FR2546879A1 - Process for the manufacture of substitution gas (SNG) from methanol or a methanol/synthesis gas mixture - Google Patents

Abstract

The present invention relates to a catalytic process for the manufacture of a gas rich in methane from methanol alone or from a methanol/synthesis gas (CO + H2) mixture. The process uses a stationary bed of catalyst 12 through which the reactants and a liquid coolant fluid pass from the top downwards. The processing takes place at a temperature from 200 to 400 DEG , under a pressure of 0.5 to 20 MPa, at a volumetric rate of the reactants of 500 to 20,000 h<-1> and at a volumetric rate of the liquid coolant of 50 to 1,000 h<-1>. The separation of the liquid and gaseous compounds is carried out at the bottom of the reactor 11. The liquid coolant is recycled 13 after heat exchange 16 at the reactor head. The steam after condensation 19 is separated from the other gaseous products 21 which are purified 24 and fractionated. The methane is collected 27.

Description

Since the crisis of 1973-1974, the problem of the gasification of coal and its subsequent transformation into "substitute" natural gas has once again been on the agenda.

As this gasification requires significant investments, the installations, in order to be profitable, must be of large capacity and be located on the site of the coal extraction.

For these industrial complexes then the crucial problems of transport and storage arise, which, unfortunately, substitute natural gas, with its low volume density, makes it very difficult to solve.

This is how we arrived at the concept of two-stage upgrading of coal.

In a first step, on the site of the extraction, the coal is gasified and transformed into methanol which is easily stored and transformed.

In a second step, the methanol is directed to areas and markets of high consumption, or, alone or combined with a synthesis gas from others. sources such as, for example, coking plants or steam reformers of naphtha9 fuels or bitumens, it is transformed in medium or small-sized units into substitute natural gas made directly available to the user.

To make the operation feasible, it remains to develop, and this is the object of the present invention9, a simple and safe process, capable of transforming into substitution natural gas, indifferently, either methanol alone or a mixture. methanol and syngas.

STATE OF THE ART.

The conversion of methanol into a substitution gas is based on reactions known in the prior art.

The catalytic transformation of methanol to methane takes place, primarily, depending on the overall reaction.

fortement endothermique exothermique fortement exothermique
Les reactions précitées sont délicates à conduire a cause de la réaction parasite dite reaction de BOUDOUARD

which actually results from the following three reactions

strongly endothermic exothermic strongly exothermic
The aforementioned reactions are difficult to carry out because of the parasitic reaction known as the BOUDOUARD reaction.

This reaction risks causing a deposit of carbon in the reactor. It is known in the art that the presence of water vapor in the feed prevents the development of this reaction.

In US Pat. No. 4,239,499, it has been proposed to maintain the water vapor / methanol molar ratio between 0.5 and 1.5.

On the other hand, US Pat. No. 3,920,716 indicates that water can have a protective effect, even in the form of dissolved liquid.

The recommended water / methanol molar ratio is between 0.09 and 1.8.

From an industrial point of view, the problem lies in controlling the rise in temperature of the reactants, due to the high exothermicity of the reaction.

This rise, depending on the composition of the reactants, can be as high as about 100 ° C to about 30 ° C per percent load converted.

As the optimum zone of activity and especially of stability of the usual catalysts extends over approximately one hundred degrees centigrade, one can easily conceive of the problems to be solved in order to bring the transformation to its economic end.

Different solutions have been proposed.

Thus, it is known for the transformation of methanol into methane to cool the gases leaving the reactor and to recycle them in part (Pipeline and Gas Journal, February 1973, p. 58-62). However, since the recycling rate is quite high, it results in a high expense due to the problems of recirculation and compression.

In the process proposed by US Pat. No. 4,239,499, provision is also made to recycle part of the mixture of gaseous products at the inlet of the reactor as a constituent of the vaporized feed in order to regulate the reaction temperature; this process therefore presents the same disadvantages.

Another technique (European Patents No. 34,014 and US 3,920,716) consists in carrying out the reaction in the presence of a liquid hydrocarbon phase containing solid catalyst in suspension.

The advantage of this technique is to be able to easily remove the heat of reaction by recirculating the liquid phase and cooling by a suitable refrigeration system.

In the present case, its application proves to be delicate because of the low catalytic concentration imposed by the necessities of the recirculation and the risks of powder catalyst deposits in hot areas of the unit where it could subsequently catalyze the process. gas phase methanation, with the resulting exothermic phenomena.

While the techniques proposed for methanizing methanol are numerous and varied, those capable of fulfilling the same function in the presence of a mixture of methanol and synthesis gas are much rarer, if not non-existent.

DETAILED DESCRIPTION OF THE INVENTION.

The process of the present invention is intended for the production of methane to be used as substitute natural gas, from a feed consisting of anhydrous or aqueous methanol and / or from a mixture in all proportions of one of these two feeds. and syngas form hydrogen and carbon oxides.

