Sustainable C3 process with integrated raw material production

Description

Field of the invention

[0001] The present invention relates to a novel process for producing organic compounds with one to three carbon atoms. This process utilizes reactants obtained from hydrogen cyanide production, which is part of a process suitable for large-scale chemical use. Furthermore, the organic compound thus obtained is used as an additional reactant in the same process. In particular, the present process relates to integration into a process for producing MMA, wherein the organic compound is a raw material, in particular an oxygenate such as acetone, methanol, ethanol, or propanol. State of the art

[0002] The global goal is to reduce greenhouse gas emissions in the coming years and to improve chemical processes, a relevant source of these emissions. CO2 emissions, in particular, are to be reduced. The state of the art describes various approaches to reducing greenhouse gas emissions with regard to fundamental aspects of chemical processes. For example, methods are described for capturing emitted CO2 and converting it into organic compounds that can then be used in industrial processes. Furthermore, processes are described that improve the CO2 balance in industrial processes through the use of non-fossil fuels, such as biogas.

[0003] For example, Liew et al. (Nature Biotech, Vol. 40, March 2022, pp. 335-344) describe the production of acetone and isopropanol from industrial waste gases by fermentation. In this process, CO 2 and hydrogen (H 2 ) from industrial plants are converted by the bacterium Clostridium autoethanogenum converted to acetone and isopropanol.

[0004] There are some chemicals produced in relevant quantities on an industrial scale that can be improved with regard to their respective greenhouse gas footprint. It is particularly interesting to improve those processes that a) are used at high capacities and b) have a high level of complexity, e.g., with regard to the multi-stage nature of the process. Such molecules include, for example, methyl methacrylate (MMA), methionine, or isophorone derivatives. A particular focus of the present invention is on improving precisely these production processes, in particular in the production of MMA: Various processes are known for the production of alkyl (meth)acrylates, in particular methyl methacrylate (MMA). These processes typically start from C2, C3, or C4 building blocks.

[0005] US Pat. No. 9,938,225 B2 describes a process for the production of methyl methacrylate with a reduced CO2 footprint. The acetone required for the C3-based synthesis is produced by fermentation of C6 and C5 sugars.

[0006] A globally used commercial process is based on acetone as the starting material and is usually referred to as the C3 process or ACH-sulfo process. In this process, acetone is reacted with hydrogen cyanide (HCN) to form the central intermediate acetone cyanohydrin (ACH). This intermediate is isolated and used in the subsequent process steps for the production of methacrylic acid (MAS) and MMA. One such process is described, for example, in US Pat. No. 4,529,816.

[0007] Processes for producing the hydrogen cyanide converted in the C3 process are also known as such and include, for example, the so-called BMA process, in which the hydrogen cyanide is obtained from methane and ammonia, and the Andrussow process. The Andrussow process is described, for example, in Ullmann's Encyclopedia of Industrial Chemistry, Volume 8, VCH Verlagsgesellschaft, Weinheim 1987, pages 161-162. In this process, the reactant gas mixture, which comprises methane, ammonia and oxygen, is passed over catalyst meshes in a reactor and reacted at temperatures of around 1000 °C. The required oxygen is usually used in the form of air. The resulting product stream contains hydrogen cyanide (HCN), unreacted ammonia and methane, as well as CO, H2, H2O and CO2 as by-products.

[0008] The BMA process is described, for example, in Ullmann's Encyclopedia of Industrial Chemistry, Volume 8, VCH-Verlagsgesellschaft, Weinheim 1987, pages 162-163.

[0009] The processes described in the prior art exhibit relatively high greenhouse gas emissions, particularly CO2 emissions. Furthermore, they do not allow for the reuse of exhaust gases and / or byproducts. Task

[0010] The object of the present invention was to provide a sustainable process for the production of organic compounds, in particular those with one to three carbon atoms.

[0011] The particular task was to integrate such a process into another synthetic production process, both with regard to the reactant streams and with regard to the use of the organic compound.

[0012] The main task was to combine the novel process with a production process for hydrogen cyanide.

[0013] In particular, the task was to realize methanol, ethanol, propanol or acetone, especially acetone, as organic compounds.

[0014] The additional task was to integrate the process into a process for producing an alkyl methacrylate, in particular MMA.

[0015] Further tasks not explicitly mentioned may arise directly or implicitly from the description, the claims or the examples. Solution

[0016] The tasks were solved by providing a novel process for the production of hydrogen cyanide and organic compounds containing one to three carbon atoms and oxygen. This process is characterized by the following process steps: I. Production of a mixture of hydrogen cyanide and hydrogen in a reaction zone A1 of a reactor A. This production starts from ammonia and a methane-containing fraction. The methane-containing fraction can in particular be natural gas. The mixture thus obtained is referred to below as process gas I. II. Separation of the process gas I into a process gas II consisting predominantly of hydrogen cyanide and a process gas III consisting predominantly of hydrogen. III. Production of a process gas IV consisting predominantly of CO2 in a reaction zone A2 of the reactor A, starting from a methane-containing fraction. IV. Mixing of at least one partial stream of the process gas III with at least one partial stream of the process gas IV to obtain a process gas VV. Converting the process gas V in a reactor B to produce a process stream VI comprising the organic compound to be produced.

[0017] The process according to the invention is particularly advantageous if, in a further process step VI, at least a portion of the hydrogen cyanide obtained in process step I and at least a portion of the organic compound obtained in process step V are used to produce a common secondary product.

[0018] There are further preferred embodiments for the individual process steps, which, unless otherwise stated, can also be combined with each other. Such combinations can also relate to several of the following embodiments in different process steps: For process step I:

[0019] Process step I is preferably carried out in a reactor A in a reaction zone A1 under the substantial exclusion of oxygen. This is particularly efficient if reaction zone A1 contains a platinum-containing catalyst as a heterogeneous catalyst.

[0020] Furthermore, the process gas I is preferably treated with an acidic aqueous absorbent in an absorption device I prior to process step II, such that any remaining ammonia is separated. The separated ammonia is obtained as an aqueous ammonium sulfate-containing solution. This ammonia separation is also referred to below as process step 1a.

[0021] Particularly preferably, the ammonia-depleted process gas II from process step 1a is treated with an aqueous, optionally carbonyl-containing absorbent in at least one absorption device II to obtain an aqueous HCN-containing solution. This solution can then be treated in at least one distillation column I to obtain, after condensation, a liquid product fraction I containing at least 99% by weight of hydrogen cyanide. For process step II:

[0022] It is particularly advantageous if, when carrying out the process according to the invention, process gas III has a hydrogen content greater than 90 vol%. Particularly preferably, this content is optionally increased to over 95 vol% in one or more processing steps between process steps II and IV. It is also advantageous if this optionally enriched process gas III is obtained at a pressure below 5 bara. For process step III:

[0023] The methane-containing fraction fed into reaction zone A2 in process step III is preferably natural gas. It has proven particularly advantageous to feed a mixture of oxygen and process gas IVa, optionally with hydrogen, to this methane-containing fraction in or upstream of reaction zone A2, preferably in countercurrent. The oxygen is particularly preferably fed directly into the reactor in countercurrent to the methane-containing gas. Pure oxygen, rather than air, is most preferably fed into the reactor. This has the advantage that the reaction can be operated virtually free of nitrogen with respect to reactor zone A2, and thus the subsequent process gases IV and V are free of disruptive nitrogen oxides.

[0024] In this context, it should also be noted that—as is already preferred in prior art processes—it is advisable to select the methane-containing fractions so that they are as sulfur-free as possible to avoid the formation of disruptive sulfur oxides. In the case of sulfur-containing (e.g., hydrogen sulfide-containing) methane-containing fractions, it is preferable to desulfurize them before introducing them into reactor zones A1 and A2.

