Process for the Production of a Substitute Natural Gas

This invention relates to a process for the production of a substitute natural gas (SNG) and in particular to such a process using methanol feedstock.

Methanol as an SNG feedstock has the following in-built advantages:

1. It is clean, free of sulphur, and any gum forming or coke forming substances;

2. It is therefore readily vapourized and requires no desulphurization step before being passed to a nickel catalyst;

3. It is a partially oxygenated compound and may be regarded as CH.sub.2.H.sub.2 O. It thus carries an in-built water supply as far as any steam reforming process is concerned;

4. If one considers the ideal overall stoichiometry of the conversion of methanol to methane viz.

it is readily appreciated that the process is a net generator of high quality water, thus no process water supply is required. Under oxygen free conditions, no deaeration or scavenging step is needed;

5. Methanol is capable of drying gases and of carbon dioxide removal.

In addition to the direct advantages of methanol as an SNG feedstock the properties of methanol are such that it constitutes an economical and convenient energy transfer medium in that it may be readily transported by pipeline without many of the practical difficulties and high cost involved in gas pipelines or pipelines for other liquids such as heavy oils. Other materials such as coal and light and heavy oils have been proposed as SNG feedstocks but the use of such materials is in many instances complicated by the necessity to transport the feedstock to the area of gas demand or indeed to manufacture the gas at source, for example, at the minehead, and then to transport the gas by pipeline to the area of gas demand. Thus it may be advantageous to convert such materials or other inconvenient solid, liquid or gaseous energy sources, often by well proven technology, to methanol for transportation, either by pipeline or sea tanker, and thereafter to manufacture SNG directly from the methanol in the area of gas demand. Previous processes for the production of SNG using methanol feedstock have involved the following steps:

A. initial gasification of methanol over a nickel catalyst to an outlet temperature of 500.degree.-550.degree.C;

b. cooling of gases so produced which contain mainly methane, hydrogen and carbon dioxide with less than 1% carbon monoxide, in the presence of steam, to 250.degree.-300.degree.C. A typical dry gas analysis would be:

c. adiabatic methanation of said gases to give a gas containing about 2 per cent water;

d. further cooling to 250.degree.-300.degree.C and final methanation to give a gas of a 1000 BTU/FT3 after carbon dioxide removal;

e. carbon dioxide removal by hot carbonate washing or other means;

f. gas drying, e.g. by a glycol wash process.

Methanol gives a high temperature rise on the initial gasification step (step (a)) and this necessitates the use of a high steam ratio in order to avoid carbon deposition in the outlet gas and further to avoid unacceptably high outlet temperatures from the reactor giving rise to problems in reactor design. This route may be improved by incorporating a number of adiabatic reactors in series with methanol addition and steam raising between each stage. Low steam ratios may then be achieved without excessive outlet temperatures.

We have now found that the production of SNG from methanol may be carried out at improved thermal efficiency and with less complicated plant design by using an isothermal reactor as opposed to an adiabatic reactor.

Accordingly, the present invention provides a process for the production of SNG from methanol feedstock which process comprises passing methanol vapour optionally admixed with a minor proportion of recycle carbon dioxide through a bed of nickel catalyst in an isothermal reactor and removing carbon dioxide from the product gas, the methanol vapour being passed into the catalyst bed at a temperature of at least 250.degree.C preferably about 250.degree.C and the bed being maintained by the reaction at a temperature of from 250.degree.C to 350.degree.C preferably about 300.degree.C.

Experiment has shown that the reaction of methanol vapour to form SNG takes place readily over a nickel catalyst at 250.degree.C. Thus, provided the reaction temperature remains at all times about 250.degree.C, there is no tendancy for the reaction to `quench`. The overall reaction is in fact strongly exothermic providing heat in excess of that required to maintain the catalyst bed in the required temperature range of from 250.degree. to 350.degree.C. The surplus heat or as much of the surplus heat as is necessary is preferably utilized in the process of the present invention for vapourizing and preheating further methanol feedstock to the required temperature of at least 250.degree.C prior to entry into the reactor.

One form of isothermal reactor which may be conveniently employed in the process of the present invention comprises a plurality of tubes surrounded by a shell such that when the reactor is in use catalyst may be contained in the tubes and boiling water may be circulated through the shell at an appropriate rate thereby conducting away surplus heat generated in the reaction tubes and maintaining isothermal or near isothermal reaction conditions.

