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Title:
PROCESS FOR SULFONATING HALOBENZENE DERIVATIVES WITH SULFUR TRIOXIDE
Document Type and Number:
WIPO Patent Application WO/2014/029666
Kind Code:
A1
Abstract:
A process for sulfonating at least one halobenzene with sulfur trioxide (SO3) comprising the following steps : Step 1. manufacturing a gaseous mixture [mixture (M)] comprising SO3 and at least one additional gas different from SO3 by oxidizing sulfur dioxide in the presence of at least one catalyst, wherein the SO3 content in mixture (M) is from 1 to 95 % by volume, relative to the total volume of mixture (M); and Step 2. contacting said mixture (M) with said at least one halobenzene.

Inventors:
DOSI, Mahendra (310 Black Swan Way, Alpharetta, Georgia, 30022, US)
HUSEIN, Ziad (1111 Rivershyre Drive, Evans, Georgia, 30809, US)
MYSONA, Ronald (876 Sturbridge Dr, Evans, Georgia, 30809, US)
Application Number:
EP2013/066909
Publication Date:
February 27, 2014
Filing Date:
August 13, 2013
Export Citation:
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Assignee:
SOLVAY SPECIALTY POLYMERS USA, LLC (4500 McGinnis Ferry Road, Alpharetta, Georgia, 30005-3914, US)
International Classes:
C07C303/06; C07C309/39
Foreign References:
US3427342A1969-02-11
US2835708A1958-05-20
US3372188A1968-03-05
US5264200A1993-11-23
US7740827B22010-06-22
US6572835B12003-06-03
GB2088350A1982-06-09
US7968742B22011-06-28
US5911958A1999-06-15
Other References:
SPRYSKOW, A.A. ET AL.: "Orientation in substitution in the aromatic nucleus. IV. Sulfonation of chlorobenzene", JOURNAL OF GENERAL CHEMISTRY USSR, vol. 28, 1958, pages 2250 - 2254, XP009168884, ISSN: 0022-1279
"Sulfur trioxide", WIKIPEDIA, THE FREE ENCYCLOPEDIA, 18 April 2013 (2013-04-18), pages 1 - 4, XP055060402, Retrieved from the Internet [retrieved on 20130419]
Attorney, Agent or Firm:
BENVENUTI, Federica et al. (Solvay SA, Intellectual Assets ManagementRue de Ransbee, 310 Bruxelles, B-1120, BE)
Download PDF:
Claims:
C L A I M S

1. A process for sulfonating at least one halobenzene with sulfur trioxide (S03) comprising the following steps :

Step 1. manufacturing a gaseous mixture [mixture (M)] comprising SO3 and at least one additional gas different from SO3 by oxidizing sulfur dioxide in the presence of at least one catalyst, wherein the SO3 content in mixture (M) is from 1 to 95 % by volume, relative to the total volume of mixture (M) ; and

Step 2. contacting said mixture (M) with said at least one halobenzene.

2. The process according to claim 1, wherein in Step 1. the additional gas different from SO3 is selected from a group consisting of air, nitrogen, carbon dioxide, oxygen, sulfur dioxide (S02), in particular unconverted S02 and mixtures thereof.

3. The process according to any one of claims 1 or 2, wherein in Step 1. the mixture (M) is a SC^/unconverted S02/air in which the SO3 content is from 1 to 95 % by volume, relative to the total volume of the SC^/air mixture.

4. The process according to any one of claims 1 to 3, wherein in Step 1. the oxidation of S02 to SO3 in the presence of at least one catalyst is carried out in the presence of air, oxygen enriched air or neat oxygen or oxygen-containing gases, preferably in the presence of air.

5. The process according to claim 4, wherein gaseous S02 is mixed with preheated air to form a gaseous sulfur dioxide/air mixture [mixture (MSo2/air)] .

6. The process according to claim 5, wherein in Step 1. the S02 content in the mixture (Mso2/air) is from 0.1 to 30 % by volume relative to the total volume of the mixture (MSo2/air).

7. The process according to any one of claims 5 or 6, wherein the 02 content in the mixture (Mso2/air) is from 15 to 25 % by volume relative to the total volume of the mixture (MSo2/air).

