WO2025186008A1 - Energy-efficient method for raw syngas purification - Google Patents

Abstract

The present invention relates to a method for purifying raw syngas, which contains at least carbon monoxide, hydrogen, hydrogen sulfide and carbon dioxide, using an absorbent. The method is characterized in that energy from the unloaded absorbent is made useful in the process by means of a heat transfer medium in order to save energy costs and CO₂.

Description

Energy-efficient process for raw synthesis gas purification

The invention relates to a process for purifying raw synthesis gas, which contains at least carbon monoxide, hydrogen, hydrogen sulfide, and carbon dioxide, using an absorbent. The process is characterized by harnessing energy from the discharged absorbent within the process by means of a heat transfer medium, thereby saving energy costs and CO2 emissions.

The purification of raw synthesis gas is a process known in the literature, the so-called sulfinol process. The sulfinol process is a method for removing acidic components (carbon dioxide, hydrogen sulfide, and/or carbon monoxide sulfide) from natural gas, long-distance gas, and synthesis gas. The acidic components are scrubbed from the raw gas under pressure using an absorbent consisting of sulfolane and diisopropanolamine in an absorption column. In a subsequent process step, the acidic gas components are expelled from the loaded absorbent.

Particularly when removing gas components from the loaded absorption medium in a desorption column, a large amount of energy must be supplied externally. Heating steam is typically used as an external heat source. Rising energy prices therefore have a significant impact on the process. Furthermore, large amounts of CO2 are released during the generation of heating steam.

The objective of the present invention was to provide a purification process for synthesis gas that requires less or no external energy input. In addition, a purification process that can save CO2 emissions should be provided.

The underlying objects could be achieved by the inventive method for purifying raw synthesis gas according to claim 1. Preferred embodiments of the method are specified in the subclaims.

The process according to the invention is a process for purifying raw synthesis gas which contains at least carbon monoxide, hydrogen, hydrogen sulfide and carbon dioxide, the process comprising at least the following steps: a) bringing the raw synthesis gas into contact with an absorption medium which consists of 10 to 20 wt.% water, 40 to 50 wt.% DIPA and 35 to 45 wt.% sulfolane, the amounts of the three components water, diisopropylamine (DIPA) and sulfolane adding up to 100%, in at least one absorption column, whereby at least a portion of the hydrogen sulfide and/or at least a portion of the carbon dioxide passes into the absorption medium and a loaded absorption medium BA is formed and whereby a purified synthesis gas stream is obtained; b) feeding the loaded absorption medium BA to at least one desorption column, wherein the loaded absorption medium BA is preheated before entering the desorption column; c) heating the loaded absorption medium in the desorption column by means of a bottom evaporator SV, which is fed with a stream taken from the lower end of the desorption column and, after passing through the bottom evaporator SV, is returned to the desorption column, whereby at least a portion of the hydrogen sulfide and/or at least a portion of the carbon dioxide are outgassed from the loaded absorption medium and a discharged absorption medium EA is formed, which is obtained at the bottom of the at least one desorption column; d) returning the discharged absorption medium EA to the at least one absorption column in step a), wherein energy is transferred from EA to a liquid or gaseous heat carrier W, whereby a heat carrier W1 is produced which has an elevated temperature compared to the heat carrier W, and whereby the discharged absorption medium EA is cooled before entering the at least one absorption column; e) compressing at least a portion of the heat carrier W1, whereby a compressed heat carrier W2 is produced which has a higher pressure than the heat carrier W1; and f) transferring energy from the compressed heat carrier W2 to the stream in the bottom evaporator SV.

The process according to the invention is characterized by heat integration, in which energy is transferred in the form of heat from one part of the process to another. A heat transfer medium is used for this purpose. The absorption process is exothermic, with the absorption heat generally being extracted from the system via a cooler. The process according to the invention, however, utilizes this heat by using a heat transfer medium. The absorption heat is not dissipated via the cooling water, but is used to evaporate a heat transfer medium and then used to heat the bottom evaporator.

