PROCESS FOR MANUFACTURING ALKANESULFONIC ACIDS

Abstract

The present invention relates to an improved process for manufacturing of alkanesulfonic acids.

Claims

1. A process for the production of alkane sulfonic acid comprising reacting sulfur trioxide and an alkane with a starter, selected from the group consisting of inorganic peroxoacids, salts of inorganic peroxoacids, organic peroxoacids, salts of organic peroxoacids, hydrogen peroxide, and mixtures thereof, and/or starter precursor, selected from the group consisting of sulfuric acid, oleum, SO3, alkanesulfonic acid, the bottom recycle stream from alkanesulfonic acid distillation, inorganic oxoacids, salts of inorganic oxoacids, hydrogen peroxide, and mixtures thereof, wherein addition of the starter and/or starter precursor to a reactor cascade is split, wherein synthesis is realized as batch process in a reactor cascade and the starter or starter precursor is divided and added in portions to a first reactor and to further reactors of the cascade, or wherein synthesis is realized as continuous process and the starter or starter precursor is divided and added continuously to a first reactor and to further reactors of the cascade.

2. The process according to claim 1, wherein the synthesis is realized as continuous process and the total amount of starter or starter precursor x is continuously divided over the n reactors or over n minus 1 reactors in a homogenous manner, wherein each of the n reactors is supplied with a continuous fraction of x/n mol of the starter.

3. The process according to claim 1, wherein the synthesis is realized as continuous process and the largest amount of starter or starter precursor is fed continuously into the first reactor of the cascade while the remaining starter is divided evenly over the remaining reactors of the cascade or over the remaining reactors but the last of the cascade.

4. The process according to claim 1, wherein the synthesis is realized as continuous process and a certain amount of starter or starter precursor is added to the first reactor while all the rest is fed into the last reactor of the cascade.

5. The process according to claim 1, wherein the starter and/or starter precursor is provided in a solvent.

6. The process according to claim 1, wherein the reaction mixture leaving the last reactor is purified by distillation.

7. The process according to claim 6, wherein a recycle stream leaving the distillation as a bottom residue is cooled prior to mixing it with the starter solution or the starter precursor solution.

8. The process according to claim 7, wherein the recycle stream is cooled to temperatures <25° C.

9. The process according to claim 7, wherein the starter solution and/or starter precursor solution is cooled while mixing it with the recycle stream.

10. The process according to claim 7, wherein the starter solution and/or starter precursor solution is cooled in a first step, and then it is mixed under further cooling at temperatures <50° C. with a pre-cooled recycle stream leaving the distillation as a bottom residue.

11. The process according to claim 1, wherein the alkane is methane and the alkanesulfonic acid is methanesulfonic acid, and wherein optionally methane is supplied to each reactor.

12. The process according to claim 1, wherein methane pressure is adjusted reactor by reactor within a range of 10 to 200 bar.

13. The process according to claim 1, wherein methane is fed at the same pressure into the first reactor(s) of a cascade, whereas the last reactor is constantly operated at a higher methane pressure than the first reactor(s) of the cascade.

14. The process according to claim 1, wherein at least one reactor type used in the cascade is a continuously stirred tank reactor.

15. The process according to claim 1, wherein at least one reactor type used in the cascade is an air lift reactor, a loop reactor, a bubble column or a trickle bed reactor.

16. The process according to claim 1, wherein conversion of SO3 and the formation of MSA are measured with density, and/or addition of water is measured with a conductivity measurement and/or ultrasound measurements.

17. The process according to claim 1, wherein side products are detected by spectroscopic methods.

18. The process according to claim 1, wherein mass flow rate of purge stream is adjusted according to the amount of unconverted SO3 and the amount of sulfuric acid formed during the synthesis of the initiator.

19. The process according to claim 1, wherein feed to the distillation in has an MSA content of 50 to 99 wt.-%, and/or feed to the reaction into the first reactor has a SO3 content of 30 to 100 wt.-%.

20. The process according to claim 1, wherein a high conversion of SO3 to alkane sulfonic acid is achieved and the amount of waste is reduced.

21. The process according to claim 1, wherein the addition of the overall starter amount is divided over several addition points in the system.

22. The process according to claim 5, wherein the solvent comprises sulfuric acid and/or alkanesulfonic acid.

23. The process according to claim 13, wherein the last reactor is constantly operated at a higher methane pressure by up to 20 bar than the first reactor(s) of the cascade.

