PROCESS FOR MANUFACTURING ALKANESULFONIC ACIDS

Abstract

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

Claims

1. A process for the production of alkanesulfonic acids from an alkane and SO3 with an initiator comprising a distillation step, wherein total energy consumption of the process is reduced by heat integration, wherein feed to the distillation step is heated by heat integration with the hot stream of purified alkanesulfonic acid, and/or feed to the distillation is heated by heat integration with the hot distillation residue, and/or feed to the distillation is first heated by heat integration with the hot purified alkanesulfonic acid and then by heat integration with the hot distillation residue, and/or feed to the distillation is heated by heat integration with the hot stream from mixing the alkanesulfonic acid with water to yield an aqueous solution.

2. The process according to claim 1, comprising the following steps a) synthesizing an initiator and reacting an alkane with SO3 to form alkanesulfonic acid, b) quenching to convert SO3 into a heavy boiler, c) removing unreacted alkane by decompression, d) purifying alkanesulfonic acid by distillation, d*) pre-heating of feed to distillation by heat integration, and e) discarding waste streams.

3. The process according to claim 1, wherein the alkane is methane and the alkanesulfonic acid is methanesulfonic acid.

4. The process according to claim 1, wherein the feed to the distillation is further heated to the distillation temperature at the feeding point after the heat integration.

5. The process according to claim 1, wherein heat exchangers are used for heat integration.

6. The process according to claim 5, wherein the heat exchangers are monoliths made of pressureless sintered silicon carbide (SSiC).

7. The process according to claim 1, wherein the heat in the heat integration step is transferred from the hot streams leaving the distillation to the cold feed directly without a heat transfer fluid in one apparatus per hot stream.

8. The process according to claim 1, wherein the heat in the heat integration step is transferred from the hot streams leaving the distillation to the cold feed with the help of a heat transfer fluid.

9. The process according to claim 1, wherein aliphatic and/or cyclic hydrocarbons and/or as silicon oils are used as heat transfer fluids for heat integration.

10. The process according to claim 2, wherein the quenching step is operated adiabatically.

11. The process according to claim 1, wherein remaining alkane from the synthesis of alkanesulfonic acids is incinerated to supply heat to the process.

12. The process according to claim 1, wherein the heat supplied by the incineration of alkane can be used to generate steam or to heat up heat transfer fluids.

13. The process according to claim 1, wherein the alkane is methane and the heat supplied by the incineration of methane can be used to heat up the feed to the distillation or to power evaporators of the distillation.

14. The process according to claim 13, wherein the evaporators of the distillation are powered by steam and steam condensate coming from the evaporators of the distillation is used to heat the feed to the distillation.

15. The process according to claim 1, wherein the energy released in condensers of the distillation step is used to heat up the feed to the distillation.

16. The process according to claim 5, wherein the heat exchangers are made of silicon carbide.

Description

DETAILED DESCRIPTION OF THE INVENTIVE PROCESS

[0051] This invention relates to an optimized process for the production of alkanesulfonic acids. In the following a detailed exemplary description of the process according to one embodiment of the present invention is given for the case that the alkane is methane and methanesulfonic acid (MSA) is formed in the reaction:

[0052] The schematic flow chart of one embodiment of the present invention is shown in FIG. 1. A further embodiment of the inventive process is shown in FIG. 3.

[0053] FIG. 2 depicts an exemplary process for the production of MSA from methane and SO3 without heat integration.

