NO304118B1 - Process for the catalytic production of liquid hydrocarbons from natural gas - Google Patents

Description

The present invention relates to a method for the catalytic production of liquid hydrocarbons from natural gas, comprising the steps of

(a) natural gas is separated into at least two fractions, a first fraction of gas enriched with methane and a second fraction enriched with C2+ alkanes (ethane, propane and higher

alkanes),

(b) at least part of the methane is selectively oxidized by means of molecular oxygen in the presence of a catalyst

for oxidative coupling of methane,

(c) at least part of the fraction enriched in C2+ alkanes is mixed with the effluent from the selective oxidation step after at least 80% of the molecular oxygen introduced in step (b) has been consumed in

step (b), and

(d) the mixture resulting from step (c) is pyrolyzed;

and the peculiarity of the method according to the invention is that (e) when the temperature of the mixture from step (d) has been brought to 300°C to 750°C, at least part of the olefins and optionally part of the C2+- the alkanes to aromatics in the presence of an aromatization catalyst.

These and other features of the invention appear in the patent claims.

The invention thus relates to a method for producing liquid hydrocarbons from natural gas.

The basic principles of the natural gas oxypyrolysis process, which combines selective oxidation of methane and pyrolysis of the C2+ alkanes that are formed and added to the oxidation effluent, are described in three French patents

(2,629,451, 2,636,627 and 2,641,531) and an article (Appl. Catal., vol. 58 (1990) 269).

In the first step (a) of the process, natural gas is separated into at least two fractions, where the first is methane with a reduced amount of C2+ alkanes (ethane, propane and higher alkanes), and the second fraction is composed of C2+ higher alkanes with a reduced amount of methane. At least part of the first fraction is then mixed with oxygen (pure or enriched) in an amount of up to 0.4 mol molecular oxygen per moles of carbon. Mix none can advantageously contain large amounts of water vapour. The molar amount of water vapor in relation to methane can be from 0 to 4, preferably from 0 to 2, even more preferably from 0 to 1, and more particularly from 0.05 to 0.50. The water vapor makes oxygen/hydrocarbon mixtures easier to handle and allows a simultaneous increase in selectivity for coupled products and conversion of methane in the step for the selective oxidation of methane.

The second step, which is step (b) of the process, comprises consumption of the molecular oxygen in a selective catalytic oxidation reaction at temperatures which are at least above 650°C, preferably above 750°C and even more preferably above 800°C. The pressure is generally from 1 to 15 bar, preferably from 1 to 10 bar and more particularly from 1 to 4 bar. The reaction, which is generally described as oxidative coupling of methane, is carried out in the presence of a catalyst which is stable at high temperature. Catalysts that are stable at high temperature are generally those that contain at least one heat-resistant oxide, such as magnesium oxide, calcium oxide, strontium oxide and the other oxides listed in Table 3 of the article published in Appl. Catal., vol. 67 (1990) 47. Catalysts that are of particular interest for oxidative coupling of methane are, among others, those described in French Patent No. 2,629,451 and the articles published in Appl. Catal., vol. 67 (1990) 47, and in Chem. Soc. Rev., vol. 18 (1989) 251.

The selective oxidation step can be carried out in a reactor with a fixed bed, a reactor with a moving bed or a reactor with an entrained bed. The use of a fixed-bed reactor is particularly advantageous in cases where the catalyst does not have good mechanical strength properties and for a relatively low methane conversion per review, e.g. less than 20%. Reactors with a moving bed, such as those with a boiling bed or an entrained bed, are very advantageous if the conversion per review e.g. is over 20%. Circulation of the catalyst provides better temperature control during heat exchange between the batch, the catalyst and the effluents.

Regardless of which reactor is used in the step for the selective oxidation of methane, this step is highly exothermic. It is therefore very important to lower the temperature in the drain in order to limit the formation of acetylene and coke which can form during long contact times at high temperature. For this reason, it is often advantageous to inject at least part of the second fraction into the hot effluent, the second fraction, comprising higher C2+ alkanes with a reduced amount of methane, flowing out from the first stage. +-hydrocarbons, which are generally paraffinic, serve to lower the temperature by thermomechanical quenching. In other words, they absorb the heat released in the oxidation step and convert it into olefins and hydrogen.