24. The process according to claim 17, wherein side products are detected by RAMAN spectroscopy and/or NMR spectroscopy.

Description

[0027] FIG. 1 shows one example how a process for the production of MSA by reaction of methane and sulfur trioxide can be set up and will be explained below in an exemplary manner.

[0028] FIG. 2 shows an exemplary reactor set up for an embodiment of the inventive process.

[0029] A Reaction: Methanesulfonic acid (MSA) is formed from methane and SO.sub.3 with a mixture of MSA and sulfuric acid as a solvent.

SO.sub.3+CH.sub.4.fwdarw.CH.sub.3—SO.sub.3H −134 kJ/mol   1 [0030] The reaction usually takes place at pressures from 10 to 200 bar and temperatures ranging from 25 to 100° C. In a preferred mode, the pressure is between 30 to 150 bar, more preferred between 50 and 110 bar, most preferred between 70 and 100 bar. In a preferred mode of operation, the temperature is between 40 to 80° C., more preferred between 45° C. and 65° C. The reaction is triggered with the help of an initiator (also called starter in the following), which is formed with the help of a peroxide, e.g. hydrogen peroxide. For the formation of the starter, hydrogen peroxide can be directly dosed into the reactor. In another option, the starter can be synthesized in a separate unit. In another embodiment, it is preferred to form the starter by reaction of the recycle stream and hydrogen peroxide as described below in step D.fwdarw.A. [0031] The raw materials methane, SO3 and H2O2, and the solvent containing sulfuric acid and/or MSA are typically provided in the feed in a molar ratio of methane/SO3 between 0.50 to 2.0 mol/mol, preferred between 0.75 to 1.2 mol/mol. The ratio of H2O2/SO3 is typically in the range between 0.001 to 0.20 mol/mol, preferred between 0.01 to 0.10. The molar ratio between SO3 and MSA in the feed to the reactor system is typically set between 0.1 to 100 mol/mol, preferred 1 to 80 mol/mol, more preferred between 5.0 to 65.0 mol/mol. In a normal and stable mode of operation these ratios are kept constant by adequately dosing fresh feed components. The absolute amount of SO3 in the feed stream is between 20% and 85 wt %, preferably 40% to 75 wt %, more preferably between 60 and 75 wt %. [0032] The liquid phase in the reactors is usually transferred to the next process step B. The reaction mixture leaving step A consists mainly of MSA, H.sub.2SO.sub.4 and unconverted reactants (SO.sub.3 and methane). Apart from that, the solution leaving the reactors in most cases contains side products of the reaction, like methanesulfonic acid anhydride (MSA-anhydride), methanesulfonic methylester (Me-MSA), Methyl bisulfate (MBS) or methyldisulfonic acid (MDSA). [0033] B Quench: The remaining SO.sub.3 and MSA-anhydride can be converted with water to H.sub.2SO.sub.4 and MSA, respectively, either when the reaction mixture is still under CH.sub.4 pressure (above ambient pressure) or when the reaction solution is already decompressed (ambient pressure or below). If conversion of SO.sub.3 is close to 100%, e.g. between 95 and 100%, and the combined selectivity towards MSA and optionally MSAA is close to 100%, e.g. 90 to 100%, it is one option within the scope of the present invention that the reaction mixture might be processed further without step B. However, it is preferred to include step B in the process even if SO3 conversion and the selectivity to MSA and optionally MSAA are high. [0034] C Degassing/Decompression: By reducing the pressure in step C a light boiling stream consisting mainly of methane and a heavy stream consisting of mainly MSA and sulfuric acid is formed. [0035] D MSA purification: The feed to the purification contains mainly MSA and sulfuric acid. Additionally, water, unreacted raw material and reaction side products can be present. MSA can be recovered by distillation. Depending on the exact composition of the feed stream entering step D and the distillation concept, purified MSA is obtained from side-discharge (WO2018/219726) or from the bottom of a distillation column (unpublished patent application EP 19190621.3, i. e. BASF internal reference: 191048EP01). The bottom residue of such a distillation mainly consists of MSA and sulfuric acid. Typically, this stream is rich of sulfuric acid and still contains 10 to 30 wt.-% MSA. In general, all distillation columns are operated under reduced pressure in the range of from 0.1 to 50 mbar, preferably 2 to 30 mbar, more preferably 3 to 20 mbar and most preferably 5 to 15 mbar (all values given as absolute pressure at the heads of the columns and evaporators). Both bottom temperatures of the distillation columns and residence times in the bottom section of the distillation (i.e. the volume in the sump of the columns including heat exchangers, pumps, piping etc.) should be kept as low as possible to avoid decomposition of MSA. For this reason, the bottom temperatures are controlled to be typically between 140° C. and 220° C., preferably between 150 and 210° C., more preferably between 160° C. and 200° C. and most preferably between 165 and 195° C. or 170 and 190° C. [0036] D.fwdarw.A Solvent Recirculation: One fraction of the residue from the MSA purification can be recirculated to step A as solvent for the reaction. H.sub.2O.sub.2 can be added to this stream or to a part of this stream to obtain fresh initiator. In one mode of operation, this recirculated stream is cooled. In another mode of operation, H.sub.2O.sub.2 is added to the cooled recirculated stream or a part of it under intensive mixing and cooling. Moreover, SO.sub.3 can be added to this stream or to a part of this stream to load the reactor(s) with reactant and solvent, respectively. In another mode of operation, SO.sub.3 is added to the cooled recirculated stream or a part of it with H.sub.2O.sub.2 under intensive mixing and cooling. Alternatively, H.sub.2O.sub.2 and SO.sub.3 can be added simultaneously to the pre-cooled recirculated stream or to a part of it under intensive mixing and cooling. In another embodiment, SO.sub.3 and/or H.sub.2O.sub.2 can directly be added into one or several of the reaction vessels of Step A. [0037] E Waste stream: Sulfuric acid is formed in the process, especially in step B, and undesired accumulation of sulfuric acid in the recirculated stream must be prevented. Therefore, the other fraction of the purification residue needs to be removed from the process and e.g. be discarded as a waste stream (also called “purge”). For an economic operation the purge stream should be kept to a minimum. [0038] Thus, the inventive process may comprise, in a preferred embodiment, the formation of methanesulfonic acid (MSA) from methane and SO.sub.3 with a mixture of MSA and sulfuric acid as a solvent at a pressure from 10 to 200 bar and a temperature ranging from 25 to 100° C. (step A), preferably conversion of remaining SO.sub.3 and MSA-anhydride to H.sub.2SO.sub.4 and MSA with water (step B), reduction of pressure (step C), recovery of MSA by distillation (step D), recirculation of one fraction of the residue from the MSA purification to step A as solvent for the reaction (step D), and discarding of the other fraction of the purification residue as a waste stream (step E), wherein, preferably, the addition of the starter or starter precursor, as defined below, to the reactor or the reactor cascade is split into two or more portions.

