Process and plant for methanol production

Abstract

A process for producing methanol from a regulated make-up gas includes the steps of regulating the stoichiometry of make-up gas fed to a methanol loop with a first stream of hydrogen recovered from a small portion of the make-up gas separated from the main stream, and a second stream of hydrogen recovered from loop refining; the process also includes the step of feeding at least a portion of the methane-rich waste gas to a reforming stage for use as a feedstock for making-up gas production.

Description

Process and plant for methanol production Description Field of application The present invention relates to a process and plant for the synthesis of methanol. Prior knowledge Industrial preparation of methanol includes synthesis gas preparation, methanol synthesis, and methanol purification. Synthesis gas is prepared in the front-end by subjecting a suitable hydrocarbon feedstock to a reforming process involving partial oxidation to produce a gaseous mixture of carbon oxides and hydrogen. Such a mixture is commonly called synthesis gas or make-up gas. Synthesis gas, after cooling and compression under methanol synthesis conditions, is fed to a methanol synthesis loop where crude methanol is synthesized in a gas-solid catalytic reactor over CuO / ZnO / Al2O3-based catalysts. Typical operating temperatures and pressures of the synthesis reactor are 250°C to 300°C and 50 to 100 bar. A methanol synthesis loop typically produces a stream of raw methanol and a purge gas exiting the loop. The purge gas contains unreacted hydrogen and inert compounds such as methane, argon, and nitrogen. It is known to recover unreacted hydrogen from the purge gas, for example using a hydrogen recovery unit (HRU). The HRU typically operates by membrane separation or pressure swing adsorption (PSA). The composition of the synthesis gas fed to the methanol reactor is characterized by the following stoichiometric number: R = (H2- CO2) / (CO + CO2). A stoichiometric number of about 2.02 to 2.1 is generally considered optimal for many catalysts used in methanol synthesis. Depending on the technology adopted, the synthesis gas produced in the front-end may have a sub-stoichiometric hydrogen content. This term indicates that the make-up gas contains less hydrogen than the target stoichiometric amount for methanol synthesis. In most cases, for example, a synthesis gas stream with a stoichiometric number R less than 2 is considered sub-stoichiometric. Typically, a make-up gas with such a hydrogen deficiency is produced from a stand-alone autothermal reformer fired with oxygen or oxygen-enriched air and operated with a relatively low steam-to-carbon ratio. This make-up gas production technique is attractive because of its low fuel consumption, however, the drawbacks of hydrogen deficiency must be overcome. A lack of hydrogen may lead to low efficiency and high production of undesirable by-products such as higher alcohols and ketones. Therefore, there is a need to adjust the hydrogen content of the make-up gas before it enters the methanol reactor. Hydrogen separated from the synthesis loop purge gas, as described above, can be used for this purpose. Unfortunately, due to the low purge gas content (i.e., maximum 5%), the hydrogen that can be recovered from the loop purge may not be sufficient to achieve the desired stoichiometry in the feed. Various attempts have been made in the prior art to solve this problem. US 6797252 discloses a process for producing synthesis gas in which the composition of the make-up gas is adjusted using a plurality of secondary reactors, wherein at least one of the secondary reactors is a low temperature water gas shift reactor and the other is a high temperature water gas shift reactor. In the prior art, provision is made for a low temperature water gas shift reactor arranged in series after a high temperature water gas shift reactor to recover the amount of hydrogen necessary to adjust the stoichiometric number of make-up gas feeds to the methanol reactor. The water-gas shift reactor converts carbon monoxide and water into hydrogen and carbon dioxide. The hydrogen is separated from the carbon dioxide through a CO2 removal unit and then fed to the methanol reactor to adjust the stoichiometry. However, this technique requires the installation of two expensive water gas shift units; another drawback is that the amount of hydrogen recovered in the WGS units is not constant over the lifetime of the installation, due to the partial deactivation of the catalyst over time, and the stoichiometry gradually deviates from the optimal value. Currently, there is no method or process available that allows adjusting and maintaining the stoichiometry of the syngas fed to the methanol reactor to the optimal value throughout the entire production cycle. Furthermore, conventional methanol synthesis processes have an additional drawback when the autothermal reformer is operated with a low steam-to-carbon