The third step, step (c), comprises adding at least part of the C2+ paraffins obtained in step (a) until the effluent from the selective oxidation of methane reaches at least 80% and preferably at least 95% of the molecular oxygen which was introduced in step (b) has already been consumed in the second step. This mode of operation has three advantages: 1) In step (b), with the selective oxidation of methane, it is not necessary for the operation to take place in the presence of -hydrocarbons which are more easily oxidizable than methane; 2) the addition of C2+ hydrocarbons to the effluent from the selective oxidation of methane after at least 80% and preferably at least 95% of the molecular oxygen introduced in step (b) has already been consumed, enables most of the molecular oxygen to be used for to activate methane instead of dehydrogenating C2+ alkanes; 3) the addition of C2+ hydrocarbons to the effluent from the selective oxidation of methane after at least 80% and preferably at least 95% of the molecular oxygen introduced in step (b) has already been consumed, enables the conversion of C2+ alkanes by heat absorption to olefins and hydrogen, as the latter can be used to recover carbon, e.g. by hydrogenation of CO to methane.

The fourth step, step (d), comprises maintaining the residence time for the combined effluents for a time which is sufficiently long to enable pyrolysis to achieve an ethylene/ethanol molar ratio of at least 1.2 and preferably above 1.4. - 1. In this step, the residence time is generally from 50 milliseconds to 10 seconds, preferably from 100 milliseconds to 2 seconds.

In a method for the production of liquid hydrocarbons from natural gas, which forms no part of this invention, the gaseous effluent leaving the fourth stage, stage (d), is brought to a temperature below 100°C, compressed and fed into the separation system . The compressed gas is then freed from water and C02 before the higher hydrocarbons (C2, C2=, C3, C3= and higher hydrocarbons) are separated from the light gases (CO, H2, CH4). The higher hydrocarbons can then: 1) be separated to produce olefins (ethylene, propylene and higher olefins) and recyclable alkanes; 2) oligomerize and/or aromatize to obtain liquid hydrocarbons and recyclable alkanes, or 3) dimerize to produce liquid C^+ olefins and recyclable alkanes.

In the method according to the invention, on the other hand, the effluent from the oxypyrolysis, in gaseous form and containing gases comprising hydrocarbons, such as ethylene, ethane and propylene, is treated with the advantage of an aromatization catalyst. It is of great interest to upgrade these C2~-C3 hydrocarbons to liquids, such as important petrochemical intermediates.

The effluent from step (d), which is then used for the rate for step (e), includes: 1) at least 40% by weight and no more than 95% by weight of the methane plus the water component, 2) at least 5% by weight, preferably at least 10% by weight % and even more preferably at least 15% by weight, of non-paraffinic hydrocarbon compounds; 3) less than 1% by weight, preferably less than 0.5% by weight and even more preferably less than 0.1% by weight, molecular

oxygen, and

4) other compounds [in amounts varying depending on the initial rate in step (a) and operating conditions in steps (a) to (d)], such as N 2 , CO, CO 2 , H 2 and C 2+ alkanes.

The fifth step, step (e), thus comprises bringing the effluent from step (d) to a temperature of from 300°C to 750°C and more particularly from 450°C to 600°C, after which the effluent is brought into contact with an aromatization catalyst which enables at least part of the olefins to be converted into liquid hydrocarbons. Conversion of the olefins and possibly at least part of the C2+ alkanes to liquid hydrocarbons can significantly affect the dimensions of the separation system. Furthermore, the placement of a hot unit, such as the aromatization unit, together with other hot units before the cold separation steps, has positive effects on the investments required for the operating unit for upgrading natural gas.

The invention has two important advantages. The at least partial conversion of the olefins and possibly the C2+ alkanes into aromatics before they enter the separation system enables:

1) a reduction in the dimensions of the separation system,

2) a reduction of necessary devices to bring together a hot (aromatization) unit with other hot (oxidation and pyrolysis) units.

The following examples illustrate different designs that can be used to obtain liquid hydrocarbons from natural gas according to the invention. The examples are illustrated using figures 1 and 2.

A preferred embodiment of the invention includes the use of a reactor with a boiling bed for the steps for selective oxidation and pyrolysis (oxypyrolysis) and a reactor with a moving bed for aromatization. In this embodiment, the gas, which is enriched with methane 3 originating from the separation of natural gas 1 in the separator 2, is mixed with molecular oxygen 5. In cases where it is advantageous to add water vapor to the batch, this water vapor is added either to the methane or to the oxygen or to the methane and oxygen before the methane mixes with the oxygen. It is very important to regulate the temperature of the gaseous mixture before it enters the reactor. For this purpose, the gases containing methane 3 and oxygen 5 can be heated independently of each other or the gases 6 can be heated together. After the gases have been preheated to a temperature which is generally below 750°C and preferably below 600°C, they are brought into contact with the catalyst for oxidative coupling of methane in the lower part 7a of the reactor 7 in fig. 1 or in the upstream part 7a of the reactor 7 in fig. 2.