[0039] In one embodiment of the invention a high MSA-yield referring to SO3 is achieved by dividing the addition of the starter or starter precursor to the reactor or the reactor cascade into two or more portions. When setting up the synthesis as batch synthesis the starter or starter precursor can be added to the same reactor in several portions rather than adding the total amount at the very beginning of the synthesis in only one portion. Alternatively, when the batch synthesis is realized in a reactor cascade, the starter or starter precursor can be added in portions to the first reactor and/or to further reactors of the cascade. When setting up the synthesis as continuous process the starter or starter precursor can be divided and added continuously to the first reactor and to further reactors of the cascade at the same time. A continuous synthesis set-up with the amount of starter being divided and added continuously to the first and at least one more reactor of the cascade at the same time is preferred. The starter is typically provided in a solvent. It is preferred that the solvent contains sulfuric acid and/or methanesulfonic acid.

[0040] In one embodiment the division of the starter (total: x mol) over the n reactors of the reactor cascade can be varied in a wide window. For example, it can be continuously divided over the n reactors in a homogenous manner, meaning that each of the n reactors is supplied with a continuous fraction of x/n mol of the starter. Alternatively, the largest amount can be fed continuously into the first reactor of the cascade, e.g. 50% or 60% or 70% or 80% or even 90% and more of the x mol starter, while the remaining starter is divided over the remaining reactors of the cascade, either evenly or in different ratios per reactor. In another embodiment the largest amount of the starter is added to the first reactor while all the rest is fed into the second reactor of the cascade. In another embodiment the largest amount of the starter is added to the first reactor while all the rest is fed into the next reactors of the cascade.