An alternative and preferred form of isothermal reactor for use in the process of the present invention comprises an inner tube sheathed by an outer tube concentric with said tube such that when the reactor is in use catalyst may be contained in the inner tube and boiling water may be circulated through the annular space between the inner and outer tubes to maintain isothermal or near isothermal conditions for the reaction occurring in the inner tube. This arrangement enables high heat transfer areas to be obtained without the need for large pressure vessels. However the tubes of the concentric tube reactor are limited in size by practical considerations, for example to an inner tube having an inside diameter of about 4 inches, thus in a preferred embodiment, to secure adequate output, a plurality of concentric tube reactors may be employed. In such an arrangement all the reactors, optionally contained in a mild steel casing and insulated from one another for example with purlite insulation, may be piped separately to inlet and outlet headers for both feedstock and cooling water.

At steam pressures above 550 psig. the boiling point of water is above 250.degree.C. Thus by raising steam in the above described isothermal reactors at 550 psig. or higher pressure, reaction quenching may be prevented and the temperature rise of the reactants may at the same time be limited such that the temperature of the catalyst bed and the outlet temperature of products is close to 250.degree.C and lies in the required range of from 250.degree. to 350.degree.C. At temperatures in this range methanol is completely converted to methane and carbon dioxide with less than 1% hydrogen in dry gas product and minor traces of carbon monoxide (.about.0.1%) only.

Circulation of boiling water through the reactors may be effected by means of a pump or by natural circulation. As stated above, some of the surplus heat conducted away by the boiling water is preferably utilized in the process of the present invention to vapourize and preheat methanol feedstock, according to known techniques. The remainder of the surplus heat may be used, for example, to drive turbines or other machinery.

As catalyst in the process of the present invention may be employed, for example, any of the nickel catalysts conventionally used in the gasification of methanol. A preferred catalyst in Catalytic Rich Gas (CRG) catalyst as described in British Patent Specification No. 820,257 and which consists of reduced nickel activated with alumina, and contains, for example, 15% of nickel. Such a catalyst may be prepared by treating an aqueous solution of water-soluble salts, for example, the nitrates, of nickel and aluminium with an alkali, such as sodium carbonate, to produce a precipitate of a mixture of nickel and aluminium compounds, washing and drying the precipitate, reducing the nickel compound to metallic nickel, and granulating or pelleting the resulting mixture of reduced nickel and alumina. Alternatively, granules of active alumina may be impregnated with an aqueous solution of nickel nitrate, then roasted, and the nickel oxide reduced to metallic nickel.

Removal of carbon dioxide from the gas produced in the reactor, usually after cooling of the gas, may be effected by any known technique. For instance the gas may be partially scrubbed of carbon dioxide using a `coarse` process such as water washing or `flash` Benfield. Cryogenic separation may be used to bring the level of CO.sub.2 down to 6 to 10% volume. The remainder of the carbon dioxide may be removed using methanol feedstock scrubbing which depends in principle on the high solubility of carbon dioxide in methanol at high pressures and low temperatures. Recycle carbon dioxide contained in methanol feedstock which has been obtained in this way has little effect on the equilibrium obtained at the outlet of the reactor.

After removal of carbon dioxide, preferably followed by enrichment in conventional manner, the gas obtained by the process of the present invention is intended to be interchangeable with natural gas.

It is to be understood that the invention also includes SNG when produced by the process of the invention and apparatus for producing SNG as specifically described herein.

The invention will now be further illustrated by the following examples of which Example 1 is a comparative example relating to the known adiabatic process and Example 2 relates to the isothermal process of the present invention.

A typical flow diagram for this process is shown in the accompanying FIG. 1.

The object of this process is to limit the temperature rise in each reactor to 250.degree.C by interstage cooling. Initially water and methanol are mixed in the ratio 4 lbs./lb vapourised and fed to the first adiabatic reactor at a preheat of 250.degree.C. Gases leave the reactor in equilibrium at 500.degree.C and are cooled to 250.degree.C. Cooling may be achieved wholly by a steam raising waste heat boiler or partly by a boiler and partly by direct injection of methanol feed. In this example liquid methanol at ambient temperature is used and this, of course, eliminates the need for a methanol vapouriser. The amount of methanol added is arranged so as to give a 250.degree.C temperature rise in the next adiabatic reactor. Temperature control is achieved by bypass control of the waste heat boiler. The process is continued in a similar manner for any number of reactors the number being limited by complexity of plant and reactor sizing. The outlet temperature of 500.degree.C from the final reactor is not compatible with the production of a satisfactory substitute natural gas and a further two adiabatic stages are required. In the first stage the gas is cooled to 250.degree.C but no further methanol is added. The resultant outlet temperature is in the range 300.degree.-350.degree.C depending on the number of preceding stages. This gas is then cooled to a temperature at which water is condensed and all the water that was originally added to the methanol is rejected. The gas is reheated to 250.degree.C and passed to a final reactor where equilibrium is reached at an outlet temperature in the range 260.degree.-300.degree.C. The final gas, after carbon dioxide removal to 1% and enrichment to 1000 Btu/scf is a fully interchangeable substitute natural gas. CRG catalyst would be used throughout the process. The major advantage of the process is the high thermal efficiency of using the multiple reactor arrangement in that the greater the stages employed the less water is required to be vapourised per unit mass of methanol gasified. If more than four stages are used the process is self supporting in heat requirement and surplus steam may be expected or used for power generation.