8. The process according to any one of claims 1 to 7, wherein in Step 2. the mixture (M) is contacted with at least one halobenzene without any preliminary separation/purification step of said mixture (M).

9. The process according to any one of claims 1 to 8, wherein in Step 2. the SOs/halobenzene molar ratio is from 0.17 to 3.

10. The process according to any one of claims 1 to 9, wherein in Step 2. the feeding of halobenzene to the reactor is realized by a one single batch addition at the start of the reaction, by sequentially multiple batch additions or a continuous feed of halobenzene.

11. The process according to claim 10, wherein the feeding of

halobenzene to the reactor is a continuous feed of halobenzene with a flow feeding rate equal to or less than 22000 liter per hour (1/h).

12. The process according to any one of claims 1 to 11, wherein in Step 2. the mixture (M) is fed to the reactor with a flow feed rate equal to or less than 1500 liter per hour (1/h).

13. The process according to any one of claims 1 to 12, wherein Step 2. is carried out at a temperature of below 200°C.

14. The process according to any one of claims 1 to 13, wherein Step 2. is carried out at a pressure of below 10 atm, preferably at atmospheric pressure.

15. The process according to any one of claims 1 to 14, wherein the halobenzene is monochlorobenzene (MCB).

Description:
PROCESS FOR SULFONATING HALOBENZENE DERIVATIVES WITH SULFUR TRIOXIDE

This application claims priority to U.S. provisional application

No. 61/684885 filed on 20 Aug 2012 and to European application

No. 12194002.7 filed on 23 Nov 2012, the whole content of each of these applications being incorporated herein by reference for all purposes.

Technical Field

The present invention relates to a process for sulfonating halobenzene with gaseous sulfur trioxide (S0 3 ) mixtures.

Background Art

Sulfuric acid (H 2 SO 4 ) and oleum (which is a solution of sulfur trioxide in concentrated sulfuric acid namely SO 3 · H 2 SO 4 ) are widely used as sulfonating agents for the sulfonation of aromatic compounds. Sulfuric acid (H 2 SO 4 ) and oleum always need to be used in large excess as water is formed in the

sulfonating reaction thereby diluting the oleum and/or sulfuric acid. This has the disadvantage of leaving large quantities of unreacted sulfuric acid. This waste acid must be separated from the reaction mixture and subsequently disposed of. This acid is difficult to dispose of, either as the free acid or in the form of soluble or insoluble sulfates, particularly now when effluent requirements are becoming more stringent.

Alternatively, sulfur trioxide itself in gaseous or liquid form, has also been used for the sulfonation of aromatic compounds. Sulfur trioxide reacts

instantaneously with aromatic compounds, and it is not necessary to use a substantial excess to realize complete sulfonation. There is, therefore, no need for the reaction product to be contaminated with substantial quantities of excess sulfonating agent. Sulfur trioxide itself can notably be prepared from oleum.

For example, distillation of oleum can generate pure SO 3 vapour which can be used as such or can be further condensed to form pure liquid SO 3 . However, pure gaseous or liquid sulfur trioxide is very highly reactive and its reactions with aromatic compounds are extremely exothermic and difficult to control and undesirable side reactions might occur. It is known that in order to moderate and control the reactions of pure gaseous or liquid SO 3 with aromatic compounds, SO 3 has been used in the presence of inert diluents. For example, air/SC^ sulfonation processes are widely used in the surfactants and detergent industry for the sulfonation of long chain and/or high molecular weight organics, such as for example low volatile long chain alkyl benzenes, which are characterized by having high viscosity and high flash points.

The diluted, gaseous SO 3 mixture is still a very aggressive/reactive material.

For this reason, it is known that the use of the air/S0 3 sulfonation process for the sulfonation of more volatile compounds such as notably toluene, xylene and other lower alkyl benzenes are problematic and said process is in general less suitable for the sulfonation of volatile aromatic compounds.

Thus, there is still a considerable need for a process for sulfonating volatile halobenzene compounds with sulfur trioxide (S0 3 ) in a controllable manner, which avoids the use of liquid S0 3 which entails some serious concern for its security, transportation, and rigorous storage requirements imposed by the hazardous nature of liquid S0 3 , capable of providing para-substituted

halobenzene sulfonic acid compounds in high yield and 4,4'-dihalodiphenyl sulfone, both in high purity and high selectivity, to minimize of any formation of oversulfonated products as a result of undesirable side reactions and avoid unconverted reactants, yet without any significant problem of disposal of excess sulphuric acid, much less hazardous for the environment and simultaneously in achieving significant cost savings.