[0041] While all these options are feasible and help to improve SO.sub.3 conversion to MSA, it is preferred to add a significant part of the starter to the first reactor, e.g. between 30% and 98%, and all the remaining starter to the last reactor to push SO.sub.3 conversion where SO.sub.3 concentration is low and reaction rates therefore decrease. In another preferred mode the total amount of starter x is divided evenly between the n reactors.

[0042] In one preferred embodiment the division of the starter (total: x mol) over the n reactors of the reactor cascade is divided only in the first n−1 reactors, while no starter is added into the last reactor of the cascade. This has the advantage, that the total amount of starter at the outlet of reactor n is further reduced and no or less decomposition of unreacted starter is expected to take place in the following reaction and work up sections. As decomposition of starter might cause gas formation, this mode of operation with prevention of further starter decomposition is beneficial for the following process steps.

[0043] By dividing the starter dosage between the n reactors the effectivity of the starter is optimized, translating into higher SO3 conversion and finally MSA synthesis. This also means that the total amount of starter x, which is required when dividing the starter dosage as described above, is the same or even lower compared to a set-up where the total starter volume y is added to the first reactor, i.e. n*x/n=x=y or n*x/n=x<y.

[0044] Our investigations have furthermore shown that the starter and the precursor of the starter—be it a single compound or a mixture of different compound in solution, e.g. in sulfuric acid or MSA/sulfuric acid—may start decomposing at temperatures as applied during the reaction of methane and SO.sub.3, typically e.g. 50° C. Therefore, the cooling steps described in the following help to suppress decomposition of the starter or starter precursor, thus foster the selective conversion of SO.sub.3 to MSA and finally result in higher MSA yields.

[0045] Therefore, in a preferred embodiment of the invention a high MSA-yield referring to SO.sub.3 is achieved by cooling the recycle stream leaving the distillation as a bottom residue prior to mixing it with the starter solution or the starter precursor solution or prior to feeding it to the reactor.

[0046] The starter precursor can be selected from sulfuric acid, oleum, SO3, methanesulfonic acid, the bottom recycle stream from MSA distillation and mixtures thereof. Further starter precursors may be inorganic oxoacids or salts thereof, e.g. oxoacids of boron, silicon, phosphorus, carbon, nitrogen or sulfur. Yet another group of precursors can be selected from organic oxoacids. The compounds can be used in pure form or in a solvent, especially a solvent comprising sulfuric acid and/or oleum and/or methanesulfonic acid. The starter itself is generated from this starter precursor or starter precursor solution by addition of hydrogen peroxide under intense mixing and cooling.

[0047] The starters may be selected from inorganic or organic peroxoacids or salts thereof, as e.g. a peroxosulfuric acid comprising at least one peroxosulfuric acid of boron, silicon, phosphorus, carbon, nitrogen or sulfur. Consequently, typical starters may be Caro's acid, Marshall's acid, dimethylsulfonylperoxide, monomethylsulfonylperoxide etc.

[0048] The recycle stream leaving the distillation as a bottom residue typically has a temperature between 140° C. and 220° C. The reactor cascade is typically operated between 25° C. to 100° C. In one embodiment of the invention the recycle is added to the reactors without further cooling.

[0049] In one option, the recycle stream is cooled to temperatures <25° C., preferred <20° C., more preferred <15° C., and most preferred <10° C. or even <5° C. or <0° C. and then the complete recycle stream is used for the preparation of the starter.

[0050] In another embodiment of the invention, preferably not the whole recirculated stream is used for the preparation of the starter, but a fraction of it, e.g. less than 50% of the recirculated stream, preferred less than 30% of the recirculated stream, more preferred less than 20% of the recirculated stream. In this case, the complete recycle stream is first cooled to the reaction temperature, e.g. 50° C., or it can be cooled to a temperature range close to the reaction temperature, e.g. 50° C. above or below the reaction temperature, preferably 20° C. above or below the reaction temperature, more preferably 5° C. above or below the reaction temperature and most preferably to 1° C. above or below the reaction temperature. It is preferred to cool down the complete recycle stream to reaction temperature or slightly below reaction temperature. Then, the fraction of the recycle stream, that is not used for preparation of the initiator is directly dosed to the reaction in step A as solvent. The fraction used for the preparation of the initiator is further cooled to temperatures <50° C., preferred <25° C., and most preferred <10° C. or even <5° C. or <0° C. or <−5° C.