The raw synthesis gas used for purification comes from the production of synthesis gas. Synthesis gas can be produced in various ways, for example by crude oil cracking. During production, hydrogen sulfide and carbon dioxide can be formed. The production of synthesis gas therefore initially produces a raw synthesis gas which contains at least carbon monoxide, hydrogen, hydrogen sulfide and carbon dioxide. The amounts of carbon dioxide in the raw synthesis gas can vary and amount to a maximum of 10 vol.%, preferably a maximum of 4 vol.%, based on the total volume of the raw synthesis gas. The amounts of hydrogen sulfide can also vary, but are usually lower and amount to a maximum of 1 vol.%, preferably a maximum of 0.5 vol.%, based on the total volume of the raw synthesis gas. Such streams are used in the process according to the invention. In the first step a), raw synthesis gas is brought into contact in at least one absorption column with an absorption medium consisting of 10 to 20 wt.% water, 40 to 50 wt.% DIPA and 35 to 45 wt.% sulfolane, the amounts of the three components water, diisopropylamine (DIPA) and sulfolane adding up to 100%, whereby at least part of the hydrogen sulfide and/or at least part of the carbon dioxide are transferred into the absorption medium and a loaded absorption medium BA is formed and a purified synthesis gas stream is obtained.

The absorbent comprises at least water, diisopropylamine (DIPA), and sulfolane. In the present invention, the absorbent used consists only of the three components water, diisopropylamine (DIPA), and sulfolane. This absorbent can contain the three components in varying amounts. The absorbent consists of 10 to 20 wt.% water, 40 to 50 wt.% DIPA, and 35 to 45 wt.% sulfolane, each based on the total absorbent. The amounts of the three components water, diisopropylamine (DIPA), and sulfolane must add up to 100%.

The absorption in step a) of the process according to the invention is carried out in at least one absorption column. However, it is also possible to use more than one absorption column in step a). According to the invention, the absorption in step a) is preferably carried out in at least two absorption columns, preferably in exactly two absorption columns.

If the absorption in step a) is carried out in two absorption columns, the unloaded absorption medium EA is first passed to the second absorption column, where a partially loaded absorption medium is formed. This is then passed to the first absorption column, where the loaded absorption medium is formed, which is then passed for desorption in step c). The raw synthesis gas is then preferably first introduced into the first absorption column and from there passed to the second absorption column. The purified synthesis gas is thus removed at the top of the second absorption column.

Any absorption column known to the person skilled in the art can be used as the absorption column. The absorption column preferably contains internals. Suitable internals include, for example, trays, unstructured packings (random packing), or structured packings. Trays typically used are bubble-cap trays, sieve trays, valve trays with fixed or movable valves, tunnel trays, or slotted trays. Unstructured packings are generally random packings. Raschig rings, Pall rings, Berl saddles, SuperRings/SuperRings Plus, or Intalox® saddles are typically used as the packing. Structured packings are marketed, for example, under the trade name Mellapak® by Sulzer. In addition to the internals mentioned, other suitable internals are known to the person skilled in the art and can likewise be used. In a particularly preferred embodiment of the present invention, the absorption column comprises a plurality of internals, preferably 2 to 50 internals, more preferably 2 to 15 internals. The absorption in step a) of the process according to the invention is preferably operated at a pressure of 20 to 50 bar, particularly preferably at a pressure of 25 to 40 bar. The temperature during the absorption in step a) is preferably in the range of 30 to 80 °C, more preferably in the range of 35 to 70 °C. If the absorption in step a) is carried out in two absorption columns, the pressure and temperature in the two absorption columns are each within the above-mentioned ranges (pressure: 20 to 50 bar, preferably 25 to 40 bar / temperature 30 to 80 °C, preferably 35 to 70 °C), but do not have to be identical. This means that the absorption columns can also be operated at different pressures and temperatures.