Summary of invention

The present invention thus relates to a process for sulfonating at least one halobenzene with sulfur trioxide (S0 3 ) comprising the following steps :

Step 1. manufacturing a gaseous mixture [mixture (M)] comprising S0 3 and at least one additional gas different from S0 3 by oxidizing sulfur dioxide (S0 2 ) in the presence of at least one catalyst, wherein the S0 3 content in mixture (M) is from 1 to 95 % by volume, relative to the total volume of mixture (M) ; and

Step 2. contacting said mixture (M) with said at least one halobenzene.

For the purpose of the present invention, the term "halobenzene" is intended to denote any halogenated derivative of benzene. It may be mono-, di-or tri-halogenated. The halobenzene is preferably a monohalobenzene where the halogen atom is chosen from chloride, fluoride, bromide and iodide. More preferably, the halobenzene is monochlorobenzene (MCB). Detailed description of embodiments

Mixture (M)

As said, the mixture (M) prepared in the first step of the process of the present invention comprises SO 3 and at least one additional gas different from S0 3 .

For the purpose of the present invention, the additional gas different from SO 3 is an inert gas.

For the purpose of the present invention, the term "inert gas" denotes a gas which is substantially unreactive during the course of the sulfonation reaction.

The additional gas different from SO 3 is preferably selected from a group consisting of air, nitrogen, carbon dioxide, oxygen, sulfur dioxide (S0 2 ), in particular unconverted S0 2 and mixtures thereof. More preferably, the additional gas different from SO 3 is nitrogen, air or a mixture of nitrogen and air or a mixture of air and unconverted S0 2 . Most preferably, the additional gas different from SO 3 is a mixture from air and unconverted S0 2 .

The expression 'air' is to be understood according to its usual meaning, i.e. atmospheric air at sea level, which typically includes about 78.1 % by volume nitrogen gas (N 2 ), about 20.9 % by volume oxygen gas (0 2 ) and less that 1 % by volume of Argon (Ar) relative to the total volume of air.

In the process according to the invention and in the particular embodiments thereof, the SO 3 content in mixture (M) is advantageously from 1 to 95 % by volume, preferably from 1 to 70 % by volume, more preferably from 2 to 50 % by volume, even more preferably from 3 to 20 % by volume, most preferably from 4 to 8 % by volume relative to the total volume of mixture (M).

If desired, the mixture (M) consists of SO 3 and the additional gas different from S0 3 .

In a particular preferred embodiment of the present invention, the mixture (M) is a SCVunconverted S0 2 /air mixture in which the SO 3 content is advantageously from 1 to 95 % by volume, preferably from 1 to 70 % by volume, more preferably from 2 to 50 % by volume, even more preferably from 3 to 20 % by volume, most preferably from 4 to 8 % by volume relative to the total volume of the SCVunconverted S0 2 /air mixture.

For the purpose of the present invention, the term "unconverted S0 2 " refers to the S0 2 that has not been oxidized.

In one embodiment of the present invention, the unconverted S0 2 content in the SCVunconverted S0 2 /air mixture is in general equal to or below 10 % by mole, preferably equal to or below 5 % by mole, more preferably equal to or below 4 % by mole, most preferably equal to or below 2 % by mole, based on the total moles of S0 3 and unconverted S0 2 .

Good results were obtained with a 2 % by mole to 4 % by mole of unconverted S0 2 content in the SCVunconverted S0 2 /air mixture, based on the total moles of SO 3 and unconverted S0 2 .

In Step 1. of the present invention, the manufacturing of mixture (M) is performed by oxidizing sulfur dioxide (S0 2 ) in the presence of at least one catalyst.

For the purpose of the present invention, any catalyst known in the art that can be used to oxidize S0 2 to SO 3 is suitable.

Typical examples of suitable catalysts that may be used in the present invention include, but are not limited to, solid particulate catalysts typically containing an alkali- vanadium or platinum-containing active phase such as notably described in U.S. Pat. Nos. 5,264,200, the whole content of which is herein incorporated by reference, commercially available vanadium

pentoxide (V 2 0 5 ), ruthenium oxide such as notably described in U.S. Pat.