molar ratio (S / C) and when relatively high pressures are considered, e.g. S / C less than 1.5 and pressures of about 35 to 50 bar, because under these operating conditions a relatively high amount of unconverted methane is retained in the synthesis gas. Such an unconverted stream is usually extracted from the methanol converter as a refinery gas and, after separation from hydrogen, can be recycled as a fuel to the fuel heater. However, if a large amount of such waste gases is valued as a fuel, the heating duty required for the process can be saturated by the waste gases alone, leading to problems with the fuel heater. As a result, not all methane-containing waste gases can be combusted optimally. Therefore, it is desirable to find a more efficient way to utilize the methane stream. Therefore, given the drawbacks mentioned above, it is desirable to design a methanol process that can exploit the unconverted methane stream in a more efficient manner while allowing for the adjustment of the stoichiometry of the syngas fed to the methanol reactor. Summary of the invention The invention is faced with the problem of how to deal with the high content of residual methane in the purge gas stream extracted from the methanol synthesis loop. Furthermore, the invention is faced with problems of how to adjust the hydrogen content in the make-up gas for methanol synthesis when the make-up gas is produced under sub-stoichiometric conditions. More specifically, the object of the invention is to provide a process in which the synthesis loop can be operated efficiently when the synthesis gas produced in the front section has a sub-stoichiometric hydrogen content and a significant methane content. Another objective is to provide a process in which the adjustment of the hydrogen content in the feed gas can be controlled in a flexible manner to adapt to different operating conditions, in particular in the case of aging and loss of efficiency of the HTS catalyst over time and the variability of the feedstock used as hydrocarbon source. Another objective is to develop a methanol plant with low capital expenditure (CAPEX) and low natural gas consumption. Still another objective is to develop a process and plant that is scalable to large capacities, e.g. 10,000 tonnes per day of methanol, under economically attractive conditions. Another objective is to reduce emissions so that a large plant is also environmentally acceptable. The above objectives are achieved by a process and a plant in accordance with the claims. In the process of the invention, a portion of the make-up gas obtained by reforming is separated during a cooling process and before the cooling process is completed, so that it still contains a significant amount of water. This separated semi-cooled make-up gas is subjected to water-gas displacement and hydrogen recovery, thus obtaining a first hydrogen stream. The remainder of the make-up gas is further cooled and the main make-up gas stream is completely cooled. A second hydrogen stream is obtained from the methanol synthesis loop refining. The first and second hydrogen streams are added to the main make-up gas stream to adjust its hydrogen content. The addition of the first hydrogen stream and the second hydrogen stream can be carried out at the same location or at different locations. In addition, a methane-rich stream is also obtained from loop refining, which is recycled as feedstock for the reforming process to produce make-up gas. The fraction separated from the make-up gas is a small fraction. Particularly preferably, this fraction may be (by volume) up to 15% of the make-up gas present in the reforming process, for example 1% to 10%. The invention provides a dual hydrogen recovery using hydrogen displacement and purification performed in the separated portion of the make-up gas and hydrogen recovery in the synthesis loop purification stream. This dual recovery allows for more flexible control in adjusting the hydrogen content of the main make-up gas stream. The separated stream of fresh make-up gas presents additional degrees of freedom that can be scaled up well to cope with changes in feed composition, providing good flexibility for feedstocks as well as flexibility from start of run (SOR) to end of run (EOR) conditions of the catalyst from HTS. In particular, an advantage of the present invention is that said dual hydrogen recovery is combined with the recovery of the methane-rich stream obtained from the synthesis loop refining. Displacement of the separated portion of fresh make-up gas allows for the recovery of useful heat due to the high temperature of said stream. The methane-rich stream recovered from the loop refining may be obtained at a pressure level compatible with the synthesis gas production section, especially when membrane separation is adopted. Accordingly, the methane-rich stream can be recycled without recompression. Preferably, the first hydrogen stream is obtained by a PSA process and the second hydrogen