It is further preferred that the absorption in step a) is operated in countercurrent, i.e., the crude synthesis gas and the absorption medium are brought into contact with one another in countercurrent. If two or more absorption columns are present, it is preferred if all absorption columns are operated in countercurrent. It is preferred if the crude synthesis gas is introduced into the lower part of the at least one absorption column and the absorption medium into the upper part of the at least one absorption column. The crude synthesis gas will then flow upwards and come into contact with the descending absorption medium. The impurities (hydrogen sulfide and carbon dioxide) are washed out in the process. The internals preferably present in the at least one absorption column can achieve better distribution and thus improved contact between the crude synthesis gas and the absorption medium.

In step a), a purified synthesis gas stream is produced. The purified synthesis gas stream can then be used as a raw material. The purified synthesis gas stream can be fed to a hydroformylation plant, where it can be used to convert olefins into aldehydes. It would also be possible to split the synthesis gas into its components (hydrogen and carbon monoxide) and feed the two components to different synthesis processes, for example, the carbon monoxide to an alkoxycarbonylation plant and the hydrogen to a hydrogenation plant.

In step b) of the process according to the invention, the loaded absorbent BA is fed to at least one desorption column, preferably only a single desorption column, wherein the loaded absorbent BA is preheated before entering the desorption column. After absorption, the loaded absorbent BA has a temperature in the range of 30 to 80 °C, preferably 35 to 70 °C. However, desorption takes place at a higher temperature than absorption. In this respect, it is advantageous to preheat the loaded absorbent BA before entering the desorption column, since not all of the desorption energy needs to be introduced via the bottom evaporator of the desorption column.

The preheating of the loaded absorption medium BA is preferably carried out by transferring energy from the discharged absorption medium EA to the loaded absorption medium BA in a heat exchanger. This involves a direct transfer of energy from the discharged absorption medium EA to the loaded absorption medium BA. Direct transfer means that the two streams Although they do not contact each other directly, energy, especially heat, is transferred from EA to BA without the presence of an additional heat transfer medium.

The preheating in step b) can be carried out using heat exchangers known to those skilled in the art. Suitable evaporators that can be used as heat exchangers include, for example, natural circulation evaporators, forced circulation evaporators, forced circulation evaporators with expansion, boiler evaporators, falling-film evaporators, or thin-film evaporators. In addition to the aforementioned evaporators, any other evaporator design known to those skilled in the art that is suitable for the desired preheating can also be used.

The separation of the acid gas from the laden absorption medium takes place in at least one desorption column. Any desorption column known to those skilled in the art can be used as the desorption column. The desorption column preferably contains internals. Suitable internals include, for example, trays, unstructured packings (random packing), or structured packings. Trays typically used are bubble-cap trays, sieve trays, valve trays with fixed or movable valves, tunnel trays, or slotted trays. Unstructured packings are generally random packings. The packings typically used are Raschig rings, Pall rings, Berl saddles, SuperRings/SuperRings Plus, or Intalox® saddles. Structured packings are marketed, for example, under the trade name Mellapak® by Sulzer. In addition to the internals mentioned, other suitable internals are known to those skilled in the art and can also be used. In a particularly preferred embodiment of the present invention, the desorption column comprises a plurality of internals, preferably 2 to 50 internals, particularly preferably 2 to 15 internals.

Step c) of the process according to the invention involves heating the loaded absorption medium in the desorption column by means of a bottom evaporator SV. The bottom evaporator SV is fed with a stream taken from the lower end of the desorption column and, after passing through the bottom evaporator SV, is returned to the desorption column, whereby at least some of the hydrogen sulfide and/or at least some of the carbon dioxide are released from the loaded absorption medium, producing a discharged absorption medium EA.