[0051] In another embodiment of the invention a high MSA-yield referring to SO.sub.3 is achieved by cooling the starter solution or starter precursor solution while mixing it with the recycle stream rather than adding the starter or starter precursor into the reactor and cooling the whole reaction mixture. Mixing of the recycle stream leaving the distillation as a bottom residue, preferred pre-cooled as described above, with hydrogen peroxide, the starter solution or the starter precursor solution, is done under intense mixing and cooling. Mixing can be done in a pipeline, in a static mixer, in a stirred or a not stirred reactor, in a vessel with injector nozzle or jet nozzle, in a heat exchanger or other devices suitable for mixing. In a preferred mode the mixing of the two streams is carried out in a static mixture with cooling function included or in a static mixer where a heat exchanger immediately follows the static mixer.

[0052] Preferably, the starter solution or starter precursor solution itself is cooled in a first step, and then it is mixed under further cooling with the pre-cooled recycle stream leaving the distillation as a bottom residue. These process steps are preferably performed at temperatures <50° C., preferred <25° C., and most preferred <10° C. or even <5° C. or <0° C. or <−5° C.

[0053] Of all the measures to increase SO.sub.3 conversion to MSA given above, executing just one already has a beneficial effect and increases SO.sub.3 conversion and thus also MSA yield. But as a matter of fact, executing several of the measures in combination will improve SO.sub.3 conversion to MSA even more, and ideally all of the measures given above are realized in the process set-up as this will have the maximum effect to increase SO3 conversion translating into increased MSA yield.

[0054] This invention is, in a preferred embodiment, related to the production of methanesulfonic acid form methane and SO3. However, the aspects of this invention can also be applied for the sulfonation of other alkanes than methane to produce the respective alkanesulfonic acid.

[0055] One further target of the present invention is the reduction of costs for waste disposal. Therefore, the amount of waste generated in step E must be reduced. As described above, the size of this stream is related to the formation of sulfuric acid in steps A and B. The amount of sulfuric acid generated in step A is determined mainly by the amount of starter solution, more specific by the amount of water introduced in step A when aqueous H.sub.2O.sub.2 is added to produce the starter or the starter precursor solution. Free water introduced via H.sub.2O.sub.2 solution reacts with SO.sub.3, reduces the MSA yield referring to SO.sub.3 and increases the amount of sulfuric acid which accumulates in the process and needs to be purged. The amount of sulfuric acid generated in step B is determined mainly by the amount of unconverted SO.sub.3 leaving reaction in step A. Consequently, to reduce waste cost, a high SO.sub.3 conversion to MSA in Steps A and B is required.

[0056] In the prior art, the formation of MSA in a cascade of reactors is proposed and all reactants including the initiator are added to the first reactor. The conversion of SO3 to MSA in the reactor(s) is determined inter alia—as one out of several parameters—by the concentration of methane in the synthesis mixture. One option is to add methane just to the first reactor. In a continuously operated cascade of reactors this implies that the pressure slightly decreases from one reactor to the following. Especially if the starting pressure is high as described above, e.g. 100 bar, the pressure remains on a fairly high level all through the reactors. Typical pressure drops may be e.g. up to 5 bar per reactor, preferably up to 2 bar per reactor, more preferably up to 1 bar per reactor or even less, e.g. about 0.5 bar or 0.1 bar per reactor of the cascade. In one embodiment of the invention, methane is not added just to the first reactor but also to further reactors or reaction chambers of the cascade.

[0057] In one embodiment of this invention, the methane pressure is adjusted reactor by reactor, e.g. by continuously increasing the methane pressure from the first to the last reactor in a reactor cascade via individual methane supply pipelines for each reactor. It needs to be mentioned that, according to a preferred embodiment of this invention, in a continuous and stable mode of operation of the inventive process the one pressure level or the different pressure levels of the reactors of a reactor cascade are always kept constant, e.g. by using pressure control valves in the methane feed lines to the reactor(s). Alternatively, methane can be fed at the same pressure into the first reactor(s) of a cascade, whereas the last reactor is constantly operated at a higher methane pressure than the first reactor(s) of the cascade, e.g. higher by up to 5 bar or up to 10 bar or even up to 20 bar. All reactors operated at the same pressure can be connected via pipeline at the gas phase of each reactor to ensure supply of methane to each reactor. Alternatively, the reactors can be equipped with an overflow pipe. In this kind of setup, a gas/liquid mixture containing methane is transported from one reactor to another. Yet another option is to reduce the methane pressure from the first to the last reactor in a reactor cascade via individual methane supply pipelines with pressure control for each reactor. It is preferred to keep the pressure over of the cascade constant, i.e. the pressure difference between first and last reactor is no more than 5 bar, preferably no more than 2 bar, more preferably no more than 1 bar or even less, e.g. no more than 0.5 bar or 0.1 bar.