stream by membrane separation. Obtaining the first hydrogen stream by a PSA process has further advantages. The PSA process recovers a pure hydrogen stream that is kept at the same feed pressure, thus allowing it to be mixed with the remaining part of the synthesis gas without any recompression. Furthermore, the hydrogen stream produced by PSA has a very high purity, so that part of it can be sent for other uses, for example for the co-production of ammonia. The modification process may include autothermal modification, optionally with one or more pre-modification steps. The invention is particularly interesting when the make-up gas is produced by autothermal reforming. The dual hydrogen recovery and methane recovery of the invention is very suitable for integration with autothermal reforming, even more so when the hydrogen recovery is carried out in the synthesis loop refining using a membrane-based HRU, allowing the recovery without recompression of part of the methane-rich reserves as feed for autothermal reforming. Yet another aspect of the invention in practice is subjecting at least a portion of the make-up gas characterized by a low steam to dry gas ratio to a high temperature water gas shift conversion step using high temperature water gas shift catalysts suitable for operating at a low steam to dry gas ratio, for example from 0.1 to 0.5, where the steam to dry gas ratio is defined as S / DG=H2O / (1-H2O). Preferred examples Cooling of the make-up gas may be carried out in a cooling zone comprising a plurality of heat exchangers arranged in series to form a cooling train such that for each successive pair of first and second heat exchangers of the cooling train, the effluent of a first heat exchanger is further cooled in a second heat exchanger. In such a case, a small portion of the make-up gas is separated at an intermediate point from the cooling zone before passing through at least one of the heat exchangers. Accordingly, the separated portion of the make-up gas does not pass completely through the cooling train and is not completely cooled. The remainder of the make-up gas is further cooled in the cooling zone until the main make-up gas stream is completely cooled. For practical reasons, a preferred embodiment includes two heat exchanger sections in series, in which a small portion of the make-up gas is separated after passing through the first heat exchanger section and before entering the second heat exchanger section. The first heat exchanger section is configured to partially cool the make-up gas obtained from the reforming. The second heat exchanger section is configured to completely cool the remaining make-up gas after separating the small portion directed to the displacement and separating the hydrogen. The temperature of the gas sub-portion after separation is preferably 320 to 450° C. The temperature of the second gas portion, after complete cooling, is preferably 25 to 60° C. and more preferably 45° C. Preferably, the make-up gas cooling is carried out by indirect heat transfer with water or steam. The heat exchanger sections may include one or more steam generators, a steam superheater and / or a water preheater. In one example, hot water produced in the make-up gas cooling zone may be fed as a reagent to the reforming process. As mentioned above, the preferred application of the invention relates to the production of make-up gas by autothermal reforming, optionally after pre-reformation. The autothermal reforming is carried out in a suitable autothermal reformer (ATR) in the presence of oxygen or an oxygen-containing gas (air or enriched air), steam and optionally carbon dioxide over a suitable catalyst under effective oxidation conditions to produce make-up gas. The pre-reformation, if provided, may be carried out in one or more pre-reformers. In a preferred embodiment, the autothermal reformer operates at a low steam to carbon (S / C) ratio, preferably between 0.5 and 1.5, more preferably between 0.8 and 1.2. Preferably, the autothermal reformer operates at a pressure of between 25 and 60 bar, more preferably between 35 and 50 bar. The water-gas shift of the separated fraction from the make-up gas may include catalytic high temperature shift (HTS). The high temperature shift may be carried out at a temperature in the range of 300 to 500°C, preferably 350 to 450°C. Preferably, the water-gas shift of the separated fraction is carried out solely in a high temperature shift reactor without subsequent intermediate or low temperature shift. The first hydrogen recovery section preferably comprises a pressure swing adsorption (PSA) unit. The second hydrogen recovery section preferably comprises a membrane-based hydrogen purification unit. In an interesting example, a portion of the hydrogen recovered from PSA may be used as a fuel to provide process energy, and / or as a feedstock for ammonia co-production and / or for other processes outside of ammonia synthesis. Using hydrogen from PSA as a fuel may