According to the invention, bottom evaporators are referred to as evaporators that heat the bottom of the desorption column. Such a bottom evaporator is usually located outside the respective desorption column. Since energy, in particular heat, is transferred from one stream to another in bottom evaporators, they act as heat exchangers. The stream to be evaporated, the at least partially loaded absorbent, is withdrawn from the bottom of the desorption column via an outlet and fed to the bottom evaporator. The evaporated stream, optionally with a residual liquid portion, is returned to the respective desorption column in the bottom region via at least one inlet. Suitable evaporators that can be used as bottom evaporators include natural circulation evaporators, forced circulation evaporators, forced circulation evaporators with expansion, kettle evaporators, falling-film evaporators, or thin-film evaporators. A tube bundle or plate evaporator is typically used as the heat exchanger for the evaporator in natural circulation evaporators and forced circulation evaporators. In addition to the above, any other evaporator design known to those skilled in the art that is suitable for use in a desorption column can also be used.

The desorption in step c) is preferably carried out at a pressure of 1 to 3 bar, particularly preferably at a pressure of 1.2 to 2 bar. The temperature in step c) of the process according to the invention should preferably be in the range of 110 to 140 °C, particularly preferably in the range of 115 to 130 °C.

The aim of desorption is to remove hydrogen sulfide and carbon dioxide from the loaded absorption medium BA. This occurs via elevated temperature. As a result, in step c), an acid gas containing at least hydrogen sulfide and/or carbon dioxide is withdrawn from the top of the desorption column. In addition, a discharged absorption medium EA is obtained, from which a large portion, i.e., preferably > 90%, particularly preferably > 99%, of the absorbed hydrogen sulfide and carbon monoxide has been removed. The discharged absorption medium EA is generally obtained at the bottom of the at least one desorption column.

The discharged absorption medium EA is returned in step d) to the at least one absorption column in step a). The discharged absorption medium is obtained at a temperature in the range of 100 to 140 °C. The high temperature is necessary to remove hydrogen sulfide and carbon monoxide from the absorption medium. At this high temperature, the discharged absorption medium cannot be returned to the at least one absorption column in step a). Instead, the discharged absorption medium must be cooled. In the present case, this is achieved by the heat integration according to the invention, in which energy is transferred from EA to a liquid or gaseous heat transfer medium W, thereby forming a heat transfer medium W1 and cooling the discharged absorption medium EA before entering the at least one absorption column.

The transfer of energy, in particular heat energy, from the discharged absorption medium EA to the heat transfer medium W in step d) preferably takes place in a heat exchanger. A direct transfer of energy from the discharged absorption medium EA to the heat transfer medium W takes place. Direct transfer means that the two streams do not directly contact each other, but that energy, in particular heat, is transferred from EA to W without the presence of an additional heat transfer medium.

The energy transfer in step d) can be carried out using heat exchangers known to those skilled in the art. Suitable evaporators that can be used as heat exchangers include, for example, natural circulation evaporators, forced circulation evaporators, forced circulation evaporators with expansion, Boiler evaporators, falling-film evaporators, or thin-film evaporators. In addition to the above, any other evaporator design known to those skilled in the art that is suitable for the desired preheating can also be used.

A variety of different substances or mixtures can be used as heat transfer media. The heat transfer media W is preferably selected from the group consisting of water; alcohols; alcohol-water mixtures; salt-water solutions; ammonia; mineral oils, such as diesel oils; thermal oils, such as silicone oils; biological oils, such as limonene; and aromatic or aliphatic hydrocarbons, such as dibenzyltoluene. The heat transfer media W is water, methanol, ethanol, propanol, n-pentane, n-butane, n-hexane, n-propane, or ammonia. The heat transfer media W is preferably n-pentane.

If the heat transfer medium W is used in the liquid phase, i.e., a liquid heat transfer medium W, in step d) and energy, preferably thermal energy, is supplied to it in step d), the heat transfer medium W is at least partially evaporated, thereby obtaining an at least partially gaseous heat transfer medium W1. If the heat transfer medium W is used in the gaseous phase, i.e., a gaseous heat transfer medium W, in step d) and energy, preferably thermal energy, is supplied to it in step d), a gaseous heat transfer medium W1 is obtained.

In the context of the present invention, the formulation of liquid heat transfer medium means that > 50 wt.%, further preferably > 55 wt.%, further preferably > 75 wt.%, further preferably > 90 wt.% and particularly preferably > 99 wt.% of the heat transfer medium used in step d) is in the liquid state, in each case based on the total weight of the heat transfer medium used in step d).