[0054] MSA can be produced from sulfur trioxide and methane in the following steps: [0055] 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 [0056] 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 25 to 100° C., preferred 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”), which is prepared with H2O2. The peroxide initiator is prone to decomposition at elevated temperatures, therefore—in addition to controlling the reaction temperature within the limits given above—cooling is preferably applied when producing the initiator precursor solution and/or during preparation of the initiator solution. Cooling can be applied to one or more of the individual raw material streams and/or to the mixing process itself when generating the initiator precursor solution and/or the initiator solution from of the initiator raw materials, comprising e.g. a solution of hydrogen peroxide, SO3, oleum, the recycle stream leaving the distillation step at the bottom of the column consisting mainly of sulfuric acid and the respective alkanesulfonic acid, when using methane as alkane methanesulfonic acid, and small amounts of other high boilers, e.g. methyldisulfonic acid. Preferably, the initiator solution or initiator 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. [0057] For the cooling of the synthesis of the starter/initiator and MSA the following cooling media can be used: water, aqueous salt solutions, organic liquids like aliphatic or aromatic oils, poly glycols etc. [0058] The liquid phase in the reactors is 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 (e.g. SO.sub.3 and methane). Apart from that the solution leaving the reactors in most cases contain side products of the reaction like methanesulfonic acid anhydride (MSA-anhydride), methanesulfonic methylester (Me-MSA), Methyl bisulfate (MBS) or methyldisulfonic acid (MDSA). [0059] Unconverted methane from the synthesis of MSA can be incinerated to supply heat for other process steps. By incinerating methane, steam can be generated. Steam can be used as a heat source in process step D and/or in process step D*. Additionally (or alternatively), by incinerating methane heat transfer fluids can be heated to supply heat especially in process step D and process step D*. [0060] 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 the 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%. [0061] 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. the reaction of SO3 with water to sulfuric is highly exothermic (ΔH.sup.R=−132 kJ/kg). Depending on the composition of the reaction mixture, the quench step might be operated isothermal or adiabatic. In an adiabatic operation of the quench step, the released heat of reaction leads to a temperature rise in the quench. Consequently, the temperature gap to the distillation step (see below, step D) is reduced. [0062] 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. Process steps B and C can be combined in one apparatus.

[0063] Unconverted methane from the synthesis of MSA can be incinerated to supply heat for other process steps. By incinerating methane, steam can be generated. Steam can be used as a heat source in process step D and/or in process step D*. Additionally (or alternatively), by incinerating methane a hot gas stream us generated which can be used as heat source directly, heat transfer fluids can be heated with steam or directly with hot gas from incineration of methane to supply heat in process step D and process step D*. However, as will be highlighted in Comparative Example 3, the power required for the MSA purification exceeds the power, that can be supplied by burning methane coming from step C. [0064] D MSA purification: The feed from step C can be sent directly to the distillation without further heat treatment. However, it is preferred that the feed to the distillation is subject to a pre-heating step to reduce energy consumption for the overall process by heat integration as described below under step D*. The feed to the distillation can enter the column at a temperature above or below the boiling point. It is preferred, that the feed entering the column is a boiling liquid or as an overheated liquid (flash). [0065] The feed to the purification contains mainly MSA and sulfuric acid. Additionally, water 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, BASF reference: 191048EP01). The distillation step may consist of one column or a set of several columns with different functionality. Depending on the capacity and the design of the column or set of columns, each column may be realized as one column or as a set of columns in parallel having the same functionality. 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. The feed to the distillation column is preheated to the distillation temperature of 150° C. till 200° C. as described in step D*. Regardless of the distillation concept, purified MSA leaves the distillation step D with a temperature of 150° C. till 200° C. The bottom residue of such a distillation mainly consists of MSA and sulfuric acid. Regardless of the distillation concept this stream has a temperature of 150° C. till 220° C. Typically, the temperature of the purified MSA leaves the distillation column at a lower temperature than the bottom residue. [0066] D* Heat integration: To heat up the feed to the distillation (step D), the heat streams leaving the distillation may be used as described earlier. This leads to a massive reduction of the energy demand of the process. The heat streams leaving the distillation are the streams with purified MSA and the bottom stream leaving the distillation mainly consisting of MSA and sulfuric acid. Only one stream or both streams can used to heat up the feed to the distillation. Preferably, both streams are used. More preferably, first the heat is first transferred from the hot purified MSA and then from hot bottom residue to the solvent. Further sources of heat streams can be used to further optimize, i.e. reduce energy consumption of the overall process, e.g. the hot stream generated when mixing MSA with water to a target concentration (formulation of MSA, step F), or the heat stream generated when incinerating methane as outlined above in step C. As a heat exchanger device, heat exchangers made of a material selected, separately from each other, from the group consisting of silicon carbide (SiC, SSiC), Titanium, Tantalum, Gold, Hastelloy or glass-lined steel are preferred. Preferably, heat exchangers made of sintered silicon carbide (SSiC) are used. More preferably the heat exchangers contain one or several SSiC monoliths or SiC tubes. The SiC monolith are tubes can be enclosed in a shell made of stainless steal, PTFE/PFA lined steel or glass-lined steel. [0067] Devices of this kind allow a direct heat transfer from the hot stream to the cold stream, either using the hot and the cold stream directly or with a heat transfer fluid. Using the hot and the cold stream directly, i.e. without a heat transfer fluid, is preferred. Alternatively, other types of heat exchangers can be used, e.g. heat exchangers containing glass-lined steel tubes. Usually these types of heat exchangers require a heat transfer medium which for example is based upon aliphatic or cyclic hydrocarbons or silicon oil. Further potential materials of construction comprise gold, tantalum, titanium, or Hastelloy. [0068] The feed to the distillation cannot be heated up completely to the temperature of the feed tray by heat integration as long as the methane set free in step C is not considered. Preferably, the feed is further heated to the feed temperature at the feeding point in the distillation column by a heat exchanger prior to the distillation column powered with steam. Alternatively, this heat exchanger can be powered electrically or with a heat transfer fluid, e.g. aliphatic or cyclic hydrocarbons or silicon oil. As soon as incinerating methane from step C is considered for the energy balance of the overall process, sufficient heat is provided process-internally to heat up the feed to the distillation up to the temperature level of the feed tray, e.g. heat in the form of steam. Providing energy process-internally by incinerating methane, i.e. without using additional external energy for the process, is preferred. [0069] While heating up the feed stream to the distillation step, the hot streams leaving the distillation are cooled. This is beneficial, because the purified MSA needs to be cooled for formulation and packing. The bottom residue of the distillation needs to be cooled for further processing as described below. Eventually, the streams leaving the distillation cannot be cooled completely to the desired temperature in step D*. Further cooling can be achieved with for example, air cooling, cooling water or other cool fluids available in a chemical production plant. Especially when cooling water is applied it may contain additives like anticorrosion additives and/or anti-fouling additives and/or bioactive compounds keeping number and amount of organisms on a low level or even reducing them to zero. [0070] Moreover, it is possible to extend the heat integration of step D* to process step B or process step C or both steps B and C. [0071] 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. [0072] 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.