[0025] In addition to the generation of process gas IV, process step III preferably has the purpose of providing heat energy for the endothermic reaction of process step I.

[0026] After exiting reaction zone A2, process gas IV preferably has a temperature between 700 °C and 1300 °C, an oxygen content of less than 5 vol%, a NO x content of less than 500 vol-ppm, preferably less than 50 vol-ppm, and a total CO, CO 2 and water content of 50 to 99 vol%, preferably 70 to 98 vol%, and particularly preferably 80 to 97 vol%. Process gas IV is then cooled to a temperature between 150 °C and 450 °C, preferably to generate medium-pressure steam.

[0027] Process stream IV preferably has an oxygen content of less than 0.1 vol%, particularly preferably less than 0.01 vol%.

[0028] In a preferred embodiment of the present invention, process gas IV can be obtained by reacting a methane-containing fraction, in particular natural gas, with oxygen. It is particularly preferred if a partial stream of this process gas IV is recycled directly to reaction zone A2 as a process gas IVa.

[0029] It has proven advantageous if 20 vol% to 50 vol%, particularly preferably 25 vol% to 35 vol% of the process gas IV is recycled as process gas IVa.

[0030] Alternatively, and just as preferably, a second partial stream IVb of the process gas IV is recycled to reaction zone A2. This can particularly preferably be recycled before or upon introduction into

[0031] Reactor zone A2 must first be mixed with oxygen. Optionally, although less preferred, mixing with the methane-containing fraction can also occur here – instead of partial stream IVb.

[0032] An embodiment of the invention has proven particularly advantageous in which a partial stream of the process gas IV is passed into reaction zone A2 for process step III and a second partial stream of the process gas IV is passed into reactor B for process step V. The ratio of these two partial streams to one another depends in particular on the respective requirements for process steps III and V or VI, with the consumption in process steps V and subsequently VI generally being decisive.

[0033] Optionally, additional partial streams of process gas IV can be diverted. These could theoretically, although not preferentially, be disposed of or temporarily stored. For process step IV:

[0034] As a rule, process gas V is a synthesis gas with a molar ratio of CO and CO2 to hydrogen between 1 to 1 and 1 to 10, preferably between 1 to 1.5 and 1 to 7.5.

[0035] It is also advantageous if the process gas IV is subjected to deoxygenation before process step IV.

[0036] In the event that process gas IV is separated into two or more partial streams as described above, deoxygenation is generally only carried out on the partial stream of process gas IV which is later fed into reactor B. This therefore concerns the partial stream which is fed to process step IV.

[0037] Alternatively or additionally, the process gas V can also be subjected to deoxygenation before process step V.

[0038] Particularly preferably, the deoxygenation is carried out adiabatically in contact with a transition or noble metal catalyst at a temperature between 200 and 800 °C. For process step V:

[0039] With regard to process step V and the preferably subsequent process step VI, various options or embodiments of the present invention are available. The particularly preferred embodiments of the present invention are discussed below, but this does not mean that there are no further, also economically interesting, embodiments of process steps V and optionally VI: In a first preferred embodiment, the organic compound formed in process step V is methanol.

[0040] Particularly preferably, the synthesis of methanol in process step V is carried out by reaction in the gas phase on a heterogeneous catalyst at pressures of 20 to 120 bar and temperatures between 130 °C and 350 °C. For many applications, after process step V, process stream VI is separated by condensation into a water phase, a phase consisting predominantly of methanol, and a non-condensed process gas VII.

[0041] The water phase can be partially recycled to the condensation stage or to a subsequent phase separator. It is also possible to first concentrate the water phase using distillation or a membrane separation stage before it is recycled. However, the concentrated or original water phase is usually disposed of.

[0042] Process gas VII can be subjected to further processing, for example, to isolate any remaining methanol. Alternatively, process gas VII, which contains in particular C1 to C3 building blocks, can be fed to one of the methane-containing fractions of process step I or, preferably, III. Furthermore, it is also possible to subject process gas VII to combustion. Combinations of these processes using partial streams of process gas VII are also possible. Preferably, however, process gas VII is mixed with process gas V and thus recycled to reactor B.

[0043] The methanol phase can optionally be further purified or directly fed to another synthesis in process step VI.

[0044] In this first embodiment of process step V, reactor B preferably comprises several sub-reactors running in series. These are particularly preferably tube-bundle reactors and / or tray reactors. In such a configuration, a partial conversion of between 20% and 90% is generally achieved in the respective sub-reactors. It has proven advantageous to cool the respective gas mixture before entering the next sub-reactor.

[0045] With regard to this methanol obtained in the first embodiment of process step V, there are several variants for the optional, but preferably carried out, process step VI:

[0046] In a first particularly preferred variant of the first embodiment, methyl methacrylate is produced in process step VI from the hydrogen cyanide produced in process step I and the methanol obtained in process step V. In particular, the methanol can be added to a mixture containing sulfuric acid, ammonium hydrogen sulfate, methacrylamide, and water at a temperature between 80°C and 130°C, and the methacrylamide can be previously produced using the hydrogen cyanide. This variant would thus mean that the process according to the invention would be integrally integrated into the so-called and well-described ACH process (see, for example, DE 744877 C, EP21 257 06) for producing MMA.

[0047] In a second variant, methyl mercaptan is first produced in process step VI from the methanol obtained in process step V together with hydrogen sulfide. The hydrogen sulfide can optionally be produced using a partial stream of process gas II. The methyl mercaptan thus obtained can then be reacted in a subsequent step with the hydrogen cyanide produced in process step I to form methionine (see, for example, 2 DE 2122491).

[0048] In a second embodiment, process step V involves a fermentation, and the organic compound is acetone. In a particularly preferred variant of this embodiment, the resulting product in process step VI is methacrylamide, from which an alkyl methacrylate or methacrylic acid can be produced.

[0049] In this variant, the acetone is particularly preferably contacted with the hydrogen cyanide from process step I at a temperature between 0 °C and 60 °C using a basic catalyst. The acetone cyanohydrin (ACH) thus produced can then, after purification with sulfuric acid, methanol, and water, be converted in at least two steps to a methacrylate via the intermediate stage of the methacrylamide.

[0050] In a third embodiment, process step V is a fermentation, and the organic compound is ethanol or propanol. In a preferred variant of this third embodiment, the subsequent product in process step VI is ethyl methacrylate or propyl methacrylate, which is obtained by esterification of a methacrylic acid amide or methacrylic acid based on hydrocyanic acid from process step I, or by transesterification of an MMA based on hydrocyanic acid from process step I with the ethanol or propanol from process step V.

[0051] In a second variant of the third embodiment, the downstream product in process step VI is isophoronediamine. The isophoronediamine is obtained by reacting hydrogen cyanide from process step I with isophorone, which was at least partially produced from acetone from process step V.

[0052] Regardless of the design or variant, process stream VI can be subjected to further separation before the optional process step VI. This produces an additional fraction consisting of an emission gas with a nitrogen oxide content of less than 150 ppm by volume.