In the context of the present invention, the formulation of gaseous heat transfer medium means that > 30 wt.%, further preferably > 50 wt.%, further preferably > 75 wt.%, further preferably > 90 wt.% and particularly preferably > 99 wt.% of the heat transfer medium used in step d) is in the gaseous state, in each case based on the total weight of the heat transfer medium used in step d).

After the described energy transfer in step d), the heat transfer medium W1 has a higher temperature and/or a higher pressure than the heat transfer medium W. In a preferred embodiment of the present invention, W1 has a temperature in the range from 30 °C to 110 °C, preferably in the range from 40 °C to 60 °C. The pressure of W1 is preferably in the range from 0.5 bar to 5 bar, more preferably 1 bar to 4 bar. It goes without saying that the heat transfer medium W1 corresponds to the heat transfer medium W and that W and W1 differ only in their respective pressure and/or temperature and, if W was used as a liquid, in their state of aggregation. In the subsequent step e), at least a portion of the heat transfer medium W1 is compressed, resulting in a heat transfer medium W2 that is more compressed than W1 and has a higher pressure than the heat transfer medium W1. The pressure of W2 after compression is preferably in the range from 1 bar to 20 bar, preferably 2 bar to 15 bar.

The compression of at least a portion of the heat transfer medium W1 in step e) can be carried out in any desired manner known to the person skilled in the art. For example, the compression can be carried out mechanically and in a single-stage or multi-stage process. In this context, "single-stage" means that compression takes place from one pressure level to another. "Multi-stage" means that compression is first carried out to a pressure level X and then from X to pressure level Y. In multi-stage compression, several compressors of the same design or compressors of different designs can be used. Multi-stage compression can be carried out with one or more compressor machines. The use of single-stage compression or multi-stage compression depends on the compression ratio and thus on the pressure to which the heat transfer medium W1 is to be compressed.

Any compressor known to the person skilled in the art, preferably a mechanical compressor with which gas streams can also be compressed, is suitable as a compressor in the process according to the invention, in particular for compressing the heat transfer medium W1 in step e). Suitable compressors include, for example, single- or multi-stage geared turbocompressors, piston compressors, screw compressors, centrifugal compressors, or axial compressors.

In step h) of the process according to the invention, energy is transferred from the compressed heat transfer medium W2 to the stream in the bottom evaporator SV. As a result of step f), the energy of W2 decreases, so that W2 preferably condenses at least partially. The transfer of energy from W2 to the stream in the bottom evaporator SV, preferably the heating of the stream in the bottom evaporator SV by W2, preferably takes place directly. Direct transfer means that W2 and the stream in SV do not come into direct contact, but that energy, in particular thermal energy, is transferred from W2 to the stream in the SV without the presence of an additional heat transfer medium. The heat transfer devices or heat exchangers familiar to the person skilled in the art, in particular evaporators, can be used as the bottom evaporator SV.

The energy transfer in step f) can be carried out by methods known to those skilled in the art or using heat exchangers known to those skilled in the art. Suitable evaporators that can be used as heat exchangers include, for example, natural circulation evaporators, forced circulation evaporators, forced circulation evaporators with expansion, kettle evaporators, falling-film evaporators, or thin-film evaporators. In addition to the aforementioned evaporators, any other evaporator design known to those skilled in the art that is suitable for use in a distillation column can also be used. The energy transfer in step f) can occur in different ways. In one variant of the process according to the invention, the compressed heat transfer medium W2 itself can be used in the heat exchanger to transfer energy from W2 to the stream in the bottom evaporator Sv. When energy is transferred from the heat transfer medium W2 to the stream in the bottom evaporator Sv, some of the absorption media mentioned are disadvantageous. For example, the evaporation temperature for n-butane, ammonia, and propanol k is above the critical point, and sufficient condensation cannot be achieved. However, a person skilled in the art would take this into account when selecting the appropriate heat transfer medium based on their usual specialist knowledge.