It is characterized in that the above feed, in vapor phase, and a hydrocarbon phase are circulated in co-current directed from top to bottom, through a fixed bed of a methanation catalyst, at a temperature of 200 to 400 "C, under a pressure of 0.5 to 20 M Pa, at a volumetric speed of vaporized methanol of 500 to 20,000 h'l and at a volumetric speed of the liquid of 50 to 1,000 h
Figure 1 gives a detailed description of an industrial unit operating according to the method of the invention,
The unit is supplied with methanol via line 1.

This load is preheated by heat exchange with the products leaving through the exchanger 2.

From exchanger 2, the methanol, via line 3, is directed to recuperative vaporizer 4.

From vaporizer 4, via line 5, the charge of vaporized methanol is introduced into line 6 where it is mixed with
- hot water (around 200 degrees centigrade) slightly methanolic, coming from line 9.

- heat transfer fluid and unconverted aqueous methanol recovered from the products at the unit outlet and which are brought through line 8.

- where appropriate, synthesis gas introduced through line 7.

These different streams are brought to operating conditions by direct contact with a stream of heat transfer fluid which arrives through line 10.

The heat transfer fluid is a hydrocarbon cut, for example a C10-C16 cut.

Its flow rate is chosen between 50 and 1,000 h 1. If it is too low, the thermal control within the catalytic bed is insufficient and the catalyst is deactivated; if it is too strong, the reaction is completely inhibited.

For optimum operation, it is preferred to keep a weight ratio between the hydrocarbon cut and the methane obtained of between 10 and 500, preferably between 50 and 250 kilograms of cut per kilogram of methane produced.

The hot water supplied via line 9 is intended to supply water in the vapor phase to the inlet of the reactor.

It is advantageous to maintain at the end of the reaction a molar ratio, water / methane produced, between 0.5 and 30, preferably between 0.5 and 20.

From line 6, the products open onto reactor 11 and pass through catalyst bed 12.

Any of the catalysts known to those skilled in the art can be used.
These are most often mass catalysts or catalysts deposited on a support (alumina, silica, silica-alumina and / or molecular sieve-based supports; it is also possible to use carbon-based supports. ; these supports may themselves be modified by various metals including those of groups I, II and those of the rare earth numbers 57 to 71 inclusive).

The active elements include at least one metal A from the following list: nickel, cobalt, iron, ruthenium, rhodium, molybdenum, tunsgtene, manganese, possibly associated with at least one metal B from the list following: copper, silver, chromium, aluminum, zinc, rare earth metals of numbers 57 to 71 inclusive. The preferred formulas include at least one metal from the nickel, cobalt, iron, manganese, ruthenium group associated with at least one metal from the copper, aluminum, zinc, magnesium, zirconium, chromium, lanthanum, cerium, neodymium, praseodymium group.

A particularly preferred catalyst is formed from a mixture of two catalysts, in weight proportions of 1; 10 to 10: 1 The first catalyst contains, by weight, 0.5 - 5% ruthenium, 5 - 30% nickel (NiO count) and 60 t 90% of a support from the silica, alumina and silica group -alumina; the second catalyst contains, by weight, 20 - 60% copper (counted as CuO), 10 P 40% zinc (counted as ZnO) and 2 - 20% aluminum (counted as Al 2 O 3).

These two catalysts can also be used in the form of a mixed bed, the first catalyst being placed at the bottom of the bed, the second catalyst being placed at the top of the bed, or alternatively in the form of multilayer beds, each layer of the second catalyst. being followed, in the direction of flow of the reactants, by a layer of the first catalyst.

These catalysts will advantageously be used in the form of calibrated particles (balls, extrudates, solid pellets, hollow pellets) with an equivalent diameter of 1 to 10 mm and preferably 2 to 5 mm.

The recommended operating conditions are as follows: the reaction temperature is between approximately 200 and approximately 400 "C and preferably between approximately 250 and approximately 350" C; the total pressure is between approximately 0.5 and approximately 20 megapascals (MPa) and preferably between approximately 1.5 and approximately 10 MPa.

The hourly volumetric speed for gas is between 500 and 20,000 h-l and, preferably, between 1,000 and 10,000 h 1.

The hourly volumetric speed for the liquid, expressed in volume of liquid per volume of catalyst and per hour, is between 50 and 1,000h 1 and preferably between 80 and 500h-1.

<tb>

<SEP> H20
<tb><SEP> Finally <SEP> the <SEP> report <SEP> H20 <SEP> or <SEP> possibly <SEP> (if <SEP> y <SEP> has <SEP> also <SEP> of the
<tb><SEP> CH3OH <SEP> H2 <SEP> 0
<tb> gas <SEP> of <SEP> synthesis) <SEP> the <SEP> ratio <SEP> 5 <SEP>, <SEP> expressed <SEP> in <SEP> moles, <SEP> measured
<tb><SEP> CH30H + CO
<tb> at the inlet of the reactor, st between 0.05 and 5 and preferably between 0.1 and 2.