[0058] To enable a high mass transfer rate of methane from the gas to the liquid phase, a high surface area at the gas-liquid boundary must be provided. This can either be achieved by use of a stirrer. Alternatively, one or more dip tubes, nozzles, ejector nozzles, frits, sieve trays, or ring spargers can be used as gas distributor. As reactors one or more stirred tank reactors, bubble column reactors, gas circulation reactors, air lift reactors, jet loop reactors, falling film reactors, tubular reactors and water-ring pumps can be used. Stirrers, nozzles, and gas distributors can also be used in combination. Another option to provide a high mass transfer rate is the use of filling material or reactor internals with a high surface area like glass balls, random column packings (e.g. Raschig rings) or structured packings (e.g. packings provided by Sulzer, Montz, Koch-Glitsch or others). The alkane, preferably methane, is introduced into the reactor at a specific height, which can be above the surface of the liquid or underneath the surface of the liquid. If the gas is introduced underneath the surface this can be done in different manners, e.g. by dip tube, by distributor, by nozzle, by jet nozzle, by perforated plate or other means suitable to introduce gas into a liquid, or a combination thereof. It is preferred to introduce the alkane, e.g. methane, underneath the surface of the liquid.

[0059] In one embodiment of the reaction the reactor cascade can comprise different types of reactors, e.g. a combination of a loop reactor or set of loop reactors and a stirred tank reactor or set of stirred tank reactors. The reactors can be aligned as series, in parallel or in a combination of parallel and serially aligned reactors. It is beneficial to provide a high gas-to-liquid mass transfer rate at the beginning of the reactor cascade, when the rate of the reaction is high. After the first reactor(s) providing a high gas-to-liquid mass transfer rate, it may be sufficient that further reactors provide a lower mass transfer rate. Technically, this can be realized by combining e.g. stirred tank reactor(s) with bubble column reactor(s) or jet loop reactor(s) with an air lift reactor(s). Irrespective of the combination of different types of reactors the addition of the initiator can be split over two or more reactors of the cascade of the reactors.

[0060] Dosing all SO3 into the first reactor of a cascade, directly or after pre-mixing, total or partial, e.g. with the recycle stream leaving the distillation as a bottom residue, is preferred. However, there may be reasons which also dictate to divide the dosing of the raw material SO3 into several reactors of the cascade rather than adding the total amount to the first reactor, irrespective if the reaction is set up as batch process or as continuous process. Such reasons may be a limited solubility of SO3 in the recycle stream, or the need to reduce the heat of mixing when adding SO3 to the recirculated stream, or vice versa, per piece of equipment. As for the starter and for methane, the addition of SO3 may then be divided between the first reactor(s) of the cascade. SO3 should not be added to the last reactor of the cascade as chances to convert the largest amount of SO3 before it leaves the reactor towards quench and distillation are minimal.

[0061] If the initiators are only added to the first reactor as described in the prior art, SO.sub.3-conversion and consequently MSA yield achieved over the reactor cascade is lower than in the inventive process when the dosing of the initiator is split. An example of a reactor cascade with the supply of reactants and initiator according to one embodiment of the present invention is depicted in FIG. 2 (and described below).

[0062] Especially in a continuous process set-up the amount of water required in the quench (step B) to convert SO3 to sulfuric acid and to hydrolyse MSAA to MSA, can be calculated based e.g. on the mass flow leaving the reactor or entering the quench assuming a constant percentage of SO3 and MSAA to be treated. Alternatively, the amount of water can be calculated relative to the mass flow of raw materials, e.g. SO3 or methane or the recycle stream, being added to the reactor cascade. In this case water is added in step B with a constant mass flow in a fixed ratio to the mass flow of e.g. one of the raw materials or the product stream leaving the reactor cascade as listed above. It is easy for a person skilled in the art to establish a similar control based on fixed ratios for a process operated in batch mode, e.g. adding water in a fixed ratio to one of the raw materials in combination with the reaction time.