reduce the carbon emissions of the process. Specifically, the use of PSA-recycled hydrogen (“PSA hydrogen”) as a fuel offers the following advantages: The purified hydrogen obtained from PSA can power the plant heaters instead of the light hydrocarbons that are typically used as both feed and fuel. Direct plant emissions can be significantly reduced or eliminated entirely by anticipating the carbon capture step. The carbon to be captured remains concentrated in the high-pressure process streams rather than being diluted in the low-pressure waste gases, making the carbon capture step less costly. The use of PSA hydrogen for ammonia co-production is also possible thanks to the high purity of hydrogen that can be produced by PSA. The highly pure PSA hydrogen stream is suitable for direct use in an ammonia synthesis loop where the hydrogen purity specifications are very strict due to poisoning of the ammonia synthesis catalyst. The ability to recover hydrogen from two independent sources, namely a portion of the make-up gas and the loop purge, allows for high flexibility in the selection of the hydrocarbon source fed to the autothermal reactor. For example, if the amount of 2H recovered from the separated portion of the make-up gas is reduced due to the deposition of carbonaceous products such as coke or soot on the catalytic surface in the WGS reactor, this reduction can be compensated by increasing the amount of 2H recovered from the loop purge. Likewise, if the feedstock is to be converted to a hydrocarbon source characterized by a higher carbon to hydrogen ratio, the resulting reduction in stoichiometry can be compensated by increasing the amount of make-up gas fed to the associated HTS hydrogen recovery unit. The hydrocarbon source is preferably a light hydrocarbon source, preferably natural gas. In an embodiment of the present invention, the hydrocarbon source may be subjected to hydrodesulfurization and subsequently subjected to first and second pre-reforming steps. Preferably, the first pre-reforming is carried out in an adiabatic fixed bed reactor at a temperature of approximately 350 to 530°C; the second pre-reforming may be carried out in an adiabatic reactor at a higher temperature than the first pre-reforming reactor. Preferably, the temperature of the second pre-reforming step is approximately 500 to 750°C. Advantageously, treating the hydrocarbon source in a first pre-reformer and in a second pre-reformer operating at a higher temperature than the first one allows for high conversion of high molecular weight hydrocarbons (>2C) to methane in the pre-reformer and conversion of methane to synthesis gas in the second pre-reformer, thus preventing soot deposition on the catalyst surface in the autothermal reformer and increasing the life of the reforming catalyst. Furthermore, the oxygen requirement for the operation of the autothermal reformer is minimized due to the additional reforming duty absorbed by the secondary pre-reformer. In another example, the hydrocarbon source is fed to a desulfurization unit and subsequently to a single pre-reformer before being fed to the autothermal reformer. In another example, the hydrocarbon source is fed to the desulfurization unit and subsequently to the autothermal reformer, and no pre-reformation step is used. The autothermal reformer may be fired with oxygen or enriched air produced in the air separation unit. Another aspect of the invention is the integration of such an air separation unit with other equipment. The high pressure steam produced from the cooling of the make-up gas may be used to operate the air separation unit, for example to feed a single steam turbine. An intermediate pressure steam extracted from the air separation unit may be further used by injection upstream of the pre-reformer, thus acting as the process steam required in the reforming section. Advantageously, all of the steam produced in this process is utilized in the plant, thus providing improved design flexibility along with improved efficiency due to reduced natural gas consumption resulting from the use of an autothermal reformer operating at a low steam to carbon S / C ratio. Additionally, steam demand and production are separated between synthesis gas production and methanol synthesis to improve plant performance and flexibility. A specific preferred example includes: A light hydrocarbon feed is converted to synthesis gas in the following steps: hydrodesulfurization, pre-reforming, secondary pre-reforming (at higher temperatures), autothermal reforming; Autothermal reforming is carried out at a low S / C ratio of 0.5 to 1.5 and at a pressure of 25 to 60 bar. The inlet stream to the autothermal reformer consists of pre-reformed feed mixed with a portion of HRU; The resulting synthesis gas is cooled by raising the high-pressure steam; A portion of the newly produced make-up gas is removed and sent to the water-gas shift reactor. The water-gas shift effluent is subjected to 2H recovery by PSA; The