Another variant of the process according to the invention provides that energy is first transferred to another heat transfer medium, such as water. The energy transfer in step f) then occurs in such a way that steam is generated via the compressed heat transfer medium W2 in an evaporator, which steam is used for heating in the bottom evaporator SV. The evaporators mentioned above can be used as evaporators. In this case, the steam is preferably water vapor.

Regardless of whether the energy in step f) is transferred by W2 itself or by the generated steam, W2 or the steam can be used for further preheating of the loaded absorbent BA. This allows for additional energy savings that would otherwise have to be supplied externally.

In a preferred embodiment of the present invention, the heat transfer medium W2 is recycled after the energy transfer in step f) and reused as heat transfer medium W in step d). The heat transfer medium is thus circulated, and only a single heat transfer medium is used in steps d), e), and f).

In a particularly preferred embodiment of the present invention, the heat transfer medium W1 is compressed twice in step e). In the first compression V1, a compressed heat transfer medium W1.1 is produced which has a higher pressure than W1, and in the second compression V2, the compressed heat transfer medium W2 is produced which has a higher pressure than W1.1. W2 is then used for the energy transfer in step f), as already described. Two individual compressors can be used for the double compression. However, the use of a single multi-stage compressor is also possible, with which both compressions V1 and V2 can be carried out independently of one another.

The heat transfer medium W2 is preferably at least partially condensed during the energy transfer in step f). It is further preferred that the at least partially condensed heat transfer medium W2 is fed into a flash tank operated at the pressure level between the two compression stages V1 and V2. Particularly preferably, the flash portion in the flash tank is passed as a gas phase between the two compression stages. This allows a portion of the energy present in the heat transfer medium W2 after the energy transfer in step f) to be utilized. The flash portion is mixed with the compressed heat transfer medium W1.1, the second compression stage V2. and led to energy transfer in step f). Only the liquid portion in the flash tank is then led to step d) as heat transfer medium W and passes through steps d), e), and f).

The present invention will be described with reference to the drawings described below.

Fig. 1 shows a non-inventive embodiment of the process. The crude synthesis gas (1) contaminated with CO2 and H2S is contacted with an absorption medium in countercurrent in the first absorption column (2). A gas stream (3) is withdrawn from the top of the first absorption column (2), passed into the second absorption column (5), and again contacted with the absorption medium in countercurrent. The purified synthesis gas (6) is withdrawn from the top of the second absorption column and used in other plants, for example in hydroformylation. The partially loaded absorption medium (7) obtained in the second absorption column is fed to the first distillation column (2). The loaded absorption medium (4) is withdrawn from the bottom of the column (2), subjected to heat integration in the heat exchanger (8), and thereby heated. Before being fed into the desorption column (10), it is further heated by means of low-pressure steam via a preheater (9). In the desorption column (10), the absorption medium is freed of hydrogen sulfide and carbon dioxide, and a sour gas (15) is removed at the top of the column (10). Energy is introduced into the system via the bottom evaporator SV (11). The discharged absorption medium (12) is removed from the bottom of the desorption column (10) and fed via a booster pump (13) for heat integration at the heat exchanger (8). Before being reused in the second absorption column (5), the absorption medium is cooled with cooling water at the heat exchanger (14).

Fig. 2 shows an embodiment of the method according to the invention, which corresponds in large parts to Fig. 1. In addition to the circuitry described in Fig. 1, the thermal energy of the discharged absorption medium is partially utilized using a heat pump. For this purpose, the heat transfer medium W, preferably n-pentane, is at least partially evaporated in the heat exchanger (25). The heat transfer medium is preheated via the heat exchanger (20) and subjected to multi-stage compression in two compressor stages (21, 22). In the heat exchanger (23), heating steam is generated via the heat transfer medium, which is used for heating in the preheater (9) and the bottom evaporator (11). The at least partially condensed steam is fed back into the heat exchanger (23) via a booster pump (26). The heat transfer medium is subjected to an intermediate flash (24). The gas phase generated in the intermediate flash (24) is recirculated to a section between the compressor stages (21 and 22). The liquid portion from the intermediate flash is fed to the heat exchanger (25).