[0053] For further clarification, it should be noted that if the organic compound obtained in process step V is methanol, this procedure will result in a fourth fraction, as just described, being obtained after the condensation from which an aqueous and a methanol-containing fraction as well as a process gas VII are obtained as described. List of reference symbols

[0054] Figure 1 shows a schematic overview of the method according to the invention. Figure 2 shows schematically the current process to which the comparative example refers. Figures 3 to 5each show a detailed view of the oxygenate reactor and its purification for the inventive examples E3 to E5. Fig. 1 A BMA reactor A1 Reaction side BMA reactor A2 Heating side BMA reactor B Absorption I (sulfuric acid wash) C Absorption II & Distillation I D Optional H2 purification (pressure swing adsorption) E Pure HCN tank F ACH reactor G ACH distillation H MMA reactor including purification J sulfuric acid reactor K Deoxo-Kat L Reactor C1-C3 Oxygenate M Purification of C1-C3 oxygenate N Reactor including purification of methyl mercaptan 1 Ammonia (NH3) 2 Methane (CH4) 3 Oxygen (O2) 4 Hydrogen (H2) 5 sulfur 6 Air 7 Sulfuric acid (H2SO4) 8 Cleavage acid 9 Fresh acetone 11 MMA 12 natural gas 13 Gas mixture for hydrogen cyanide production 14 Flue gas BMA reactor 15 Flue gas recycling (process gas IV) 16 Mixture of natural gas, flue gas recycle stream (process gas IV), oxygen and optionally hydrogen 17 Heating gas mixture with optional hydrogen 18 Optional hydrogen for reactor heating 19 Hydrogen for oxygenate synthesis 20 Hydrogen cyanide-containing process gas 21 Process gas after sulfuric acid scrubbing 22 Pure hydrogen cyanide 23 Hydrogen-rich residual gas from hydrogen cyanide production 24 Optional purified hydrogen-rich residual gas 25 Hydrogen-rich residual gas 26 Raw ACH 27 Rein-ACH 28 CO2-containing exhaust gas from BMA for oxygenate synthesis 29 CO2-containing exhaust gas after deoxygenation 30 Gas mixture for oxygenate synthesis 31 Optional reactor recycle stream oxygenate synthesis 32 Optional hydrogen from oxygenate synthesis 33 Raw oxygenate 34 Optionally discharged oxygenate 35 Acetone from oxygenate synthesis for ACH synthesis 36 Methanol from oxygenate synthesis for MMA synthesis or methyl mercaptan synthesis 37 Derivatives from methyl mercaptan AG exhaust Fig. 2 (additional currents) 10 Fresh methanol 38 Mixture of natural gas and air and optionally hydrogen 39 Optionally discharged hydrogen Fig. 3 (additional devices and currents) O compressor P Methanol reactor Q separator R Pre-run column S Product column 40 Mixture of synthesis gas and crude methanol 41 Degassed crude methanol Fig. 4 (additional devices and currents) T CO generator U Acetone reactor V Pressure swing adsorption (PSA) W column 42 CO2-rich process gas 43 CO2-rich gas for acetone synthesis from PSA 44 Synthesis gas in acetone reactor 45 Gas mixture of CO and CO2 Fig. 5 (additional devices and currents) X Caphenia reactor 46 Mixture of natural gas and water Examples

[0055] Comparative Example V1: Production of hydrogen cyanide from methane and ammonia and subsequent conversion to methyl methacrylate (MMA) using non-process-coupled methanol

[0056] The synthesis of hydrogen cyanide is carried out according to the BMA process (details can be found in DE 1013636 B) as follows: The analysis of the reaction and process gas was carried out using online µGC, which has an accuracy of ± 0.2 %.

[0057] The hydrogen cyanide is produced in reactor A, which is spatially divided into two parts A1 and A2.

[0058] 636 kg / h of synthetic methane, with a typical composition of 95.1 wt% methane, 4.47 wt% hydrogen, 0.23 wt% carbon monoxide, < 0.1 wt% nitrogen, ethane and ethylene each, are fed to reactor section A1 with 724 kg / h of ammonia, with a typical composition of 99.5 wt% ammonia and 0.5 wt% water, and reacted.

[0059] Reactor A1 consists of 1,350 platinum-coated (according to DE 3923034, DE 19617040) reaction tubes with a length of 2,100 mm and an internal diameter of 16 mm, arranged in several chambers with up to 66 tubes each. The reactor operates at a temperature greater than 1,000 °C and a negative pressure of -50 mbarg. This temperature is generated by burning 486 kg / h of a mixture of commercially available natural gas with a methane content greater than 90% and hydrogen in a mass ratio of two to one with 9,315 kg / h of air in reactor section A2. This produces an exhaust gas stream (process stream II) of 9818 kg / h, containing 775 kg / h of CO2 and 2.5% residual oxygen, as well as 2000 mg / m3 of NOx, which is subsequently reduced to 300 mg / m3 by non-catalytic reduction with ammonia. The exhaust gas stream is then released into the atmosphere.

[0060] The process stream I leaving reactor A1 typically has a composition of 22 vol% hydrogen cyanide, 3 vol% ammonia, 1 vol% methane, and 74 vol% hydrogen. This stream is cooled to <200°C in a heat exchanger associated with the reactor and fed to the subsequent absorption stage I at a pressure of -25 mbar.

[0061] The ammonia is separated from process gas stream I by feeding it into absorption stage I, which is carried out as a two-stage acid scrubbing process, using 30 wt% sulfuric acid, thereby obtaining process gas stream II. 1100 kg / h of the resulting ammonium sulfate solution is discharged from the process.

[0062] By feeding the process gas stream II into the absorption stage II, the hydrogen cyanide contained in the process gas stream II is absorbed at a temperature of 8 °C and a pressure of -25 mbarg, with 24,000 kg / h

[0063] Water is washed out of the gas stream, producing 25,000 kg / h of 4 wt% aqueous hydrogen cyanide solution and 267 kg / h of a residual gas with a composition of 98% hydrogen, 1% methane, and 1% nitrogen. 75% of this stream is recycled to reactor section A2 for the heating described above.

[0064] The 25,000 kg / h of hydrogen cyanide solution is then separated by rectification into water, which is returned to absorption stage II, and hydrogen cyanide with a purity of >99%.

[0065] 1000 kg / h of hydrogen cyanide is produced, which corresponds to a yield of 98.4% based on the methane input. The production of 1000 kg / h of HCN requires 35 GJ / h.

[0066] The hydrogen cyanide is further processed into acetone cyanohydrin (ACH) in a second process step. For this, 1000 kg / h of HCN are mixed with 2140 kg / h of acetone and 2 kg / h of diethylamine and reacted in a four-stage reactor system at temperatures of 33 °C, 18 °C, 5 °C, and -5 °C, respectively. This yields 3650 kg / h of crude ACH, with a composition of 93% ACH, 2.5% acetone, and 2.5% HCN (residual inert by-products), which is then purified in a two-stage evaporator system. In the first stage, the low-boiling components acetone and HCN are separated and condensed as a vapor stream at 120 mbar and 82 °C, and in the second stage at 40 mbar and 85 °C. A distillate stream of 475 kg / h, with a composition of 83% ACH, 15% acetone, and 2% HCN, is obtained, which is recycled to reaction stage 1. A pure ACH stream (99.5%) of 3000 kg / h is produced at the bottom outlet of the evaporator.

[0067] In the third process step, ACH is converted to methacrylamide. For this purpose, 1834 kg / h of ACH are fed with 4941 kg / h of highly concentrated sulfuric acid at <100°C to the first amidation reactor. 550 kg / h of natural gas are required for combustion to produce the sulfuric acid. Together with the combustion of residual organic matter in the decomposition acid, this results in 2569 kg / h of CO2 emissions. A further 1166 kg / h of ACH are then added to the product stream, and the mixture is fed to a second amidation reactor. The mixture is then converted at 160°C with a residence time of several minutes. 7933 kg / h of the resulting amide mixture, containing 36 wt% methacrylamide, are then fed to a six-stage reactor cascade for esterification with 1125 kg / h of methanol and 1200 kg / h of water. After purification by crude and two-stage pure rectification, 3426 kg / h of methyl methacrylate are finally obtained with a yield of 91% based on the ACH used.