resulting stream of pure hydrogen obtained from the PSA is mixed with the larger remaining portion of fresh make-up gas and with the hydrogen-enriched stream from a membrane HRU unit to form a regulated make-up gas that is fed to the synthesis loop compressor; The compressed regulated make-up gas is mixed with the circulating recycle stream and fed to a methanol reactor; The reactor effluent is cooled and the crude methanol is condensed and sent to the distillation section; A purge stream is removed from the unreacted gas mixture; The remainder (and the majority) of the unreacted gas mixture (recycle stream) is recycled through a recycle device to the methanol reactor inlet, mixing with the regulated make-up gas; The refining stream is sent to a membrane-based HRU, the hydrogen-reserved effluent is partially discarded and used as fuel, and partially recycled to the autothermal reformer; The hydrogen-enriched permeate is mixed with fresh, pure 2H make-up gas from the PSA to form the regulated make-up gas. According to another aspect of the invention, the make-up gas has a low steam to dry gas (S / DG) ratio. For example, this ratio is in the range of 0.1 to 0.5. In the case of a make-up gas with such a low S / DG ratio, the displacement reaction is preferably carried out on an iron-free catalyst. Preferred embodiments of the invention offer, among others, the following advantages: low consumption due to the low S / C used in the front section; the oxygen requirement for the operation of the autothermal reformer can be minimized thanks to the additional reforming duty absorbed by the secondary pre-reformer; steam production can be limited due to the low S / C used in the front section so that the heater does not have to be discharged. The steam network may be configured to take advantage of the low steam generation. In particular, steam demand and production between the syngas production and methanol synthesis sections of the plant can be separated, thus increasing the plant's performance and modularity. Preferably, the crude methanol extracted from the methanol synthesis loop is purified to high purity methanol by a distillation scheme comprising four columns. Preferably, the four columns include an overhead column to remove volatile components present in the crude and three purification columns to separate the methanol from water and higher alcohol by-products. Desirably, the four-column distillation scheme produces purified methanol using a reduced amount of steam for reboiling. Reduced steam is necessary to capitalize on the reduced gas consumption and steam generation provided by the low S / C front end. Preferably, the three column pressure levels are in the range of 12 to 16 bar, 6 to 10 bar and 0.5 to 3 bar, respectively. Desirably, the associated heat required for reboiling can be provided at a temperature level compatible with the low pressure steam produced by the operation of the steam turbines. Even more advantageously, the amount of reboil duty for the recovered distillation from the make-up gas cooling may be minimized and limited to a high reboil duty. The purification duty may be obtained from the condensation of steam extracted from the operation of a steam turbine. As a result, the steam demand and production between the synthesis gas production and methanol synthesis sections of the plant may be separated, increasing the plant's performance and modularity. In one example, the high pressure steam produced in the syngas cooling section / zone is used to drive an air separation unit that feeds pure oxygen to the autothermal reformer. The intermediate pressure steam may be extracted from the air separation unit steam turbine and used partly as process steam to drive the reforming reactions in the syngas production section and partly to drive further steam turbines for plant services. In an embodiment of the invention, the steam generated by the heat exchange in the methanol synthesis reactor is superheated in a fuel-fired reheater and used to drive the regulated gas compressor and the circulating recycle stream. Preferably, the low pressure steam is extracted from the steam turbine of the make-up gas compressor and condensed in the reheater of the first purification column. Description of shapes Figure 1 shows a methanol synthesis process according to an example of the invention. Detailed description of preferred examples Figure 1 shows a schematic 150 of a plant for producing methanol 56 from a light hydrocarbon 101, for example natural gas. Hydrocarbon 101 is supplied to hydrodesulfurization unit 41 via line 40, and sulfur-free hydrocarbon 42 is fed to a first prereformer 43, where high molecular weight hydrocarbons (≥2C) are partially converted to methane, hydrogen, and carbon oxides. The gas mixture 44 leaving the pre-reformer is then fed to a second pre-reformer 45 where further conversion of the hydrocarbons takes place, resulting in a hydrocarbon-containing gas 1 consisting mainly of methane, hydrogen and carbon