Fig. 3 shows an embodiment of the method according to the invention, which largely corresponds to Fig. 1. In addition to the circuitry described in Fig. 1, the thermal energy of the discharged absorption medium is partially utilized with a heat pump. In the present case, the heat transfer medium is used directly as a heating medium. For this purpose, the heat transfer medium W, preferably n-pentane, is at least partially evaporated in the heat exchanger (25). The heat transfer medium is preheated via the heat exchanger (20) and subjected to a multi-stage compression in two compressor stages (21, 22). During the first compression, a partial flow of the heat transfer medium can be used for preheating (9). The other part of the heat transfer medium is fed to the second compression stage (22) and subsequently used for heating in the bottom evaporator (11).

The following examples are intended to illustrate the invention without limiting its scope of application, which is apparent from the description and the claims.

Examples

The same feed stream is used as the basis for all subsequent examples. The raw syngas composition is considered as follows: 51 vol.% CO, 45 vol.% H2, 3.5 vol.% CO2, 0.3 vol.% H2S, and 0.2 vol.% H2O. The calculations were performed using Aspen Plus® Version 10 and an adapted property model based on operationally validated property data.

Example 1 (not according to the invention)

The simulation was carried out using the process configuration shown in Fig. 1. The following parameters were taken into account in the simulation. The absorbent (fresh: 15% water, 45% DIPA and 40% sulfolane) is fed at a rate of 80 t/h into the first desorption column (2), which is operated at a pressure of 36 bar abs. The purified synthesis gas (6) is withdrawn from the second absorption column (5) at 35 bar abs and 47°C. The loaded absorbent is withdrawn from the bottom of the first absorption column (2) at approx. 60°C and heated to 102°C in the heat exchanger (8). It is heated to 118°C by the preheater (9). The preheater is operated with 1.2 t/h of low-pressure steam. Energy is introduced into the system via the bottom evaporator (11) on the desorption column (10) using 4.2 t/h of low-pressure steam to provide the required 80 t/h of discharged absorption medium (12). Before being reused in the second absorption column (5), the absorption medium is heated to the optimal absorption temperature of approximately 47°C using cooling water in the heat exchanger (14). At the heat exchanger (14), 2.3 MW of energy is removed from the system using recooling water. A total of 5.4 t/h of steam (= 3 MW heating capacity) is required to operate the purification of raw synthesis gas.

Example 2 (according to the invention)

The simulation was carried out using the process configuration shown in Fig. 2. The following parameters were taken into account in the simulation. During heat integration, n-pentane is evaporated in the heat exchanger (25) at 1.2 bar abs and 40°C. The n-pentane is preheated via the heat exchanger 20 and subjected to multi-stage compression (21, 22) using a machine. This requires an electrical output of 1.1 MW. In the heat exchanger 23, 5.4 t/h of heating steam at 4 bar abs is generated from the upgraded waste heat, which is used to heat the preheater (9) and bottom evaporator. A total of 3 MW of heating output (by 5.4 t/h steam) was replaced with 1.1 MW of electrical output. Example 3 (according to the invention)

The simulation was performed using the process configuration shown in Fig. 3. During heat integration, n-pentane is vaporized in the heat exchanger (25) at 1.2 bar abs and 40°C. The n-pentane is preheated via heat exchanger 20 and subjected to multi-stage compression (21, 22). This requires a total electrical power of 0.95 MW. A total of 3 MW of heating power (by 5.4 t/h of steam) was replaced with 0.95 MW of electrical power.

Examples 2 and 3 according to the invention demonstrate a process for raw synthesis gas scrubbing in which steam can be replaced by heat integration. By substituting steam with green electricity, up to 9000 CO2eqt/a can be saved and operating costs reduced.