Nos. 7,740,827, the whole content of which is herein incorporated by reference.

Good results were obtained with commercially available vanadium

pentoxide (V 2 0 5 ).

If desired, the oxidation of S0 2 to SO 3 can be carried out in a multiple stage catalytic converter such as notably described in U.S. Pat. Nos. 7,740,827 and references therein. A specific example of a multiple stage catalytic converter is notably a three-stage vanadium pentoxide catalytic converter.

In Step 1. of the process of the invention, the oxidation of S0 2 to SO3 in the presence of at least one catalyst is typically carried out in the presence of air, oxygen enriched air or neat oxygen or oxygen-containing gases. Preferred, said oxidation is carried out in the presence of air.

In one specific embodiment of Step 1. of the present invention, gaseous S0 2 is mixed with preheated air to form a gaseous sulfur dioxide/air mixture

[mixture (M S0 2/air)] .

In this specific embodiment, the preheated air has advantageously a temperature equal to or less than 600°C, more preferably equal to or less than 480°C.

The preheated air has advantageously a temperature equal to or above 350°C, more preferably equal to or above 400°C. Preheated air having a temperature of 450°C gave particularly good results.

The preheated air, mentioned above, is preferably dry and the removal of water in the preheated air may be accomplished by any conventional means such as notably vapor tubes and the like. The water content in the preheated air is advantageously equal to or less than about 0.01 % by volume, preferably equal to or less than 0.001 % by volume, based on the total volume of the

mixture (M S0 2/air).

It is known in the art that a specific temperature range from 416 to 454°C is favorable to initiate the catalytic conversion of S0 2 to S0 3 . If desired, the temperature can be changed after the conversion reaction is initiated.

The conversion of S0 2 to SO 3 is an equilibrium reaction

(S0 2 +[l/2]0 2 → SO 3 ). The oxygen required to convert S0 2 to SO 3 is typically provided by the air in the mixture (M S o 2 /air). The percentage of S0 2 which can be converted to SO 3 varies with temperature and with the concentration (partial pressure) of the gaseous initial reactants, namely S0 2 and 0 2 . The person skilled in the art will use standard techniques and routine work so as to determine temperature and concentration of the gaseous initial reactants for obtaining the final desired conversion of S0 2 to SO 3 . The conversion of S0 2 to SO 3 is typically in the range from 96 % to 98 %, preferably exceeds 96 %.

It is known that the lower the temperature in the standard temperature range at which the conversion reaction occurs, the greater the conversion of S0 2 to SO 3 . For a given concentration of reactants and assuming the conversion reaction proceeds to equilibrium, there is a theoretical conversion percentage of S0 2 to SO 3 at each temperature within the range at which conversion can be sustained. The conversion temperature range has maximum and minimum temperatures. Maximum theoretical conversion occurs at the minimum temperature at which conversion can be sustained. Depending upon the concentration of the reactants, maximum theoretical conversion can be

99 percent or more, at a minimum sustaining temperature of for example 400°C.

As noted above, there is a maximum temperature at which the conversion reaction can be sustained, and the maximum sustaining temperature decreases as the conversion percentage increases. For example, depending upon the concentration of the initial reactants, at a temperature of about 600°C the conversion reaction reaches equilibrium when the theoretical SO 3 percentage is about 70 percent ; a lower temperature, e.g., about 480°C or below, may be required to obtain a theoretical conversion of 95 percent, and a temperature of about 400°C may be required to obtain a theoretical conversion of 99 percent.

In this specific embodiment of Step 1. according to the present invention, the S0 2 content in the mixture (M S o2/air) is from 0.1 to 30 % by volume, more preferable from 0.5 to 15 % by volume, most preferably from 1 to 5 % by volume relative to the total volume of the mixture (Mso2/air).

In this specific embodiment of Step 1. according to the present invention, the 0 2 content in the mixture (M S o2/air) is from 15 to 25 % by volume, most preferably from 19 to 21 % by volume relative to the total volume of the mixture (M S0 2/air).