The catalyst for this type of application must have good mechanical strength. The particles generally have a diameter from 20 µm to 4 mm. The size of the catalyst particles varies according to the operating conditions of the unit and the density of the catalyst. Although it is not possible to name all catalysts for the oxidative coupling of methane that are of potential interest for use in a fluidized bed reactor, mention may be made of catalysts described in French Patent No. 2,629,451, in Appl. Catal., vol. 76 (1990) 47, in Chem. Soc. Rev., vol. 18 (1989) 251 and e.g. catalysts such as BaC03/Al2C<3 (generally, but not necessarily in the presence of a few ppm of a chlorine source in the batch); Pb/Al2C>3 (often in the presence of an alkali metal and/or an anion containing sulfur or phosphorus); mixed oxides containing Na, B, Mg and Mn (such as NaB2Mg4Mn2O<x>;La/MgO (often in the presence of one or more alkaline earth metal or alkali metal oxides and/or other lanthanide oxides and optionally boron); combinations of Na or K, one or more alkaline earth metal oxides and optionally boron; MCeO3 perovskites (where M = Sr or Ba) and mixed oxides containing thorium or yttrium.

After completion of selective oxidation in the reactor 7 and in a position where at least 80% of the molecular oxygen introduced in step (b) has been consumed, at least part of the gas fraction enriched with C2+ alkanes 4 is added to the reactor. This fraction can be added in the expanded catalyst layer (Fig. 1) or at a level above the catalyst (such as in the release zone).

After a sufficiently long residence time in the reactor [step (d) in the process] to achieve the desired olefin content, the gases are led through pipeline 8 to a heat exchanger 9 where the temperature of the gases is lowered to 300°C - 750°C, and more specifically to 450°C - 600°C. The drain is then led through pipeline 10 in fig. 1 to the aromatization reactor lic. If the operating pressure in the aromatization reactor lic is above the pressure in pipeline 10, the effluent can advantageously be compressed at this point.

In one of the preferred embodiments, the aromatization reactor receives lic in fig. 1 batch to be aromatized together with fresh catalyst through the upper part. In this way, the catalyst that is deactivated by coke formation at the bottom of the reactor is removed. The deactivated catalyst is then regenerated in lid before being fed back through pipeline 13 and into the upper part of reactor lic. The batch entering the reactor is brought into contact with the aromatization catalyst. Although all known aromatization catalysts are suitable for the invention, the catalysts described in the applicant's French patent applications 91/06.193, 91/06.194 and 91/06.195 are preferred. The aromatization catalyst is used for the reaction where oxypyrolysis gases are aromatized. This reaction is of particular interest as it enables hydrocarbon gases to be upgraded to higher total value liquids (mainly benzene, toluene and xylenes).

The batch, which is the oxypyrolysis effluent and which contains ethylene and/or ethane and/or propylene together with other compounds, is brought into contact with the aromatization catalyst at a temperature from 300°C to 750°C and more particularly from 450°C to 600°C, with one mass review per hour of the batch in relation to the weight of the catalyst (PPH) from 0.5 to 150 hours<1>, preferably from 1 to 80 hours 1. The operating pressure is advantageously from 1 to 18 bar, preferably from 1 to 12 bar. The effluent from the aromatization reactor lic, enriched with aromatics, is transferred through the pipeline 12 to a separation system.

Another preferred embodiment of the invention (Fig. 2) includes the use of a reactor 7 with a fixed layer for the steps for selective oxidation and pyrolysis, preferably, but not necessarily, in the same chamber (there are thus two separate zones 7a and 7b ). In a special method (see Fig. 2), aromatization can take place in at least one reactor designated as 11a and 11b, which can be taken "out of the loop" to regenerate the catalyst located in the reactor (a swing reactor) . In this embodiment of the invention, the catalysts used for selective oxidation of methane and aromatization of at least part of the olefins and possibly part of the C2+ alkanes do not necessarily have to be particularly resistant to wear. The catalyst can thus be suitably used in the form of extrudates, fragments or particles. All catalysts for use in a fluidized bed reactor can be used in a fixed bed reactor. The catalyst for this type of design does not necessarily have to have particularly high mechanical strength.