[0068] Based on the production of one kg of MMA, consumption figures or input factors for the individual raw materials are derived, as well as the emissions from sulfuric acid production and HCN synthesis (BMA) (see Table 1).

[0069] The following life cycle assessment is intended to evaluate the environmental impacts of the produced MMA with regard to CO2 emissions using the cradle-to-gate system boundary - the so-called Product Carbon Footprint (PCF). This takes into account not only the direct CO2 emissions from the process (Scope 1) but also the CO2 emissions from the raw material upstream chains (Scope 3). The indirect CO2 emissions from the energy (Scope 2) are neglected, since the production of electrical power is assumed to come from purely renewable sources and steam is produced as a heat source from the exothermic energy or waste heat of the MMA process itself. All CO2 emissions are related to 1 kg of MMA produced (functional unit FU). The data source for the primary data can be found in the examples; the secondary data, i.e. the specific CO2 emissions of the raw materials, come from PlasticsEurope (EcoProfiles).For methanol, a value from the literature is used (Methanol - Managing greenhouse gas emissions in the production chain by optimizing the resource base; AIMS Energy; DOI: 10.3934 / energy.2018.6.1074). The raw materials mentioned represent more than 99% of the total mass used. All other material flows are not considered here. This life cycle assessment was generally conducted according to ISO 14040 / 14044. For the environmental impact assessment, the impact assessment of the Dutch Research Center for Environmental Sciences Leiden (CML) with characterization factors from 2011, updated in April 2013, was used.

[0070] The quantity of raw materials and CO2 emissions released refer to the comparative example described above for 1000 kg / h hydrogen cyanide and 3424 kg / h MMA respectively.

[0071] Using the quantity and the specific emissions, the individual CO2 emissions of the raw materials can be calculated both specifically and absolutely. Table 1: Data basis and calculation of the product carbon footprint of MMA for the comparative example V1. PCF reactant / kgCO 2 eq. / kg Amount used per hour / kg CO2eq emissions kgCO2 eq. Upstream emissions methane 0,52 636 330 Natural gas (BMA + H2SO4) 0,52 826 429 276 +550 NH3 2,4 724 1737 acetone 1,6 2140 3424 Methanol (natural gas based) 0,83 1152 956 H2 export 1,6 66,7 -106,8 Subtotal 6769 production Direct CO2 emissions (BMA + H2SO4) 3344 (775+2569) 3344 (775+2569) MMA 3424 10113 PCF 2.95 kgCO2eq. / kg MMA

[0072] The quotient of total CO2 emissions and MMA production results in a PCF of 2.95 kgCO2eq. / kg MMA.

[0073] Table 1 shows that approximately two-thirds of the CO2 emissions from MMA production are attributable to the raw materials. 23% of the direct emissions come from the BMA process and 76% from the sulfuric acid plant. Inventive example E1: Production of hydrogen cyanide and methanol from methane and ammonia and their subsequent conversion to methyl methacrylate (MMA) with reduced CO2 emissions.

[0074] In Example E1 according to the invention, analogous to Comparative Example C1, 636 kg / h of synthetic methane, with a typical composition of 95.1 wt% methane, 4.47 wt% hydrogen, 0.23 wt% carbon monoxide, <0.1 wt% nitrogen, ethane, and ethylene, are fed into reactor section A1 with 724 kg / h of ammonia, with a typical composition of 99.5 wt% ammonia and 0.5 wt% water. Analogously, 1000 kg / h of hydrogen cyanide and 267 kg / h of hydrogen-rich residual gas are produced.

[0075] To heat the reactor, a 300 to 400 °C hot flue gas recycle stream of 3881 kg / h, consisting of 55 wt% CO 2 , 44 wt% H 2 O, and 1 wt% O 2 , is thermally reacted with 2751 kg / h of technically pure oxygen and 690 kg / h of commercially available natural gas with a methane content of > 90 wt% in reactor section A2. This produces 7323 kg / h of flue gas, of which 53% is continuously separated and used for the described flue gas recycle stream. The remaining part of the stream is deoxygenated using a precious metal catalyst and a carbon-containing reducing agent. After cooling to 20 °C and separating the condensed water, a stream of 2019 kg / h is obtained with a CO2 content of 98 wt% and < 25 mg / m3 NOx (residual water). In this context, process stream IV, including flue gas, is listed as described in the present example.

[0076] In the inventive example E1, in a next process step, this CO 2 stream is mixed with the 267 kg / h of residual gas with an H 2 content of 98 wt.% obtained from reactor section A1 and compressed to 85 bar by multi-stage compression. Between the compression stages, condensed water is separated from the gas phase after cooling. The composition of the mixed gas is then 72.1 vol.% H 2 , 24.3 vol.% CO 2 , and 3.6 vol.% H 2 O. This corresponds to a stoichiometry of 1.96. (The stoichiometry of methanol synthesis is defined as: SN = (H 2 - CO 2 ) / (CO + CO 2 ) in mol.)

[0077] The mixed and compressed synthesis gas is then fed into a first heat exchanger together with the recycle gas. The inlet temperature is 67 °C and the outlet temperature is 220 °C. The preheated gas has a composition of 77.2 vol% H 2 , 20.3 vol% CO 2 , 2.1 vol% CO, 0.28 vol% MeOH and 0.11 vol% H 2 O with a volume flow of 19061 Nm 3 / h. The gas is then fed to the tube side of a tube bundle reactor, where the reaction begins with heat evolution on a copper / zinc / aluminum oxide catalyst (e.g. Clariant MegaMax MM800) in pellet form (6 x 4 mm). The generated heat is carried away by the boiling water at 40 bar pressure on the shell side. The single-pass conversions for H2 and CO2 are 20% and 25%, respectively, with a residence time of 8.7 seconds and a catalyst bed porosity of 0.38. The saturated steam produced is 500 kg / h.The reacted reactor outlet gas has an outlet temperature of 251 °C with a composition of 69 vol% H 2 , 16.9 vol% CO 2 , 5.8 vol% MeOH, 5.7 vol% H 2 O and 2.3 vol% CO and a volume flow rate of 17163 Nm 3 / h. The gas is then passed into the first heat exchanger and cooled to 130 °C. In the final cooler, the process gas is cooled to 40 °C using cooling water (with a flow temperature of 30 °C), during which liquid products condense out. The liquid phase (so-called crude methanol) has a composition of 62.9 wt% MeOH, 35.9 wt% H 2 O and 1.12 wt% dissolved CO 2 at 83.7 bar. Other by-products include 700 ppm w / w EtOH and 50 ppm w / w dimethyl ether. The total mass flow is 2150 kg / h.

[0078] After separating the liquid and gas phases, the unreacted gas is remixed with the synthesis gas using a recycle compressor. The volume flow of the recycle gas is 15,237 Nm³ / h. The ratio of recycle gas to synthesis gas is 3.8. To prevent the accumulation of inert components, a purge gas stream is branched off from the recycle gas at a mass flow rate of 49 kg / h. The total conversion of H² is 98%. The liquid crude methanol is expanded to 3 to 5 bar in a crude methanol tank, releasing dissolved gases (primarily CO²).

[0079] In the subsequent pre-end column, further gases (DME) are separated from the crude methanol at ambient pressure and a bottom temperature of 75 °C. The top temperature is 65.4 °C with a reflux ratio (mol / mol) of 1.