oxides. Hydrocarbon containing gas 1 is fed with an oxygen containing gas 25 to an autothermal reformer 2 to produce a make-up gas 3 having a substoichiometric hydrogen content, which is a hydrogen deficiency exceeding a stoichiometric number determined by the methanol reaction. The oxygen-containing gas 25 is obtained from an air separation unit 47. The air separation unit 47 operates with the high pressure stream 52 obtained in the make-up gas cooling zone 60. An intermediate pressure steam 48 discharged by the air separation unit 47 is introduced into the first pre-reformer 43 and is used as process steam downstream (not shown). The hot make-up gas 3 exiting the reformer 2 is fed to a cooling zone 60 which, in the example shown, comprises a first heat exchanger section 4 and a second heat exchanger section 8. In the first heat exchanger of section 4, the hot make-up gas 3 is cooled to produce semi-cooled make-up gas 5 and a high-pressure stream 52 is produced. A small portion 26 of the semi-cooled make-up gas 5 is separated from an intermediate point 6 of the cooling zone 60, i.e. after passing through the first heat exchanger section 4 and before entering the second heat exchanger section 8. Said separated portion 26 is sent to a train comprising the HTS displacement reactor 27, the cooling section 57 and the PSA unit 33 for the separation of a hydrogen stream 34. The separated stream 26 is preferably about 2% of the make-up gas 5. The remaining portion 7 of the make-up gas (after separation from the above fraction 26) is sent to the second section of the heat exchanger 8 where it is further cooled and a stream 9 of fully cooled make-up gas is obtained. In more detail, the separated make-up gas 26 is fed to a high temperature shift reactor 27 to produce a hydrogen-enriched shift gas 28. Said shift gas 28 is then cooled in a cooling section 57 comprising a first heat exchanger 29 and a second heat exchanger 31. The effluent 30 of the first heat exchanger 29 is cooled in the second heat exchanger 31. The resulting cooled gas 32 is fed to a pressure swing absorption unit 33 to produce a hydrogen stream 34 and a waste gas 35 comprising methane and carbon dioxide. Said waste gas 35 may be sent to combustion. The temperature of the make-up gas 3 exiting the autothermal reformer 2 may be about 1000°C. The temperature of the separated make-up gas 26 at the inlet of the high temperature water gas shift reactor 27 is typically about 350°C. The temperature of the shifted gas 28 exiting the water gas shift reactor 27 may be about 470°C, and the temperature of the cooled gas 32 entering the pressure swing absorption unit 33 after appropriate cooling is about 45°C. The fully cooled make-up gas 9 is mixed at a mixing point 10 with at least a portion of the hydrogen stream 34 and with the hydrogen-rich permeate 20 recovered from the loop refining as described below. By mixing with the hydrogen stream 34 and the permeate 20, the hydrogen content in the make-up gas 9 is adjusted, i.e. the initial hydrogen deficiency is compensated. Typically, a hydrogen stream obtainable from a PSA unit is of high purity. It should be noted that the hydrogen stream 34 may contain unavoidable impurities. The regulated make-up gas 11 is fed to the methanol synthesis loop 14 after being compressed in the synthesis gas compressor 12. In some examples, the hydrogen stream 34 leaving the pressure swing absorption unit 33 may be fed directly to the mainline mixing point 10 , i.e., no compression step is required because the hydrogen stream 34 is extracted at sufficient pressure. A condensed crude methanol stream 15 and a refining stream 16 are extracted from the methanol synthesis loop 14. The condensed crude methanol stream 15 is purified in a distillation section 49 to obtain a pure methanol stream 56. Preferably, the distillation section 49 comprises four distillation columns operating in cascade. Particularly preferably, the four-column arrangement described in EP 2 617 478 may be adopted, which has the advantage of low steam consumption. The refining stream 16 is fed to a membrane-based hydrogen purification system 19 after suitable cooling in a heat exchanger 17, and a hydrogen-rich permeate 20 and a methane-rich bulk 21 are obtained. Typically, the 2H recovery in the hydrogen refining system 19 is carried out in such a way that the recovered 2H stream 20 is at a pressure compatible with direct mixing with stream 9. The hydrogen-rich permeate 20 is returned to the mixing point 10 where it is combined to adjust the hydrogen content in the make-up gas 9. At least a portion of the chamber 21 may, after cooling in a heat exchanger 24, be fed to the reformer 2 and used as additional feedstock for the synthesis of make-up gas 3. Recycling of the methane-rich retained material 21 eliminates the problem of methane slippage and prevents the heating duty from being saturated by plant waste gases. Portion 22 may be separated from the chamber and sent to combustion.