In cases where it is advantageous to add water vapor to the batch, this water vapor is generally added either to the methane or to the oxygen, or to the methane and oxygen before the methane is mixed with the oxygen. It is very important to regulate the temperature of the gas mixture before it enters the reactor. In the case of a reactor with a fixed bed, for this purpose the gases containing methane 3 and oxygen 5 are heated independently of each other before the methane is mixed with the oxygen. The methane and oxygen are mixed in the chamber in the reactor 7 before they come into contact with the catalyst. As mixing is carried out in the reactor, the pipeline 6 does not necessarily need to be used in this embodiment. It is nevertheless often advantageous that part of the preheating of the gases takes place in the reactor. The water present in the batch is capable of receiving heat released by the catalyst by radiation. In this way, part of the preheating can take place in the oxidation reactor when using this method. The gases, which are preheated to a temperature which is generally below 750°C and preferably below 600°C, are then brought into contact with the catalyst for oxidative coupling of methane in the reactor 7.

After the selective oxidation in the reactor 7 has taken place and at a point where at least 80% of the molecular oxygen introduced in step (b) has been consumed, at least part of the gas fraction enriched with C_ - alkanes is added 4 to the reactor. The place where this fraction is added is generally after the catalytic layer in order to make the operation simple.

After a residence time which is sufficiently long to achieve the desired olefin content (molar ratio ethylene/ethane of at least 1.2 : 1), the gas is transferred to the aromatization reactor through pipelines 8 and 10 and through unit 9. The reactor unit for aromatization, indicated as 11a and 11b , is operated discontinuously. The part lia of the reactor unit receives e.g. gas to be aromatized, at a temperature and for a time long enough to convert at least part of the olefins and possibly part of the C2+ alkanes into aromatics. For at least part of this conversion in part 11a in the reactor unit, part 11b is put on regeneration to remove the carbon deposited on the catalyst during the previous cycle. After the regeneration step, part 11b is put back into operation, and the process thus continues.

The aromatized effluent in the aromatization reactor 11a and 11b is then fed to the separation system through the pipeline 12.

The batch entering the reactor is brought into contact with an aromatization catalyst which, in this invention, is preferably of the gallium aluminum silicate type, and which is characterized by the following weight composition:

(a) 0.01 - 10%, preferably 0.03 - 7%, gallium,

(b) 0.1 - 99.49% of a matrix selected from the group consisting of alumina, silicon dioxide, magnesium oxide, a clay and any combination of at least two of the

the above compounds, and

(c) 0.5 - 99.89% of an MFI zeolite synthesized in an OH~ or F medium in the absence or presence of organic compounds, its approximate chemical formula being generally as follows: M+, . (Siro, , X,A1 Ga 0. Q_) (x+y) "

(x+y) [96-(x+y)J x y 192

where M represents an alkali metal cation and/or an ammonium cation and/or a proton, x is a number from 0.1 to 12 and y is a number from 0 to 13.7.

The MFI zeolite preferably has a fluorine content of from 0.01 to 2% by weight, preferably from 0.02 to 1.5% by weight, fluorine being preferably incorporated during the synthesis.

The MFI zeolite used in the invention can be produced using any of the known production methods. It can thus e.g. are synthesized in a conventional OH~ or F medium in the presence of or in the absence of organic compounds.

Gallium can be deposited on the zeolite obtained in this way at this point, after which it can be processed by any of the known techniques. The processing can also take place before the gallium deposition. In this case, gallium is deposited on the formed solid.

The MFI zeolite can in particular be mixed with a generally amorphous matrix, e.g. a moist alumina gel powder.

The mixture is then shaped, e.g. by extrusion through a die. The zeolite content of the resulting carrier is generally from approx. 0.5 to 99.89%, the amount of matrix in the catalyst is from 0.1 to 99.49% and advantageously from approx. 10 to 60%. Processing can be carried out with matrices other than aluminum oxide, such as magnesium oxide, silicon dioxide-alumina, natural clay types (kaolin, bentonite), and by methods such as extrusion, pelleting, tableting, oil dripping or spray drying.

The deposition of gallium on the MFI zeolite before shaping or on the support after shaping is carried out by means of any method known in the art for the deposition of metal in zeolite. It is possible to use techniques for competitive cation exchange, where the competing agent is preferably ammonium nitrate, or techniques for depositing gallium on the catalyst by impregnation or precipitation. The solutions for gallium replacement or impregnation can be prepared from gallium compounds such as gallium oxide, nitrate, sulphate, halides or hydroxide. These methods for ion exchange, impregnation or precipitation can also be used to deposit the metal directly on the zeolite powder before it is optionally mixed with a matrix. The amount of gallium deposited on the catalyst at the end of the ion exchange and/or impregnation or precipitation steps is from 0.01 to 10% by weight of the entire catalyst and preferably from 0.03 to 7% by weight.