Claims

Patent claims 1. A process for purifying raw synthesis gas which contains at least carbon monoxide, hydrogen, hydrogen sulfide and carbon dioxide, the process comprising at least the following steps: a) bringing the raw synthesis gas into contact with an absorption medium which consists of 10 to 20 wt.% water, 40 to 50 wt.% DIPA and 35 to 45 wt.% sulfolane, the amounts of the three components water, diisopropylamine (DIPA) and sulfolane adding up to 100%, in at least one absorption column, whereby at least a portion of the hydrogen sulfide and/or at least a portion of the carbon dioxide passes into the absorption medium and a loaded absorption medium BA is formed and a purified synthesis gas stream is obtained; b) feeding the loaded absorption medium BA to at least one desorption column, the loaded absorption medium BA being preheated before entering the desorption column; c) heating the loaded absorption medium in the desorption column by means of a bottom evaporator SV, which is fed with a stream taken from the lower end of the desorption column and, after passing through the bottom evaporator SV, is returned to the desorption column, whereby at least a portion of the hydrogen sulfide and/or at least a portion of the carbon dioxide are outgassed from the loaded absorption medium and a discharged absorption medium EA is formed, which is obtained at the bottom of the at least one desorption column; d) returning the discharged absorption medium EA to the at least one absorption column in step a), whereby energy is transferred from EA to a liquid or gaseous heat transfer medium W, whereby a heat transfer medium W1 is formed which has an elevated temperature compared to the heat transfer medium W, and whereby the discharged absorption medium EA is cooled before entering the at least one absorption column; e) compressing at least a portion of the heat transfer medium W1 to produce a compressed heat transfer medium W2 having a higher pressure than the heat transfer medium W1; and f) transferring energy from the compressed heat transfer medium W2 to the stream in the bottom evaporator SV. 2. Process according to claim 1, wherein step a) is carried out in at least two absorption columns, preferably in exactly two absorption columns. 3. Process according to one of the preceding claims, wherein step a) is operated in countercurrent, i.e. the raw synthesis gas and the absorption medium are brought into contact with each other in countercurrent. 4. The process according to claim 3, wherein the raw synthesis gas is introduced into the lower part of the at least one absorption column and the absorbent is introduced into the upper part of the at least one absorption column. 5. Process according to one of the preceding claims, wherein the heat transfer medium is water, methanol, ethanol, propanol, n-pentane, n-butane, n-hexane, n-propane or ammonia, particularly preferably n-pentane. 6. Method according to one of the preceding claims, wherein the heat transfer medium W2 is returned after the energy transfer in step f) and used again as heat transfer medium W in step d). 7. The process according to claim 2, wherein, when the absorption in step a) is carried out in two absorption columns, the discharged absorbent EA is passed to the second absorption column, where a partially loaded absorbent is formed, which is then passed to the first absorption column, where the loaded absorbent is formed, which is passed for desorption in step c). 8. The process according to claim 7, wherein the raw synthesis gas is introduced into the first absorption column. 9. Method according to one of the preceding claims, wherein the preheating of the loaded absorption medium BA takes place in that energy is transferred from the discharged absorption medium EA to the loaded absorption medium BA in a heat exchanger. 10. The method according to any one of the preceding claims, wherein the transfer of energy in step f) is carried out in such a way that steam, preferably water steam, is generated via the compressed heat carrier W2 in an evaporator, which steam is used for heating in the bottom evaporator SV. 11. Method according to one of the preceding claims, wherein the heat transfer medium W1 is compressed twice, wherein in the first compression a compressed heat transfer medium W1 .1 is produced which has a higher pressure than W1 and in the second compression the compressed heat transfer medium W2 is produced which has a higher pressure than W1 .1. 12. The method according to claim 10 or 11, wherein the heat transfer medium W2 condenses at least partially during the energy transfer. 13. The method according to claim 12, wherein the at least partially condensed heat transfer medium W2 is fed into a flash tank which is operated at the pressure level which lies between the two compression stages. 14. The method according to claim 13, wherein the flash portion in the flash container is passed between the two compression stages.