[0080] All exhaust gases from the purge gas stream, raw methanol tank, and pre-run column are combusted together with air in a common combustion chamber to produce CO2 and H2O. The resulting CO2 emissions amount to 75 kg / h.

[0081] The stabilized crude methanol obtained in the bottoms is rectified in the subsequent product column at ambient pressure, yielding H2O in the bottoms at a temperature of 99 °C, a composition of 0.11 wt% MEOH, 376 wt ppm EtOH, and a total mass flow of 676 kg / h. The finished methanol product is obtained at the top of the column at a temperature of 64 °C and an EtOH content of 8.2 wt ppm. The production rate is 1410 kg / h. The column's reflux ratio is 3.4. To stabilize the column and ensure the product and wastewater specifications, by-products (mainly EtOH) with a composition of 92.5 wt% H 2 O, 4.1 wt% MeOH and 3.3 wt% EtOH and a total mass flow of 32.9 kg / h are separated by a sidestream separation in the lower part of the column.

[0082] The methanol produced will be reacted with methacrylamide and water to form methyl methacrylate, analogous to comparative example V1. Table 2: Data basis and calculation of the Product Carbon Footprint of MMA for example E1 PCF reactant / kgCO 2 eq. / kg Amount used per hour / kg CO2eq emission / kgCO 2eq / h Upstream emissions methane 0,52 636 330 Natural gas BMA + H2SO4 0,52 1240 (690 +550) 644 NH3 2,4 724 1737 acetone 1,6 2140 3424 Methanol (natural gas based) 0,83 -258 (1152-1410) -214 O 2 (gaseous) 0 2751 0 Subtotal 5921 H2 export 1,6 0 0 production Direct CO2 emissions (H2SO4 + MeOH) 2644 (2569+75) 2644 MMA 3424 8565 PCF 2.5 kgCO2eq. / kg MMA

[0083] The quotient of total CO2 emissions and MMA production results in a PCH of 2.5 kgCO2eq. / kg MMA.

[0084] By converting direct emissions from hydrogen cyanide with the hydrogen produced there, not only can direct emissions be reduced by 21%, but the Scope 3 share is also reduced by 5%. The excess methanol production is calculated using a CO2 credit. The increased use of natural gas partially offsets the credit from methanol production. Oxygen production via cryogenic air separation is powered by renewable electrical energy, so no (indirect) emissions are generated.

[0085] In the end, the example can show that CO2 emissions could be reduced by 15% compared to the comparison example. Inventive example E2a: Production of hydrogen cyanide and methanol from methane and ammonia and their subsequent conversion to methyl methacrylate (MMA) under improved process conditions.

[0086] In inventive example E2a, analogous to inventive example E1, 636 kg / h of synthetic methane, with a typical composition of 95.1 wt% methane, 4.47 wt% hydrogen, 0.23 wt% carbon monoxide, and < 0.1 wt% each of nitrogen, ethane, and ethylene, are fed into reactor section A1 with 724 kg / h of ammonia, with a typical composition of 99.5 wt% ammonia and 0.5 wt% water. Analogously, 1000 kg / h of hydrogen cyanide and 267 kg / h of hydrogen-rich residual gas are produced.

[0087] The reactor is heated by substoichiometric combustion of 690 kg / h of commercial natural gas, with a methane content of over 90%, with 2688 kg / h of oxygen and the addition of 3809 kg / h of a 300 to 400 °C hot recycle stream, consisting of 55 wt% CO 2 , 44 wt% H 2 O and 1 wt% CO, in reactor section A2.

[0088] This produces 7187 kg / h of flue gas, 53% of which is separated and used for the described flue gas recycling stream. After cooling to 20 °C and separation of condensed water, 1942 kg / h of a carbon monoxide stream with a CO 2 content of 97 wt.%, a CO content of 1 wt.%, an oxygen content of < 1000 wt. ppm, and a NOx content of < 25 mg / m 3 are obtained. In this context, flue gas refers to process stream IV as described in the present example.

[0089] Analogously to Example E1 according to the invention, the carbon oxide stream is then reacted with the hydrogen-containing residual gas stream to form methanol, which, like the hydrogen cyanide produced, is then used for the production of methyl methacrylate.

[0090] The methanol production increases by 2% compared to example E1 to 1440 kg / h with the same H 2 input.

[0091] The individual mass flows are summarized in Table 3 for the analysis of the CO 2 footprint Table 3: Data basis and calculation of the Product Carbon Footprint of MMA for the example E2a PCF reactant / kgCO 2 eq. / kg Amount used per hour / kg CO2eq emission / kgCO 2eq / h Upstream emissions methane 0,52 636 330 natural gas 0,52 1240 (690 +550) 644 NH3 2,4 724 1737 acetone 1,6 2140 3424 Methanol (natural gas based) 0,83 -288 -239 1152-1440 O 2 (gaseous) 0 2688 0 H2 export 1,6 0 0 Subtotal 5896 output Direct CO2 emissions (H2SO4 + MeOH) 2645 (2569+76) 2645 (2569+76) MMA 3424 8541 PCF 2.49 kgCO2eq / kg MMA

[0092] The quotient of total CO2 emissions and MMA production results in a slightly lower PCF of 2.49 kgCO2eq / kg MMA. Inventive example E2b: Production of hydrogen cyanide and methanol from methane and ammonia and their subsequent conversion to methyl methacrylate (MMA) under improved process conditions.

[0093] In inventive example E2b, analogous to inventive example E1, 636 kg / h of synthetic methane, with a typical composition of 95.1 wt% methane, 4.47 wt% hydrogen, 0.23 wt% carbon monoxide, and < 0.1 wt% each of nitrogen, ethane, and ethylene, are fed into reactor section A1 with 724 kg / h of ammonia, with a typical composition of 99.5 wt% ammonia and 0.5 wt% water. Analogously, 1000 kg / h of hydrogen cyanide and 267 kg / h of hydrogen-rich residual gas are produced.

[0094] By further reducing the C / O ratio for heating the reactor while maintaining the reaction heat or temperature, it was surprisingly possible to carry out parallel synthesis gas production (analogous to a partial oxidation).

[0095] Here, the reactor is heated analogously to example E2a by substoichiometric combustion of 1467 kg / h of commercially available natural gas, with a methane content of over 90%, with 4032 kg / h of oxygen with the addition of 3809 kg / h of a 300 to 400 °C hot recycle stream, consisting of 16 wt% CO2, 42 wt% H2O, 17 wt% CO and 24 vol% H2 as well as residues of methane, in reactor section A2.

[0096] This produces 13,746 kg / h of flue gas or synthesis gas, 40% of which is separated and used for the described recycle stream. After cooling to 20 °C and separation of condensed water, 3,450 kg / h of a synthesis gas stream with a composition (in vol%) of 41% H2, 29% CO, 27% CO2, and 2.3% H2O, as well as traces of methane, is obtained. In this context, flue gas refers to process stream IV as described in the present example.

[0097] The mixture, together with hydrogen-containing residual gas, produces a synthesis gas with a composition of 67 vol% H2, 16 vol% CO, 14.9 vol% CO2, as well as residues of H2O and methane. Analogous to inventive example E1, the synthesis gas stream is then converted to methanol, which, along with the hydrogen cyanide produced, is used for the production of methyl methacrylate.

[0098] Methanol production was significantly increased to 2550 kg / hr compared to Example E1. This corresponds to an increase of 77%. The excess methanol production amounts to 1398 kg / hr.