Claims

Claims 1) A process for producing methanol comprising the following steps: a. reforming a hydrocarbon-containing source into a make-up gas (3) comprising hydrogen, carbon oxides and water; .b subjecting the make-up gas (3) produced in step (a) to a cooling process; .c separating a portion (26) of the make-up gas during the cooling process of step b) and before the cooling process is completed, and subjecting the remaining portion (7) of the make-up gas to complete cooling, obtaining a main make-up gas stream (9), the separated gas (26) at a higher temperature than the main stream (9); .d subjecting said separated make-up gas (26) to at least one water gas shift (WGS) conversion step (27) to obtain a hydrogen-enriched shift gas (28); e. cooling said shift gas (28) and feeding the cooled shift gas (32) to a first hydrogen recovery section (33) and obtaining a first hydrogen stream (34); f. adding the main stream of make-up gas (9) with said first stream of hydrogen (34) and with the second stream of hydrogen (20) obtained in step (i), thus obtaining a regulated make-up gas (11) with a regulated hydrogen content; g.Feeding the regulated make-up gas (11) to a methanol synthesis loop (14) in which the catalytic conversion of carbon oxides to methanol is carried out under methanol synthesis conditions and a condensed crude methanol stream (15) is obtained; .hPurifying the condensed crude methanol stream (15), preferably in a distillation section (49), to obtain a methanol product (56); i. Feeding a refining stream (16) exiting the methanol synthesis loop to a second hydrogen recovery section (19), to obtain a second hydrogen stream (20) containing hydrogen removed from the refining stream, and a waste gas (21) containing methane; .jAdding the second hydrogen stream (20) to the main make-up gas stream (9) according to step (f); k. Using at least a portion of the waste gas (21) as a feedstock for producing the make-up gas (3) of step (a). 2) A process according to claim 1 wherein said separated portion (26) of the make-up gas is a small portion. 3) A process according to claim 2 wherein the volumetric flow rate of the separated portion of said make-up gas (26) is not more than 15% of the total volumetric flow rate of said make-up gas (3), preferably 1% to 10%. 4) A process according to any one of claims 1 to 3, wherein the cooling process of step c) is carried out in a cooling section (60) comprising a plurality of heat exchangers arranged in series, and the separate portion (26) of said make-up gas is separated after passing through at least one of the heat exchangers. 5) A process according to any preceding claim, wherein the modification of step a) comprises an autothermal modification optionally preceding the modification. 6) A process according to claim 5, wherein the autothermal reforming is carried out with a steam to carbon (S / C) ratio comprised between 0.5 and 1.5, preferably between 0.8 and 1.