[0063] However, in a preferred mode to better control the feed to the distillation and the discharge of a waste stream in step E according to the conversion of SO3 in steps A and B, accurate process analytics is helpful. The generation of MSA and the conversion of SO3 can be measured with a density measurement, e.g. in step B. As the density of pure MSA is lower than the density of Oleum (H2SO4+SO3), an increasing conversion can be monitored by a decreasing density. In step B, the addition of water can be controlled with the help of a conductivity measurement. The conductivity rises proportional to the water content of the solution. To avoid the formation of side products in step D, it is crucial to remove SO3 completely. For this reason, equimolar amounts of water relative to SO3 and MSAA or an excess of water is beneficial. However, if too much water is added the purification effort in step D would increase. Furthermore, the ratio of MSA to sulfuric acid can be checked with a density measurement in the feed to step D. As an alternative to conductivity measurements, sound velocity measurements can be used. Side products of the reaction, e.g. methanesulfonic acid anhydride (MSAA), can be identified using an online spectroscopic method like RAMAN-, IR- or NMR-spectroscopy. As MSAA can be hydrolyzed in Step B, the information gathered with spectroscopic methods can also be used to adjust the water addition in this step. Preferably, RAMAN-spectroscopy is used. As an alternative, chromatographic methods can be employed. When applying the suitable analytic method the mass flow rate of the purge stream in step E can be adjusted according to the amount of unconverted SO3 and the amount of sulfuric acid formed during the synthesis of the initiator and consequently can be minimized.

[0064] All analytic methods can be applied as offline measurement, e.g. for measurements on a daily, weekly or monthly basis, or as inline measurement for continuous surveillance of the respective parameters

[0065] A high conversion of SO.sub.3 in step A or in other words a reduced formation of sulfuric acid further helps to minimize the concentration of sulfuric acid in the distillation in step D.

[0066] It is advantageous, to feed MSA with a concentration between 50 to 99 wt.-%, preferred 60 to 90 wt.-%, more preferred 65 to 85 wt.-%, into the distillation to keep the distillation set-up simple and the energy demand to operate the distillation low. This implies that the feed to the reaction has a SO3 content of 30 to 99 wt.-%, preferred 50 to 99 wt.-%, more preferred 60 to 90 wt.-%, and most preferred 60 to 80 wt.-%.

[0067] For the general understanding of this invention it needs to be mentioned that according to the inventive process the process parameters at stable operation, be it in continuous or in batch mode of operation, especially in continuous mode of operation, are preferably kept constant, i.e. at or close to a target value, e.g. regarding temperature level(s), pressure level(s), feed ratios, liquid level(s) etc., as it is common practice in chemical processes, e.g. by applying temperature control systems, pressure control systems, mass flow and/or volume flow systems, these systems normally consisting of a sensor or a set of sensors and the respective actor or set of actors.

[0068] To illustrate some aspects of the present invention, several experimental examples may be found below.

EXAMPLES

Comparative Example 1

[0069] The synthesis of MSA from SO3 and methane (process step A) is carried out in a cascade of three reactors in continuous operation. Each of these stainless steel reactors has a volume of 181 and is equipped with stirrer, thermo element, pressure indicator, cooling coils and a blow-out disc. Prior to use the reactors are inerted by rinsing with nitrogen for 5 min. For the start-up of the reactors, the first reactor is filled with a mixture of H2SO4 and MSA (mass ratio 12.5:1) until the stirrer is covered and the temperature is set to 50° C. Methane is provided to each reactor at a pressure of 90 bar and the feed supply is started. The reactors of the cascade are successively filled with the fresh reaction mixture until a stationary operation point is reached. The initiator is produced by adding H.sub.2O.sub.2 60% to the MSA/H2SO4-mixture recycled from the bottom of the distillation under constant cooling at −5° C. The amount of initiator solution added constantly to the reaction mixture is such that the H.sub.2O.sub.2 concentration in the whole feed mixture is about 1 wt.-%. The total liquid feed to the first reactor (SO.sub.3, MSA, H.sub.2SO.sub.4, H.sub.2O.sub.2) has a mass flow rate of about 2.5 kg/h with a concentration of fresh SO.sub.3 of 65 wt.-%, the rest being a mixture of MSA (3 wt.-%), sulfuric acid (31 wt.-%), initiator solution and small amounts of side products from the recycle of the bottom of the distillation column (<<1 wt.-%, e.g. MDSA). When the reactors are at a stationary operation point, the concentrations of the main components in the reactor outlets are as follows:

TABLE-US-00001 Concentration/wt.-% Reactor 1 Reactor 2 Reactor 3 MSA 45 52 57 SO.sub.3 18 11 8 H.sub.2SO.sub.4 37 37 35

[0070] Samples from each reactor are taken and analyzed by 1H NMR, ion chromatography and titration. The MSA-concentration after the last reactor is determined by density measurement to be 57 wt.-%. The SO3 content cannot be determined by direct measurement. It was calculated by a component balance considering the compositions of the reactor feed and the reactor outlet and the MSA concentrations in each reactor. The MSA concentration comprises MSA and MSA anhydride. MSA anhydride is completely hydrolyzed to MSA in the quench step. A SO.sub.3 conversion of 86% and a yield with regard to SO3 of 74% were reached.

[0071] If these results are scaled to a production process with a capacity of 10 kt/a MSA 100% (8000 operating hours per year) the purge stream in process step E has a mass flow rate of 880 kg/h under the assumption that the purge stream still contains 30 wt.-% MSA.

Example 1

[0072] The reactor cascade of example 1 is prepared and operated in the same manner as in comparative example 1, i.e. at the same temperature, pressure and with the same feed rate and feed composition as in comparative example 1. In contrast to comparative example 1, the initiator is not only added to the first reactor, but also to the second reactor of the cascade. The split between initiator added to reactor one and two of the cascade is 2 (reactor 1): 1 (reactor 2). When the reactors are at a stationary operation point, the concentrations of the main components in the reactor outlets are as follows:

TABLE-US-00002 Concentration/wt.-% Reactor 1 Reactor 2 Reactor 3 MSA 45 59 62 SO.sub.3 18 5 3 H.sub.2SO.sub.4 37 36 35

[0073] Samples from each reactor are taken for analysis (1H NMR, ion chromatography and titration) and the MSA-concentration after the last reactor is determined by density measurement to be 62 wt.-%. The SO3 content cannot be determined by direct measurement. It was calculated by a component balance considering the compositions of the reactor feed and the reactor outlet and the MSA concentrations in each reactor. The MSA concentration comprises MSA and MSA anhydride. MSA anhydride is completely hydrolyzed to MSA in the quench step. Due to the split of the starter a SO.sub.3 conversion of 95% and yield with regard to SO.sub.3 of 85% were reached.

[0074] If these results are scaled to a production process with a capacity of 10 kt/a MSA 100% (8000 operating hours per year) the purge stream in process step E has a mass flow rate of 596 kg/h under the assumption that the purge stream still contains 30 wt.-% MSA.

[0075] This example highlights that the measures claimed in this application lead to a massive reduction of waste.

Example 2

[0076] Hydrogen peroxide was added to Oleum 32 under cooling to 0° C. to yield an initiator solution with a concentration of 0.7 wt.-% hydrogen peroxide. The peroxide concentration was confirmed by titration. The degradation of the peroxide at several temperatures is compared with the time t.sub.1/2, in which half of the peroxide was decomposed. The results are shown in Table 1. The decomposition rate of the initiator shows a strong dependency on the temperature and quickly increases with increasing temperature. This example illustrates the advantage of cooling the recirculated stream from step D to step A for the formation of the initiator to ensure, that the required amount of initiator reaches the reactor.

TABLE-US-00003 TABLE 1 Decomposition of the initiator at different temperatures characterized by the time t.sub.1/2, in which half the peroxide was decomposed. T/° C. t.sub.1/2/min 40 1497 50 216 60 58

[0077] These results are surprising since industrially available peroxides, for example hydrogen peroxide, are expected to be stable also at elevated temperatures, for example at 50° C. Technical Data Sheet of hydrogen peroxide Interox® Technical Grade 50%, downloaded from Solvay homepage in July 2020, mentions a stability of 98%. This stability is based on gasometric stability measurements of hydrogen peroxide by Solvay performed at 100° C. (taken from the Technical Data Sheet downloaded from Solvay homepage in July 2020).

[0078] The experimental examples in general show a higher yield of the inventive process. Furthermore, the amount of waste stream could be reduced. These technical advantages could be obtained without any further reactor equipment, but only with an optimized reaction management.