EXAMPLES

[0073] In the following paragraphs, several examples are described to illustrate some aspects of the present invention.

[0074] In all examples the energy demand refers to 1 kg of purified MSA. For the heat transfer a temperature difference of 10 K was assumed. The heat capacity of all streams is 1.41 kJ/(kg K).

Comparative Example 1

[0075] In a process for the production of MSA from methane and SO3 without heat integration (FIG. 2), the feed stream to the distillation is heated in a heat exchanger prior to the distillation step D. In a typical operation, the stream leaving step C has a temperature of about 50° C. and the feed tray of the first distillation column has a temperature of about 170° C. To produce 2000 kg/h of purified MSA, the feed to the distillation column has a flow rate of 3100 kg/h and consists of 75 wt.-% MSA and 25 wt.-% sulfuric acid with side products <<1 wt.-%. The energy required to heat up this stream from 50 to 160° C. is 241 kJ/kg(MSA).

[0076] At the same time, the hot purified MSA with a flow rate of 2000 kg/h and the same heat capacity has to be cooled down from 160° C. to 30 ° C. for formulation and packing. Additionally, the bottom stream leaving the distillation has flow rate of 1111 kg/h. It leaves the distillation column with a temperature of 185° C. and must also be cooled to a temperature of 30° C. for further processing. These cooling operations together require a cooling duty of 305 kJ/kg(MSA).

Comparative Example 2

[0077] In a process for the production of MSA from methane and SO3 without heat integration (FIG. 2) the feed stream to the distillation is heated directly in the distillation column. In a typical operation, the stream leaving step C has a temperature of about 50° C. and the feed tray of the first distillation column has a temperature of about 170° C. To produce 2000 kg/h of purified MSA, the feed to the distillation column has a flow rate of 3100 kg/h and consists of 75 wt.-% MSA and 25 wt.-% sulfuric acid with side products <<1 wt.-%. The energy required to heat up this stream from 50 to 160° C. is 241 kJ/kg(MSA).