2. 7) A process according to claim 5 or 6, wherein the autothermal reforming is carried out at a pressure between 25 and 60 bar abs, preferably between 35 and 50 bar abs. 8) A process according to any one of the preceding claims, wherein the gas water displacement conversion of step d) comprises displacement at elevated temperature, preferably between 300 and 500°C, more preferably between 350 and 450°C. 9) A process according to any one of the preceding claims, wherein the first hydrogen recovery section (27) comprises a pressure swing absorption unit. 10) A process according to any one of the preceding claims, wherein the second hydrogen recovery section (19) comprises a membrane-based hydrogen recovery unit. 11) A process according to any one of the preceding claims, wherein a portion of said first hydrogen stream (34) is used as a fuel to provide the energy required for the process and / or as a feedstock for the co-production of ammonia. 12) A process according to any one of the preceding claims, wherein said hydrocarbon-containing gas is obtained from a natural gas source (101) by hydrodesulfurization (41), pre-reforming (43) and secondary pre-reforming (45), wherein said secondary pre-reforming (45) is carried out at a higher temperature than said pre-reforming (43). 13) A process according to any one of the preceding claims, wherein the reforming of step a) is carried out with oxygen or an oxygen-containing stream produced in the air separation unit (47) and steam (52) produced in the make-up gas cooler of step b) is used to operate said air separation unit. 14) A process according to any one of the preceding claims, wherein the make-up gas (3) obtained in the reforming process has a steam to dry gas ratio of not more than 0.5, preferably 0.1 to 0.

5. 15) A process according to any preceding claim, wherein the purification of the condensed crude methanol stream (15) from step (h) is carried out in a distillation section (49) comprising four columns operating in cascade, wherein one of the four columns is a top column for removing volatile components and the other three columns are purification columns designed to separate methanol from water and higher alcohol by-products. 16) A methanol production plant from synthesis gas containing hydrogen, carbon oxides and optional inert components, comprising: a) a reforming section suitable for reforming a hydrocarbon-containing feedstock to a make-up gas (3) comprising hydrogen, carbon oxides and water; b) a cooling section arranged to cool the make-up gas (3) produced in step (a); c) a line arranged to separate a portion (26) of the make-up gas from an intermediate location of said cooling section and prior to complete cooling and a line arranged to subject the remaining portion (7) of the make-up gas to complete cooling in the section and to obtain a main stream of fully cooled make-up gas (9) at a temperature lower than the separated make-up gas (26); d) a water gas displacement section (27) connected to said line carrying the separate portion of make-up gas (26) and configured to produce a hydrogen-enriched displaced gas (28); e) a cooling section of the displaced gas and a first hydrogen recovery section (33) arranged to receive said displaced gas after cooling and producing a first hydrogen stream (34); f) a line for adding said first hydrogen stream (34) to the main make-up gas stream (9)and a line for adding the second hydrogen stream (20) obtained in step (i) to said make-up gas, thus obtaining a regulated make-up gas (11) with a regulated hydrogen content; g) a methanol synthesis loop (14) and a line for feeding the regulated gas (11) to said loop, in which the catalytic conversion of carbon oxides to methanol is carried out under methanol synthesis conditions and condensed crude methanol (15) is obtained; h) a purification section for the condensed crude methanol (15), preferably a multi-column distillation section (49), methanol (56) is obtained; i) a second hydrogen recovery section (19) arranged to receive a refining stream (16) exiting the methanol synthesis loop and to obtain said second hydrogen stream (20) and a waste gas (21) containing methane removed from the refining stream; j) a line arranged to feed at least a portion of the waste gas (21) as feedstock to the reformer section for producing make-up gas (3). 17) A plant according to claim 16, comprising one or more of the following: the reforming section comprises an autothermal reformer, optionally with one or more pre-reformer(s); the first hydrogen recovery section is a PSA unit; the second hydrogen recovery section is a membrane separation unit; said water-gas shift section from the separated make-up gas comprises a high temperature shift reactor. 18) A plant according to claim 16 or 17, wherein the condensed crude methanol purification section (15) of step (h) comprises four columns operating in cascade, wherein one of the four columns is an overhead column for removing volatiles, wherein one of the four columns is an overhead column for removing volatiles and the other three columns are purification columns designed to separate methanol from water and higher alcohol by-products.

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