occured and no material was added 
to or withdrawn from the recycled catalyst solution, over about 7 
consecutive days. During the period of operation there was no evidence of 
inhibition of the reaction, and reaction conditions were varied to 
investigate the effects of individual reaction parameters. Results 
obtained over 8 hours towards the end of this period are given in Table 4. 
TABLE 4 
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Example 3 
"Light-ends" Recycled 
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Reaction Temperature (.degree. C) 
130.5 
Reaction Pressure (bar gauge) 
48.3 
Butane Feedrate (kg/h) 
ca. 1.5 
Air Feedrate (m.sup.3 /h at STP) 
ca. 4.6 
Cobalt Concentration in Reactor 
Product (Butane free) (% w/v) 
0.28 
Proportion as Co(+3) (%) 26 
Cobalt Concentration in Recycled 
Catalyst Solution (% w/v) 0.86 
Proportion as Co(+3) (%) 10 
Oxygen Consumption Rate (Kg/h) 
1.15 
Acetic Acid Production 
Rate (kg/h) 1.07 
Weight Selectivities 
(g/100 g butane consumed) 
Acetic Acid 169 
Propionic Acid 3 
Butyric Acid 2 
Carbon Dioxide 39 
Carbon Monoxide 2 
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