[0099] The individual mass flows are summarized in Table 3 for the analysis of the CO 2 footprint Table 4: Data basis and calculation of the Product Carbon Footprint of MMA for the example E2b PCF reactant / kgCO 2 eq. / kg Amount used per hour / kg CO2eq emission / kgCO 2eq / h Upstream emissions methane 0,52 636 330 natural gas 0,52 2017 (1467 +550) 1048 NH3 2,4 724 1737 acetone 1,6 2140 3424 Methanol (natural gas based) 0,83 -1398 -1160 1152-2550 O 2 (gaseous) 0 4032 0 H2 export 1,6 0 0 Subtotal 5380 output Direct CO2 emissions (H2SO4 + MeOH) 2645 (2569+76) 2645 (2569+76) MMA 3424 8024 PCF 2.34 kgCO2eq / kg MMA

[0100] The results of the CO2 emissions analysis surprisingly show that, despite the higher natural gas volumes required to provide thermal energy for the BMA reaction and for syngas production within the BMA reactor's combustion chamber, methanol production has a positive impact on the PCF of the MMA. Assuming that the methanol overproduction can be compensated with a conventional CO2 factor, this results in savings of 1160 kg / hr CO2eq. The final PCF for this example is 2.34 kgCO2eq / kg MMA, corresponding to a reduction of 22%. Inventive example E3: Production of hydrogen cyanide and acetone from methane and ammonia and their subsequent conversion to methyl methacrylate (MMA) with reduced CO2 emissions.

[0101] In inventive example E3, analogous to inventive example E1, 636 kg / h of synthetic methane, with a typical composition of 95.1 wt% methane, 4.47 wt% hydrogen, 0.23 wt% carbon monoxide, <0.1 wt% nitrogen, ethane, and ethylene, are fed into reactor section A1 with 724 kg / h of ammonia, with a typical composition of 99.5 wt% ammonia and 0.5 wt% water. Analogously, 1000 kg / h of hydrogen cyanide and 267 kg / h of hydrogen-rich residual gas are produced.

[0102] The reactor is heated analogously to Example E1 according to the invention. After cooling to 20 °C and separation of condensed water, a flow of 2019 kg / h is obtained with a CO 2 content of 98 wt% and < 25 mg / m 3 < NO x . The remainder is predominantly water.

[0103] 62% of this CO2-rich gas stream is converted into a CO2-rich stream by electrical power in a CO2 generator based on a SOEC (Solid Oxide Electrolysis Cell) (see article "Small-Scale CO from CO2 using Electrolysis" in "Chemical Engineering World; pages 44-46; March 2017). The composition of this product stream is 99 vol% CO and 1 vol% CO2 with a volume flow of 778 kg / h. 438 kg / h of oxygen is also produced. The electrical power consumption is 4.4 MW.

[0104] The CO2-rich stream is then blended with a portion (5%) of the hydrogen-rich residual gas from the BMA plant, the bypass of the carbon dioxide stream, and the CO2 recycle gas stream to produce a synthesis gas with a composition of 50 vol% CO, 30 vol% CO2, 10 vol% H2, and 10 vol% N2. The majority of the H2 stream (247 kg / h) is used in the combustion of the H2SO4 plant to replace the natural gas there, thus avoiding the (direct) generation of any hydrogen.

[0105] The mixed synthesis gas is then converted to acetone in a bioreactor according to the procedure described in the literature (US 2012 / 0252083, FE Liew et al. Carbon-negative production of acetone and isopropanol by gas fermentation at industrial pilot scale, Nature Biotechnology 40, 335-344 (2022)). The synthesis gas described above is fed at 1568 Nm³ / h at 35 °C into a continuously operated stirred tank (V=550 m³) at 1.4 bar, containing modified Clostridium autoethanogenum.

[0106] The straight-pass conversion based on CO is 80%, and the selectivity of CO to acetone is 80%. The product stream will be 237 kg / h, consisting of 80 wt% acetone, 15 wt% acetoacetic acid, 2.5 wt% ethanol, and 2.5 wt% isopropanol.

[0107] This mixture is then purified by rectification. The overhead product is 194 kg / h of acetone with a purity of 99.5%. The bottom product, 43 kg / h (35 wt% acetone, 50 wt% acetoacetic acid, 7 wt% ethanol, and 8 wt% isopropanol), is discharged and processed in a further process step.

[0108] The produced acetone is converted together with 1946 kg / h of conventional acetone, analogous to comparative example V1, first to ACH and then further to methacrylamide and water as well as to methyl methacrylate.

[0109] The unreacted gas is first compressed to 15 bar using a compressor before being fed into a pressure swing adsorption plant. Here, the CO is recovered at an efficiency of 90% and fed back into the bioreactor as a recycled gas stream (246 Nm³ / h) with a composition of 69 vol% CO and 30 vol% CO². The tail gas stream from the PSA, with a composition of 76 vol% CO², 17.8 vol% N², and 2.1 vol% CO² and a volume flow of 873 Nm³ / h, is thermally utilized in the H²SO² plant. Table 5: Data basis and calculation of the Product Carbon Footprint of MMA for the example E3 PCF reactant / kgCO 2 eq. / kg Amount used per hour / kg Product carbon footprint / kgCO 2 eq. Upstream emissions methane 0,52 636 330 natural gas 0,52 1240 644 NH3 2,4 724 1737 acetone 1,6 1946 3113 Methanol (natural gas based) 0,83 1152 954 O 2 (gaseous) 0 2250 (2699 - 438) 0 H2 export 1,6 0 0 Partial result 6778 output Direct CO2 emissions (H2SO4+acetone) 2353 (1324+1029) 2353 MMA 3424 9131 PCF 2.66 kgCO2.eq / kg MMA

[0110] The PCF is reduced to 2.66 kgCO2.eq / kg MMA in this example, which corresponds to a reduction of 10% compared to the comparison example. Inventive example E4: Production of hydrogen cyanide and acetone from methane and ammonia and their subsequent conversion to methyl methacrylate (MMA) with reduced CO2 emissions.

[0111] In inventive example E3, analogous to inventive example E1, 636 kg / h of synthetic methane, with a typical composition of 95.1 wt% methane, 4.47 wt% hydrogen, 0.23 wt% carbon monoxide, <0.1 wt% nitrogen, ethane, and ethylene, are fed into reactor section A1 with 724 kg / h of ammonia, with a typical composition of 99.5 wt% ammonia and 0.5 wt% water. Analogously, 1000 kg / h of hydrogen cyanide and 267 kg / h of hydrogen-rich residual gas are produced.

[0112] The reactor is heated analogously to Example E1 according to the invention. After cooling to 20 °C and separation of condensed water, a flow of 2019 kg / h is obtained with a CO 2 content of 98 wt% and < 25 mg / m 3 < NO x . The remainder is predominantly water.

[0113] The CO2-rich gas stream is converted into a CO2-rich stream by electrical power in a CO2 generator based on an SOEC (Solid Oxide Electrolysis Cell) (see article "Small-Scale CO from CO2 using Electrolysis" in "Chemical Engineering World; pages 44-46; March 2017). The composition of this product stream is 99 vol% CO and 1 vol% CO2 with a mass flow of 1282 kg / h. A further 737 kg / h of oxygen is produced, which can then be added for combustion in the BMA reactor. The electrical power consumption is 7.4 MW.