[0078] At the same time, the hot purified MSA with a flow rate of 2000 kg/h and the same heat capacity has to be cooled down from 160° C. to 30° C. for formulation and packing. Additionally, the bottom stream leaving the distillation has flow rate of 1111 kg/h. It leaves the distillation column with a temperature of 185° C. and must also be cooled to a temperature of 30° C. for further processing. These cooling operations together require a cooling duty of 305 kJ/kg(MSA).

Comparative Example 3

[0079] In order to produce 2000 kg/h of purified MSA, MSA is synthesized from SO3 and methane in step A at 100 bar and 50 ° C. Then, the liquid reactor outlet is passed from step A to step B and the rest of SO3 is quenched at the pressure of step A. Afterwards about 3100 kg/h of liquid are transferred to step C and are depressurized. At the reaction conditions of 100 bar and 50° C. the solubility of methane is 0.003 kg/kg. When the liquid is depressurized to atm. pressure in step C, 9.3 kg/h of methane are released. If this amount of methane is burned (assumptions: 100% efficiency, ca. 800 kJ/mole) about 130 kW of heat can be generated. To power the purification of 2000 kg/h MSA in step D at least 1000 kW are required. Therefore, additional measures are needed, the reduce the energy of the process.

Example 1

[0080] In an example according to the inventive process for the production of MSA from methane and SO3 with heat integration, the feed to the distillation is first heated with the hot purified MSA and then with the hot bottom residue of the distillation (cf. schematic illustration in FIG. 1). The stream leaving step C has a temperature of about 50° C. and the feed tray of the first distillation column has a temperature of 170° C. To produce 2000 kg/h of purified MSA, the feed to the distillation column has a flow rate of 3100 kg/h and consists of 75 wt.-% MSA and 25 wt.-% sulfuric acid with side products <<1 wt.-%. Purified MSA leaves the distillation step with a temperature of 160° C. and a flow rate of 2000 kg/h. The bottom residue of the distillation has a temperature of 185° C. and a flow rate of 1111 kg/h.

[0081] First, the feed stream to the distillation is heated in a heat exchanger made from SSiC with the help the hot purified MSA from 50° C. to 88° C., while at the same time the purified MSA is cooled to 98° C. Then, the feed is further heated in a heat exchanger made from SSiC with the help of hot bottom residue to 112° C., while the bottom residue is cooled to 122° C. Afterwards the feed is heated to 170° C. with steam in a heat exchanger made from SSiC. For this 127 kJ/kg(MSA) are required. Finally, the purified MSA and the bottom residue of the distillation are cooled to 30° C. with cooling water using a heat exchanger made from glass-lined steel. These cooling operations require a cooling duty of 168 kJ/kg(MSA).

[0082] Compared to a process without heat integration (Comparative Example 1), in the inventive process the energy for heating up the feed to the distillation is reduced by 52% and the energy demand for cooling the hot streams leaving the distillation is reduced by 45%.

Example 2

[0083] In an example of the optimized process according to the present invention for the production of MSA from methane and SO3 with heat integration the feed to the distillation is first heated with the hot bottom residue and then with the hot purified MSA (see schematic illustration in FIG. 3). The stream leaving step C has a temperature of about 50° C. and the feed tray of the first distillation column has a temperature of 170° C. To produce 2000kg/h of purified MSA, the feed to the distillation column has a flow rate of 3100 kg/h and consists of 75 wt.-% MSA and 25 wt.-% sulfuric acid with side products <<1 wt.-%. Purified MSA leaves the distillation step with a temperature of 160° C. and a flow rate of 2000 kg/h. The bottom residue of the distillation has a temperature of 185° C. and a flow rate of 1111 kg/h.

[0084] First, the feed stream to the distillation is heated with the help the bottom residue from 50° C. to 81° C., while at the same time the bottom residue is cooled to 91° C. Then, the feed is further heated with the help of hot purified MSA to 107° C., while MSA is cooled to 117° C. Afterwards the feed is heated to 170° C. For this 139 kJ/kg(MSA) are required. Finally, the purified MSA and the bottom residue of the distillation are cooled to 30° C. These cooling operations require a cooling duty of 85 kJ/kg(MSA).

[0085] Compared to a process without heat integration (Comparative Examples 1 and 2), in the inventive process the energy for heating up the feed to the distillation is reduced by 47% and the energy demand for cooling the hot streams leaving the distillation is reduced by 50%.