Gaseous S0 2 is generally obtained by evaporation of liquid S0 2 which is for example commercially available in large containers as well as in small lab cylinders. The evaporated S0 2 can then be mixed with preheated air to form the mixture (M S o2/air), as described above.

In an alternative embodiment, a gaseous sulfur dioxide/air mixture can be produced by reacting sulfur and air in a sulfur burner such as notably described in U.S. Pat. Nos. 6,572,835 Bl the whole content of which is herein incorporated by reference and UK published patent application GB 2 088 350 A, the whole content of which is herein incorporated by reference.

The Applicant has surprisingly found that the mixture (M) as mentioned above, is effective in sulfonating at least one halobenzene by contacting said at least one halobenzene.

In a preferred embodiment of the present invention, the mixture (M) is contacted with at least one halobenzene without any preliminary

separation/purification step of said mixture (M).

If desired, the mixture (M) can be further purified by notably removing the unconverted S0 2 , or can be further enriched or diluted with the at least one additional gas different from SO3, as described above, before said mixture (M) is contacted with at least one halobenzene.

If desired, said mixture (M) can be stored.

In a particular preferred embodiment of the present invention, the process for sulfonating at least one halobenzene with sulfur trioxide (SO 3 ) comprising the Steps 1. and Steps 2. are carried out with no intermediate separation or storage of the mixture (M). Step 2. according to the present invention is advantageously carried out by contacting the mixture (M) with the at least one halobenzene in a suitable reactor.

The choice of the reactor is not critical, provided that the reactor can enable an efficient contact between the mixture (M) and at least one halobenzene, a removal of the exothermic heat of reaction, and is foreseen with means for escape of effluent vapor streams and avoid S0 3 loss in said effluent vapor streams.

Among suitable reactors, mention can be made of, but not limiting to film reactors, including notably the Falling Film Reactor as described for example in U.S. Pat. Nos. 7,968,742 B2 and 5,911,958, the whole content of which is herein incorporated by reference, dispersed phase or jet reactors, stirred tank reactors, continuous stirred tanks (CSTR) and cascades of at least two or 3 continuous stirred tanks which can optionally be equipped with a means to enable a counter- current flow of the halobenzene and the mixture (M). Preferred will be the Falling Film Reactors, stirred tank reactors and cascades of at least two or 3 continuous stirred tanks optionally equipped with a means to enable a counter-current flow of the halobenzene and the mixture (M).

In general, the reactor is also additionally equipped with at least one nozzle-set ; stirring means ; means for feeding the reactants including notably gas flow means, dosing/metering systems such as for example a rotameter for converter gas dosing or a semi-automatic burette for liquids, and the like ;

temperature control means, such as for example a thermocouple and heat exchangers so as to obtain carefully controlled conditions of SOs/halobenzene molar ratio, temperature and SO 3 concentration.

Any nozzle-set suitable for mixing of the halobenzene and the mixture (M) can be used.

Among suitable nozzle-sets, mention can be made of, but not limiting to a ring fitted with multi-nozzles, 12-nozzle disperser, or a flow pipe with a sparger. Non limiting examples of stirring means in the reaction medium may be notably means of internal stirring such as a turbine or an agitator, or by means of a recirculation pipe exterior to the reactor. Optimal stirring can advantageously ensure good mixing, in particular good dispersion, of the mixture (M) into halobenzene.

Any heat-exchanger suitable for reconverting by cooling the volatized reaction medium to liquid, in particular the halobenzene, so that said liquid can be returned to the reaction mixture can be used. A typical heat-exchanger is notably a water-cooled condenser. It has been found that the heat-exchanger advantageously avoid the loss of the halobenzene in the effluent vapor stream.

In Step 2. of the process according the present invention, the

SOs/halobenzene molar ratio is advantageously from 0.17 to 3, preferably from 0.25 to 1.2, more preferably from 0.5 to 1.0.

In Step 2. of the process according the present invention, the

SOs/halobenzene molar ratio is advantageously above 0.17, preferably

above 0.25.

For the purpose of the present invention, the SOs/halobenzene molar ratio is defined as the ratio between the total number of moles of S0 3 , present in the mixture (M), as described above, fed to the reactor divided by the total number of moles of halobenzene fed to the same reactor.