[0114] The CO2-rich stream is then mixed with the hydrogen-rich residual gas from the BMA plant to form a synthesis gas with a composition of 61 vol% H2, 37 vol% CO, 0.4 vol% CO2 and the remainder H2O and N2. This mixed synthesis gas is then converted to acetone in a bioreactor according to the process described in the literature (US 2012 / 0252083, FE Liew et al. Carbon-negative production of acetone and isopropanol by gas fermentation at industrial pilot scale, Nature Biotechnology 40, 335-344 (2022)). The synthesis gas described above is fed at 1549 kg / hr at 35 °C into one or more continuously operated stirred tanks with a total volume of 1400 m3< at 1.4 bar, which contains modified Clostridium autoethanogenum. The conversion based on CO or H2 is 80% and the selectivity of CO to acetone is 80%.After cooling the product stream to 5°C and separating the unreacted gaseous reactants (H2 and CO), the process water is first separated by distillation. This yields a product stream of 385 kg / h, consisting of 80 wt% acetone, 15 wt% acetoacetic acid, 2.5 wt% ethanol, and 2.5 wt% isopropanol.

[0115] This mixture is then purified by further rectification. The overhead product is 308 kg / h of acetone with a purity of 99.5%. The bottom product, 77 kg / h (35 wt% acetone, 50 wt% acetoacetic acid, 7 wt% ethanol, and 8 wt% isopropanol), is discharged and processed in a further process step.

[0116] The unconverted synthesis gas, with a composition of 75 vol.% H2, 19 vol.% CO, and residual H2O and CO2, with a mass flow of 864.3 kg / hr, is thermally converted in the sulfuric acid plant, where it replaces 99% of the natural gas. Due to the higher H / C ratio of the purge gas compared to natural gas, the direct CO2 emissions from H2SO4 production are reduced by 14%.

[0117] The produced acetone is combined with 1795 kg / h of conventional acetone in the same way as

[0118] Comparative example V1 was first converted to ACH and then further to methacrylamide and water as well as to methyl methacrylate. Table 6: Data basis and calculation of the Product Carbon Footprint of MMA for example E4 PCF reactant / kgCO 2 eq. / kg Amount used per hour / kg Product carbon footprint / kgCO 2 eq. Upstream emissions methane 0,52 636 330 natural gas 0,52 782 406 NH3 2,4 724 1737 acetone 1,6 1832 (2140 - 308) 2931 Methanol (natural gas based) 0,83 1152 954 O 2 (gaseous) 0 1951 (2688-737) 0 H2 export 1,6 0 0 Partial result 6358 output Direct CO2 emissions 2357 2357 MMA 3424 8715 PCF 2.54 kgCO2eq. / kg MMA

[0119] This example demonstrates that the CO2 footprint of the entire MMA process can be further reduced by optimizing the conditions for the biogenic production of acetone. The PCF is 2.54 kgCO2eq. / kg MMA, corresponding to a reduction of 14%. Inventive example E5: Production of hydrogen cyanide and acetone from methane and ammonia and their subsequent conversion to methyl methacrylate (MMA) with reduced CO2 emissions.

[0120] In inventive example E5, analogous to inventive example E1, 636 kg / h of synthetic methane, with a typical composition of 95.1 wt% methane, 4.47 wt% hydrogen, 0.23 wt% carbon monoxide, <0.1 wt% nitrogen, ethane, and ethylene, are fed into reactor section A1 with 724 kg / h of ammonia, with a typical composition of 99.5 wt% ammonia and 0.5 wt% water. Analogously, 1000 kg / h of hydrogen cyanide and 267 kg / h of hydrogen-rich residual gas are produced.

[0121] The reactor is heated analogously to Example E1 according to the invention. After cooling to 20 °C and separation of condensed water, a flow of 2019 kg / h is obtained with a CO 2 content of 98 wt% and < 25 mg / m 3 < NO x . The remainder is predominantly water.

[0122] The CO2-rich gas stream is then fed into a partial oxidation unit with an integrated electric plasma arc (see Caphenia patents US 10,906,806 and US 10,927,007). Combined with 1101 kg / hr of natural gas and 1198 kg / hr of steam (200°C) at a temperature of 1000°C and 20 bar, a synthesis gas containing 58 vol.% H2, 34 vol.% CO, and 4.9 vol.% CO2 (the remainder being methane) is produced. With its H2 to CO ratio of 5 to 3, the resulting synthesis gas is ideal for the production of acetone. The power required to generate the plasma is 7.1 MW.

[0123] The acetone production process is described analogously to the process in Example E4. Additionally, in this example, the unreacted purge gas is recirculated to the bioreactor using a blower. This increases the overall CO conversion to 95% while maintaining a constant CO to acetone selectivity of 80%. After product purification (analogous to Example E4), a product flow of 1276 kg / hr of acetone is obtained.

[0124] The resulting purge with 958 kg / hr and a composition of 30 vol.% H2, 36 vol.% CO2, 13 vol.% and 12 vol.% methane is fed into the combustion of the H2SO4 plant and replaces 40% of the natural gas for the auxiliary combustion.

[0125] A portion of the hydrogen-rich electricity (134 kg / hr, 50%) from the BMA reactor is also used to fuel the H2SO4 plant, replacing the remaining 60% of the natural gas. The other half, 134 kg / hr, is intended for export and is remunerated with a CO2eq factor of 1.6 kgCO2eq / kg.

[0126] The produced acetone is converted together with 864 kg / h of conventional acetone, analogous to comparative example V1, first to ACH and then further to methacrylamide and water as well as to methyl methacrylate. Table 7: Data basis and calculation of the product carbon footprint of MMA for example E5. PCF reactant / kgCO 2 eq. / kg Amount used per hour / kg Product carbon footprint / kgCO 2 eq. Upstream emissions methane 0,52 636 330 natural gas 0,52 1791 931 NH3 2,4 724 1737 acetone 1,6 864 (2140 - 1276) 872 Methanol (natural gas based) 0,83 1152 956 O 2 (gaseous) 0 2751 0 H2 export 1,6 134 -214 Partial result 5122 output Direct CO2 emissions 1931 1931 MMA 3424 6919 PCF 2.06 kgCO2eq. / kg MMA

[0127] This example shows that the co-reforming of CO2 with natural gas, steam, and electricity significantly reduces the PCF. 2.06 kgCO2eq. / kg MMA,which corresponds to a reduction of 31% compared to the comparison game.

[0128] Summary of results: (see Table 8)

[0129] The results presented in Table 8 for the comparative and patent examples show that all examples were able to achieve the CO2 reduction target for Scope 1 and Scope 3 emissions, despite increased natural gas demand. Since the oxygen production and the conversion of CO2 to CO2 in examples E3 to E5 are based on largely CO2-free electrical energy, the Scope 2 emissions of the energy sources can be neglected here.

[0130] Furthermore, it can be seen that the air-free or N2-free operation of the BMA reactor furnace achieves a significant reduction in NOx emissions, which would otherwise only be possible with additional ammonia denitrification. Table 8: Results according to comparative example V1 and inventive examples E1-E5. Selected process parameters for hydrogen cyanide production Comparison example V1 Examples according to the invention E1 E2a E2b E3 E4 E5 HCN production in kg / h 1000 1000 1000 1000 1000 1000 1000 CO 2 emissions (BMA process) / kg / h 747 0 0 0 0 0 0 Specific energy demand GJ / t HCN 35 28 28 60 28 28 28 NO x emission / mg / m 3 300 20 18 18 20 20 20 Natural gas consumption in kg / h 334 690 690 1467 690 690 1791 Air requirement in kg / h 9315 0 0 0 0 0 0 Pure O 2 requirement in kg / h 0 2751 2688 4032 2250 1951 2751 Residual oxygen concentration of flue gas (process gas IV) in % 2,5 1 0 0 1 0 1 Selected process parameters for H 2 SO 4 production H 2 SO 4 production in kg / h 4941 4941 4941 4941 4941 4941 4941 Natural gas consumption in kg / h 550 550 550 550 0 18 0 in kg / h 2569 2569 2569 2569 1029 2357 1931 Selected process parameters for MMA production MMA production (>99.85%) in kg / hr 3426 3426 3426 3426 3426 3426 3426 Methanol demand for MMA production in kg / h 1125 1125 1125 1125 1125 1125 1125 Acetone requirement for MMA production in kg / h 2140 2140 2140 2140 2140 2140 2140 Selected process parameters for oxygenate production Acetone from BMA exhaust gas in kg / h 0 0 0 0 194 308 1276 Methanol from BMA exhaust gas in kg / h 0 1410 1440 2550 0 0 0 CO2 emissions from oxygenate production / kg / h 0 75 75 75 1324 0 0 Product Carbon Footprint PCF MMA in kg CO 2 q / kg MMA 2,98 2,5 2,49 2,35 2,66 2,54 2,06