The composition of the gas produced at each stage of the process illustrated in FIG. 1 is given in the following Table 1.

In this process the temperature rise due to the exothermic reaction is controlled by directly cooling the catalyst vessel with a jacket of boiling water and steam. This technique enables methanol to be gasified directly without the necessity to add water to the feed. This example is also used to demonstrate the use of methanol to recycle carbon dioxide round the process. Methanol containing 20 mole percent of carbon dioxide is vapourised and preheated to 250.degree.C prior to introduction to the isothermal reactor. The reactor (the design of which is discussed hereinbelow) is cooled with boiling water at 850 lb/in..sup.2 and the flows are adjusted to give an outlet temperature of 300.degree.C. The gases are cooled and passed to the first stage of carbon dioxide removal. There is insufficient heat in the gases to vapourise and preheat the feed and this must therefore be accomplished by heat exchange with steam. In the coarse carbon dioxide removal unit, carbon dioxide formed from the methanol feed is removed. The recycle carbon dioxide is then removed in the `Rectisol` unit leaving a gas which is substantially free of carbon dioxide and which after enrichment to 1000 Btu./scf, is interchangeable with natural gas.

Two types of reactor design may be used:

a. the catalyst may be contained in tubes in a bundle contained in a pressure shell through which the cooling water and steam circulates. The major limitation with this design is the size of the outer shell. For example a 10 mmscfd plant would have a heat release of 64.10.sup.6 Btu./hr. in the reactor. For an assumed heat flux of 1000 Btu./ft..sup.2 a heat transfer area of 6400 ft..sup.2 is required. If 3 inch I.D., 3.5 inch O.D. tubes 20 ft. long are used, 407 tubes would be required and the methanol loading to the catalyst would be 2250 lb/ft..sup.2 hr. If the tubes are arranged on a 5 inch triangular pitch the shell would be 9 ft. I.D. and 9 ft. 6 in. O.D. This type of design results in comparatively low heat transfer area to cross sectional area ratio because of practical and economic restrictions on the outer pressure shell. This ratio cannot be improved by reducing the catalyst tube diameter beyond 3 inches because of difficulties with loading and unloading of catalyst. The length of the reactor could be extended but such a vessel 100 ft. long for example would be very expensive.

b. The catalyst may be contained in tubes each of which is surrounded by another tube with boiling water in the annular space between them. This eliminates the problems with a large pressure vessel and the loading to the catalyst, the length of the tubes and the number of tubes can all be increased. For example a 50 mmscfd plant with a loading of 7000 lb/ft..sup.2 hr to the tubes and an assumed heat flux of 1000 Btu./ft.sup.2 hr would have 346 4.5 inch O.D. 4 inch I.D. tubes 100 ft. long each surrounded by a cooling water tube. The catalyst tubes would probably contain 1/4 inch CRG catalyst. Such an arrangement could be contained in a mild steel casing with purlite insulation between the tubes and the casing size would be approximately 12 ft. by 20 ft. Each tube would be piped separately to inlet and outlet headers for both gas and cooling water and the water would be circulated by pumps to a steam drum situated above the tubes.

A typical flow diagram for the process of the invention incorporating a reactor as described at (b) above is shown in the accompanying FIG. 2. The composition of the gas produced at various stages of this process is presented in the following Table 2.

It is believed that in addition to methanol gasification to SNG, other exothermic reactions could be profitably carried out in the apparatus of the present invention employing an isothermal reactor. Examples of such reactions would be methanation of carbon oxides to convert a lean gas, such as LURGI gas, to SNG and reduction of hydrogen content of GRH effluent or FBH effluent. The invention could also be applied to the direct gasification of naphtha feedstocks provided the reactor feed is comprised, in part, of elemental hydrogen. The presence of hydrogen reduces the "quench" temperature of the gasification reaction so that, although the reactants may be cooled to the jacket temperature by passage over a layer of deactivated catalyst, the hydrogen enables the reaction to proceed. The source of hydrogen for this type of process would be typically the effluent from a CRG reactor or the effluent from a high temperature tubular reformer.