In a specific preferred embodiment of the present invention, in Step 2. of the process, the SO 3 /MCB molar ratio is advantageously from 0.17 to 3, preferably from 0.25 to 1.2, more preferably from 0.5 to 1.0.

In this specific preferred embodiment, the SO 3 /MCB molar ratio is advantageously above 0.17, preferably above 0.25.

The SOs/halobenzene molar ratio can be varied or kept constant during the reaction time depending on the feed of the mixture (M) and the feed of the halobenzene, in particular MCB, to the reactor.

In Step 2. of the process according the present invention, the feeding of halobenzene, in particular MCB to the reactor can be realized by a one single batch addition at the start of the reaction, by sequentially multiple batch additions or a continuous feed of halobenzene.

In one embodiment of Step 2. of the present invention, the halobenzene is charged into the reactor in one single batch at the start of the reaction.

In another embodiment of Step 2. of the present invention, the halobenzene is sequentially fed to the reactor in multiple batch additions.

In another embodiment of Step 2. of the present invention, the halobenzene is fed to the reactor by a continuous feed of the halobenzene. In this embodiment, it is beneficial, to keep the SOs/halobenzene molar ratio constant by controlling the feed flow rate of the halobenzene, in particular MCB, during the reaction time. Generally the feed flow rate of the halobenzene, in particular MCB is adapted to the feed flow rate of the mixture (M) in order to comply with the desired SOs/halobenzene molar ratio, as mentioned above. Generally the flow rate of said feeding is equal to or less than 22000 1 per hour (1/h), preferably equal to or less than 15000 1/h, more preferably equal to or less than 2000 1/h, even more preferably equal to or less than 240 1/h.

Generally, the flow rate of the feeding of the halobenzene is equal to or more than 0.1 1/h, preferably equal to or more than 10 1/h, more preferably equal to or more than 140 1/h.

In Step 2. according to the invention and in the particular embodiments thereof, it is especially beneficial to control the feeding of the mixture (M) to the reactor so as to avoid notably that halobenzene is not volatilized by the heat of the reaction more rapidly than the heat-exchanger is able to reconvert it to the liquid state. Generally the flow rate of said feeding is equal to or less than 1500 liter per hour (1/h), preferably equal to or less than 500 liter per hour (1/h), more preferably equal to or less than 380 liter per hour (1/h).

Generally, the flow rate of the feeding of the mixture (M) is equal to or more than 10 liter per hour (1/h), preferably equal to or more than 300 liter per hour (1/h), more preferably equal to or more than 320 liter per hour (1/h).

Flow rates of the feeding of the mixture (M) In Step 2. of the process according the present invention from 320 1/h to 380 1/h gave particularly good results, a flow rate of 350 1/h gave excellent result.

The flow rates of the feeding of the mixture (M) are typically controlled by gas flow means known to the person skilled in the art such as for example a gas rotameter, mass flow meter and the like.

Step 2. of the process according the present invention is preferably carried out at a temperature of below 200°C, more preferably of below 160°C, still more preferably of below 130°C and most preferably of below 100°C.

Step 2. of the process according the present invention is preferably carried out at a temperature of above -40°C, more preferably of above 0°C, still more preferably of above 20°C and most preferably of above 40°C.

Step 2. of the process according the present invention is preferably carried out at a pressure of below 10 atm, more preferably of below 7 atm, still more preferably of below 5 atm and most preferably of below 2 atm.

Step 2. of the process according the present invention is preferably carried out at a temperature of above 0.5 atm, more preferably of above 0.6 atm, still more preferably of above 0.7 atm and most preferably of above 0.8 atm.

Excellent results were obtained when Step 2. of the process according the present invention was carried out at atmospheric pressure. It has been found that the process according to the invention allows a decreased loss of SO 3 in the effluent vapor streams.

It is understood that the different processes and embodiments disclosed herein apply in a most preferred way to the sulfonation of monochlorobenzene (MCB) with sulfur trioxide (SO3) yielding a sulfonated product mixture comprising mainly the desired para-chlorobenzene sulfonic acid (para-CBS A) and the desired 4,4'-dichlorodiphenyl sulfone together with minor amounts of ortho-chlorobenzene sulfonic acid (ortho-CBS A), meta-chlorobenzene sulfonic acid (meta-CBSA), 2,4'-dichlorodiphenyl sulfone, 3,4'-dichlorodiphenyl sulfone, 2 sulfone sulfonic acid isomers, sulfuric acid and unreacted halobenzene. It has been suggested that in the sulfone sulfonic acid isomers, an additional sulfonating of one of the two benzene rings in the sulfone has taken place.