Claims

1. Process for preparing hydrogen cyanide and organic compounds including one to three carbon atoms and oxygen, characterized in that the process comprises the following process steps: I. preparing a mixture of hydrogen cyanide and hydrogen as process gas I in a reaction zone A1 of a reactor A, proceeding from ammonia and a methane-containing fraction, in particular natural gas, II. separating process gas I into a process gas II consisting predominantly of hydrogen cyanide and a process gas III consisting predominantly of hydrogen, III. producing a process gas IV consisting predominantly of CO2 in a reaction zone A2 of reactor A, proceeding from a methane-containing fraction, IV. mixing at least a substream of process gas III with at least a substream of process gas IV to obtain a process gas V, and V. converting process gas V in a reactor B to produce a process stream VI including an organic compound.

2. Process according to Claim 1, characterized in that, in a process step VI, at least a portion of the hydrogen cyanide obtained in process step I and at least a portion of the organic compound obtained in process step V are used to prepare a joint conversion product.

3. Process according to Claim 1 or 2, characterized in that process gas IV is obtained by the reaction of a methane-containing fraction, in particular natural gas, with oxygen, in that a substream of this process gas IV, as process gas IVa, is returned directly to reaction zone A2 and optionally mixed first with oxygen and then optionally with further methane-containing fraction before this gas mixture is returned to reaction zone A2 as process gas IVb.

4. Process according to at least one of Claims 1 to 3, characterized in that process step I is effected with substantial exclusion of oxygen in a reactor A in a reaction zone A1 having a platinum-containing catalyst, and in that the process gas I is treated before process step II with an acidic aqueous absorbent in an absorption apparatus I in such a way that remaining ammonia is separated off and obtained in the form of an aqueous ammonium sulfate-containing solution.

5. Process according to Claim 4, characterized in that the ammonia-depleted process gas II is treated in at least one absorption apparatus II with an aqueous, optionally carbonyl-containing absorbent so as to obtain an aqueous HCN-containing solution, and in that this solution is then treated in at least one distillation column I in such a way that a liquid product fraction I containing at least 99% by weight of hydrogen cyanide is obtained after condensation.

6. Process according to at least one of Claims 1 to 5, characterized in that process gas V is a synthesis gas having a molar ratio of CO and CO2 to hydrogen between 1:1 and 1:10.

7. Process according to at least one of Claims 1 to 6, characterized in that process gas III has a hydrogen content > 90% by volume and in that this content is optionally increased between process step II and process step IV to > 95% by volume in one or more workup steps, and in that the optionally enriched process gas III is obtained at a pressure < 5 bara.

8. Process according to at least one of Claims 1 to 7, characterized in that the methane-containing fractions are natural gas, where a mixture of oxygen and process gas IVa, optionally with hydrogen, is fed into the methane-containing fractions in or upstream of reaction zone A2, and in that process step III serves to provide thermal energy for the endothermic reaction of process step I.

9. Process according to at least one of Claims 1 to 8, characterized in that the process gas IV after exiting from reaction zone A2 is at a temperature between 700°C and 1300°C and has an oxygen content of less than 5% by volume, an NOx content of less than 500 ppm by volume and a total content of CO, CO2 and water of 50% to 99% by volume, and in that process gas IV is then cooled to a temperature between 150°C and 450°C to produce mid-pressure steam.

10. Process according to at least one of Claims 1 to 9, characterized in that process stream VI has an oxygen content of less than 0.1% by volume, preferably less than 0.01% by volume.

11. Process according to at least one of Claims 1 to 10, characterized in that one substream of process gas IV is guided into reaction zone A2 for process step III, and a second substream of process gas IV into reactor B for process step V.

12. Process according to at least one of Claims 1 to 11, characterized in that process gas IV before process step IV, or a substream of process gas IV guided into reactor B at a later stage, or process gas V before process step V, is subjected to deoxygenation, preferably adiabatically in contact with a transition metal catalyst or precious metal catalyst at a temperature between 200 and 800°C.

13. Process according to at least one of Claims 2 to 12, characterized in that the organic compound which is formed in process step V is methanol, where the conversion in process step V is effected by conversion in the gas phase over a heterogeneous catalyst at pressures of 20 to 120 bar and a temperature between 130°C and 350°C, and wherein, after process step V, process stream VI is separated by condensation into a water phase, a phase consisting predominantly of methanol, and an uncondensed process gas VII.

14. Process according to at least Claim 13, characterized in that process gas VII is mixed with process gas V, and in that reactor B comprises a plurality of series-connected subreactors, preferably shell-and-tube reactors and / or tray reactors, with achievement of a partial conversion between 20% and 90% in the respective subreactors, and where the respective gas mixture is cooled before entering the respective next subreactor.

15. Process according to Claim 13 or 14, characterized in that the hydrogen cyanide prepared in process step I and the methanol obtained in process step V are used to prepare methyl methacrylate in process step VI, where the methanol has been added to a mixture comprising sulfuric acid, ammonium hydrogensulfate, methacrylamide and water at a temperature between 80°C and 130°C and the methacrylamide has been prepared beforehand using the hydrogen cyanide.

16. Process according to Claim 13 or 14, characterized in that the methanol obtained in process step V together with hydrogen sulfide that has optionally been prepared using a substream of process gas II are first used to obtain methyl mercaptan in process step VI, and the methyl mercaptan is then reacted in a subsequent step with the hydrogen cyanide produced in process step I to give methionine.

17. Process according to at least one of Claims 2 to 12, characterized in that process step V is a fermentation and the organic compound is acetone, and in that the conversion product in process step VI is methacrylamide, from which an alkyl methacrylate or methacrylic acid is prepared in turn, where the acetone together with hydrogen cyanide is contacted with a basic catalyst at a temperature between 0°C and 60°C, and the acetone cyanohydrin thus obtained, after purification, is reacted with sulfuric acid, methanol and water in at least two stages to give a methacrylate.

18. Process according to at least one of Claims 2 to 12, characterized in that process step V is a fermentation and the organic compound is ethanol or propanol, and in that the conversion product in process step VI is ethyl methacrylate or propyl methacrylate which is obtained by esterification of a methacrylamide or methacrylic acid based on hydrogen cyanide from process step I or which is obtained by transesterification of an MMA based on hydrogen cyanide from process step I with the ethanol or propanol from process step V.

19. Process according to at least one of Claims 2 to 12, characterized in that process step V is a fermentation and the organic compound is acetone, and in that the conversion product in process step VI is isophoronediamine, where the isophoronediamine has been prepared by reacting hydrogen cyanide from process step I with isophorone that has been at least partly prepared from acetone from process step V.

20. Process according to Claim 13, characterized in that process stream VI is subjected to a further separation, where after the condensation and an optional workup is a fourth fraction which is an emission gas having a nitrogen oxide content of less than 150 ppm by volume.