After transformation on the catalytic bed 12, the products arrive at the bottom of the reactor Li where they separate.

The liquid leaves through line 13 from where it is taken up by recirculation pump 14 to be introduced through line 15 into exchanger 16.

The exchanger 16 evacuates the heat of reaction to the outside by producing high pressure steam. From the exchanger 16, the coolant returns to the reactor via lines 10 and 6.

The vapor phase which contains the synthesized methane leaves through line 17. This vapor phase passes through exchanger 4, line 18, exchanger 19 and line 20 to end in separator tank 21.

Through exchanger 4, the vapor effluent heats up and vaporizes the methanol feed as has been said above.

In exchanger 19, most of the water and the solvent contained in the reaction effluent is condensed.

It is advantageous to carry out this condensation at high temperature, around 100 to 250 degrees, preferably 150 to 230 degrees centigrade.

This way of proceeding appreciably increases the quantity of high pressure steam produced in the exchanger 16.

In the separator flask 21, two liquid phases are collected.

The aqueous liquid phase which mainly contains water with a little methanol and dissolved solvent passes into line 22 from where it is taken up by pump 33 to be reintroduced into the reactor via lines 9 and 6.

The liquid hydrocarbon phase passes through line 23 in unit 24 for separating and purifying the products,
From the separator flask 21, the vapor phase passes through the intermediary of line 25 through the methanol preheater from feed 2.

It is then directed by means of the pipe 26 on the separation and purification unit 24,
From unit 24, the purified and conditioned methane leaves via line 27, carbon dioxide via line 28, the water of reaction via line 29, methanol and the solvent recovered via line 30.

Part of the hydrocarbon phase is purged via line 31 and replaced by a clean cut of hydrocarbons supplied via line 32.

EXAMPLE
In an installation of the type described above, a feed consisting of half of methanol and half of synthesis gas is converted.

Thus, we introduce, from above, into the reactor (11)
- 4100 kgh of vaporized methanol.

- 4,100 kg / h of synthesis gas including 512.5 kg / h of hydrogen and
3587.5 kg / h of carbon monoxide.

- 6300 kg / h of heat transfer fluid (cut of hydrocarbon C10'618)
recovered from products leaving the unit (line 18).

- 5080 kg / h of water approximately 200 "C and slightly methanol (line 9)
- 400,000 kg / h of recycled heat transfer fluid (line 13).

The recovery heat transfer fluid contains aqueous methanol.

All of this mixture is brought to operating conditions (temperature = 325 ° C., pressure = 6 MPa) by direct contact with the recycled coolant, withdrawn from the bottom of the reactor.

The reactor with a diameter of 0.9 m and a height of 7 m contains 2700 kg of a catalyst having the following overall composition (% by weight)
SiO2 = 50
CuO = 23.4
Al203 = 3.1
ZnO = 12.5
Ru = 1
NiO = 10
This catalyst is in the form of balls, with a diameter of between 2.4 and 4 millimeters.

After conversion on the catalytic bed, a vapor phase is recovered which contains 3000 kg / h of synthesized methane, 280 kg / h of unconverted methanol, 7240 kg / h of water and 6300 kg / h of heat transfer fluid, hydrogen and unconverted carbon monoxide as well as carbon dioxide, by-product.

Claims (6)

- CLAIMS - 1.- A method of manufacturing substitution gas with a high methane content, characterized in that a mixture of vaporized methanol and a liquid hydrocarbon phase is circulated, in co-current, from top to bottom, to through a fixed bed of a. methanation catalyst, at a temperature of 200 to 400 C, under a pressure of 0.5 to 20 MPa, at an hourly volumetric speed of vaporized methanol from 500 to 20,000 h 1 and at an hourly volumetric speed of the liquid phase of 50 at 1000h \ and in that the gas thus produced is collected. 2.- Method according to claim 1, wherein the reaction temperature is between 250 and 3500C, the pressure is between 1.5 and 10 MPa, the hourly volumetric speed of the vaporized methanol is between 1000 and 10,000 h-1 and the hourly volumetric speed of the liquid phase is between 80 and 500 h 1. 3. A method according to claim 1 or 2, wherein the methanol is used as a mixture with synthesis gas containing carbon monoxide and hydrogen. 4.- Method according to one of claims 1 to 3 according to which, in addition to methanol, or the methanol-synthesis gas mixture, the feed also contains 0.05 to 5 moles of water per mole of methanol or of mixture ( methanol + CO), the water vapor then being present at the end of the reaction at a rate of about 0.5 to about 20 moles of water per mole of methane produced. 5. A method according to one of claims i to 4 wherein the weight ratio between the liquid hydrocarbon phase and the methane obtained is between 50 and 250 kg of liquid phase per kg of methane produced. 6. A method according to one of claims 1 to 5, wherein the steam effluent at the outlet of the reactor is condensed under pressure at around 150 to 300 degrees centigrade.