Should the disclosure of any patents, patent applications, and publications which are incorporated herein by reference conflict with the description of the present application to the extent that it may render a term unclear, the present description shall take precedence.

EXAMPLES

The invention will now be described in more details with reference to the following examples, whose purpose is merely illustrative and not intended to limit the scope of the invention.

General procedure for the sulfonation of monochlorobenzene (MCB)

Liquid S0 2 supplied in commercial lab cylinders was heated and vaporized, said vaporized S0 2 was flowed into to the three-stage vanadium pentoxide (V 2 O 5 ) catalytic converter along with an air stream preheated at 450°C. The flow rates of SO 2 and air in the feed were adjusted to yield a desired SO 3 content of 8 % by volume relative to the total volume of air/SC^ product stream. The SO 2 was converted to SO 3 up to 96-98 % thereby providing an air/S03/S02(unconverted) mixture which is subsequently used without further purification. Said air/SCVSC^unconverted) mixture obtained was continuously fed with a flow rate of 350 liter/hr (controlled by a rotameter) into a jacketed cylindrical glass flask, which was one-liter in size, fitted with a stirrer, a condenser and a thermocouple. Said air/SCVSC^unconverted) mixture was dispersed into the liquid content of the reactor through a 12-nozzle disperser. The reactor temperature, as indicated in Table 1 was controlled by circulating water in the reactor jacket from a water bath. Liquid monochlorobenzene

(MCB) was initially charged into the glass flask and/or added in batch to the flask and/or continuously fed into the reactor using a semi-automatic burette with a specific flow rate as specified below. The loss of SO 3 , present in the non- condensable gas leaving the condenser was scrubbed with water in three stages. The loss of SO 3 in reactor effluent vapor was determined using acid analysis of the scrubbed water samples. During the reaction and at the end of the reaction, the reaction mixture and the final reaction product mixture, respectively were analysed by GC, LC, GC/MS (gas chromatography coupled to mass

spectrometry) and LC/MS (liquid chromatography coupled to mass

spectrometry). The experimental data are summarized in Tables 1 and 2.

Example 1 :

The general procedure, as detailed above, was followed whereby liquid monochlorobenzene (MCB) was initially charged into the glass flask in an amount of 250 ml and with no additions of MCB. The sulfo nation was carried out at a temperature of 30°C.

Example 2 :

The general procedure, as detailed above, was followed whereby liquid monochlorobenzene (MCB) was initially charged into the glass flask in an amount of 250 ml and after every 30 minute, 20 ml of MCB was added in batch. Initially the reaction temperature was maintained at 30°C. The temperature was increased to 45°C, 60°C and 75°C at 60, 90 and 120 minutes, respectively. The reaction at 75°C was extended for another 30 minutes with no further addition of MCB.

Comparative example 3 (Cex 3)

Liquid monochlorobenzene (MCB) and liquid SO 3 were charged continuously to a larger production-scale continuous reactor. The liquid SO 3 and liquid MCB were fed with a flow rate of 500 Kg/hr (260 liter/hr) and 1500 Kg/hr (1350 liter/hr), respectively. The mole ratio liquid SCVliquid MCB was 0.47. The MCB entered the reactor through the reactor vent scrubber which eliminated any loss of SO 3 vapour into the vent. The reactor temperature was maintained constant at 70°C. Since the sulfonation reaction is exothermic, reactor liquid was cooled by recirculation through an external heat exchanger. After steady state reaction conditions were attained, the reactor products were removed continuously keeping the liquid level constant in the reactor. Table 1

It has been found that the overall chemical reactions taking place in

Examples 1 and 2, carried out according to he present invention by using an air/SCVSC (unconverted) mixture yielded no new products when compared with the products obtained by sulfonation of MCB using liquid SO 3 (Cex 3) and the desired para-chlorobenzene sulfonic acid and 4,4'-dichlorodiphenyl sulfone could be obtained in higher yield.