The catalyst which is obtained by the above methods and which can optionally undergo calcination treatment in air at a temperature which is generally from 350°C to 690°C, is used for the reaction where the oxypyrolysis gases are aromatised. This reaction is of particular interest as it enables the upgrading of hydrocarbon gases to higher value liquids (mainly benzene, toluene and xylenes).

The following examples illustrate the procedure.

The catalyst used in the following examples contains 20% by weight of a matrix and 80% by weight of zeolite in the form of extrudates.

Example

Preparation of aromatizing catalyst A

An MFI zeolite in hydrogen form, supplied by Conteka with the designation CBV 1502, is used. It is characterized by a Si/Al ratio of 75, a sodium content of 0.016% by weight and a pore volume, measured by nitrogen adsorption at 77K, of 0.174

cm/g.

The MFI zeolite is mixed with a moist alumina gel powder. The mixture is then shaped by extrusion through a die.

After the resulting extruded solid has been dried and calcined, gallium is deposited on the zeolite by ion exchange. The exchange solution is prepared from 0.2 M gallium nitrate. The pH value of the gallium nitrate solution is adjusted to a value of 2.2 with ammonia.

The gallium content of the resulting catalyst A, obtained after 4 exchanges, is 3.2% by weight.

Conversion of an effluent from oxypyrolvse

The catalyst A is used to aromatize an effluent 10 from oxypyrolysis, the weight composition of which is given in table 1. The oxypyrolysis effluent shall be described in the following as "the batch" and the conversion products of this effluent on the aromatization catalyst as "products".

The aromatization catalyst is fed into a reactor bed, made of stainless steel, with a fixed bed and where the operating conditions are as follows:

temperature: 53 0°C

pressure : atmospheric

review per hour at a floating rate corresponding to 20

times the weight of catalyst.

The results, denoted mass percentages, of products found in pipeline 12 are shown in table 2.

Claims (8)

1. Process for the catalytic production of liquid hydrocarbons from natural gas, comprising the steps of (a) separating natural gas into at least two fractions, a first fraction of gas enriched with methane and a second fraction enriched with C2+ alkanes (ethane, propane and higher alkanes ), (b) at least part of the methane is selectively oxidized by means of molecular oxygen in the presence of a catalyst for oxidative coupling of methane, (c) at least part of the fraction enriched in C2+ alkanes is mixed with the effluent from the selective oxidation steps after at least 80% of the molecular oxygen introduced in step (b) has been consumed in step (b), and (d) the mixture resulting from step (c) is pyrolyzed, characterized in that (e) when the temperature for the mixture from step (d) has been brought to 300°C to 750°C, at least part of the olefins and optionally part of the C2+ alkanes are converted to aromatics in the presence of an aromatization catalyst. 2. Method according to claim 1, characterized in that the temperature in the mixture from step (d) under step (e) is brought to a temperature of from 450°C to 600°C. 3. Method according to claim 1 or 2, characterized in that at least part of the first fraction (the methane fraction) at the end of step (a) is mixed with oxygen, the oxygen content can be up to 0.4 mol molecular oxygen per moles of carbon. 4. Method according to claim 3, characterized in that a mixture is used which contains a molar amount of water vapor of up to 4 in relation to the methane. 5. Method according to any one of claims 1-4, characterized in that at least part of the second fraction (C2+) obtained in step (a) is injected into the zone for selective oxidation of methane (step b). 6. Method according to any one of claims 1-5, characterized in that the residence time for the combined effluents obtained in step (c) in step (d) is kept until a molar ratio of ethylene/ethane of at least 1.2:1 has been achieved by pyrolysis , the residence time being from 50 milliseconds to 10 seconds. 7. Method according to any one of claims 1-6, characterized in that the aromatization catalyst contains (<*>) 0.01 - 10%, preferably 0.03 - 7%, gallium, (<*>) 0.1 - 99.49% of a matrix selected from the group consisting of alumina, silicon dioxide, magnesium oxide, a clay and any combination of at least two of the above compounds, and (<*>) 0.5 - 99.89% of an MFI zeolite synthesized in an OH~ or F medium in the absence or presence of organic compounds, its approximate chemical formula being generally the following: where M represents an alkali metal cation and/or an ammonium cation and/or a proton, x is a number from 0.1 to 12 and y is a number from 0 to 13.7. 8. Method according to any one of claims 1-7, characterized in that the MFI zeolite used in step (e) has a fluorine content of 0.01 to 2% by weight, fluorine being incorporated during the synthesis of the zeolite.