DK142624B - Process for producing a methane-rich gas. - Google Patents

Description

(11) FOREIGN / ANNOUNCEMENT 142624 DENMARK «ο int.ci.» o 10 κ 3/02 (21) Application No. 1621/78 (22) Filed on 1J. April 1978 mi (24) Race day 1 3. April 1978 (44) The application presented in the writ of annulment published on 1. d.00. 1 980

DIRECTORATE OF

PATENT AND TRADEMARK LABEL <30) Woritet requested from the (71) HALDOR TOPS / E Å / s, Nymøllevej 55 »2800 Lyngby, DK.

(72) Inventor: Carsten Sejrbo Nielsen, Poppelgade 9, 1st TV., 2200 Purchases harbor N., DK.

(74) Plenipotentiary in the proceedings:

Office of Industrial Energetic v. Svend Schørmlng .__ (54) Procedures for the production of a methane-rich gas.

The present invention relates to a process for producing, in at least one adiabatic methanization reactor, a methane-rich gas by catalytic conversion of an input stream, obtained by combining a preheated synthesis gas consisting essentially of hydrogen and carbon monoxide in the volume ratio H2 / CO 3: 1 with a recycle stream from the reactor, by means of a methanization catalyst whose active component is mainly nickel, and a shift catalyst to shift convert a portion of the carbon monoxide with water vapor into the recycle stream, the combined stream being passed through the catalyst bed in the reactor and an output stream thereof at a temperature between 500 and 700 ° C are divided into the recirculating stream which is cooled and a product gas stream which is diverted for use or further treatment.

In recent years, there has often been a local and / or periodic shortage of natural gas. Methanisation processes have therefore attracted considerable attention in connection with the production of methane-rich gases suitable for natural gas substitute gas, also known as synthetic natural gas or SNG.

The replacement of natural gas will normally take four steps.

The first sub process is a gasification of carbonaceous material, usually solid or liquid fossil fuel such as coal or fuel oil. In this process, the coal or hydrocarbons are converted to carbon oxides as well as smaller amounts of hydrogen and methane. The pressure gasification is usually followed by a so-called shift conversion. In this process, the following so-called shift reaction takes place: (1) CO + H2 O ^ = in CO 2 + H2 ΔΗ298 ° Κ = ~ 9'9 kcal / mol

The purpose of the shift process is to convert part of the carbon monoxide to hydrogen. Thus, a complete conversion of the carbon monoxide is not sought, but usually a ratio of hydrogen to carbon monoxide in the exhaust gas from the shift reactor of 3 or just over 3.

Before the gas can be subjected to methanization, it must be purified as the methanization catalyst will otherwise be deactivated. The gas purification is usually a wash of the gas either with an alkaline aqueous solution or an organic liquid such as methanol. Many gas purification processes have been developed and used industrially. In the purification process, according to its nature, carbon dioxide and sulfur compounds are first and foremost removed, but also other trace elements which can act as catalyst poisons. You can do the gas cleaning before and / or after the shift conversion, depending on how it fits with the current operating conditions.

The methanization is carried out according to the reaction equations below: (2) CO + 3 H2 * = ^ CH4 + H2O δΗ298 ° κ = -49.3 kcal / mol (3) CO2 +4 H2i = ^ CH4 + 2H20 AH2ggoK = -39.4 kcal / mol 3 142624

The formation of methane, whether according to reaction (2) or (3), is thus accompanied by strong heat generation. Therefore, the temperature of the reactants and products will rise during passage through a catalyst bed in an adiabatic reactor.

Such a rising temperature, according to Le Chatelier's principle, will tend to shift the equilibrium towards lower methane concentration. Accordingly, complete or almost complete conversion will only be possible if the rise in temperature is limited by cooling the reacting gas in some way, for example, by recirculating cooled product gas.

Danish Patent Application No. 4956/75 discloses the preparation of a methane-rich gas in at least one adiabatic methanization reactor containing a bed of a methanization catalyst in which a stream of a pre-heated methane synthesis gas containing hydrogen and carbon monoxide is combined with recycled methane stream from recirculation stream. is passed through the catalyst bed in the reactor, while an output stream from the reactor is divided into said recycle stream and a stream of product gas for further processing or collection for cooling and use. In this known process, the methanization reactor is operated in such a way that the output stream has a temperature between 500 ° C and 700 ° C and is cooled to a temperature between 250 ° C and 350 ° C, yet at least 50 ° C above the dew point of the output stream. (i.e., the condensation temperature of the steam in the output stream) at its actual pressure and composition, and the recycle stream is then extracted from the output stream after this cooling and combined without further treatment with the input stream.

The critical point in this and other adiabatic reactions is the recirculation ratio, defined as the ratio of recirculation flow to synthesis gas flow. All else being equal, it is desirable to keep the recirculation ratio as low as possible, since the means used to drive the recirculation stream back to the synthesis gas stream, compressors and the like are extremely energy intensive and therefore have an adverse effect on the entire economy of the process.

It is important to show that by calculating a methanization process in which a portion of the cooled output stream is recycled to the reactor input where it is mixed with the synthesis gas stream, it can be shown that all else equal the recycle ratio U 2624 4 is reached when the temperature difference across the reactor, ie. the temperature difference between the input current and the output current increases.

Therefore, in order to achieve the most economical recirculation methanization process, one should strive to obtain as high a temperature in the output stream and as low a temperature in the input stream as possible. However, practical, kinetic and thermodynamic conditions limit how much temperature difference can be achieved over the reactor.

As mentioned, the two methanation reactions are very exothermic, and therefore at high temperatures there will be a tendency for the decomposition of methane into carbon monoxide and hydrogen. Of course, this will result in the higher the temperature of the output stream, the less the methane content will be when it is assumed that reactions (2) and (3) run to equilibrium. Therefore, it is not appropriate to operate the methanization process at temperatures above 700 ° C, usually even barely above 600 ° C. In practice, it will be difficult to produce a methanization catalyst which is both mechanically stable and has the appropriate catalytic activity after being exposed to such a high temperature for a long time. Suitable catalysts are disclosed in DK patent applications 2968/75 & 2388/76.

However, although the formation of methane is thermodynamically favored by low temperatures, the reaction rate sets a lower limit on how low a temperature in the inlet stream to the reactor it is reasonable to have. In practice, however, the lower limit has been found to be linked to the nature of the catalyst. Surprisingly, it turns out that the nickel catalyst is destroyed if the input temperature becomes too small. It is not known exactly where this limit is, but it is believed that if the inlet temperature of the reactor is kept below 280-300 ° C, then in the long run it will result in the destruction of the catalyst.

It is an object of the invention to provide a methanization process of the kind initially described, wherein the process economy is improved by reducing the recycling work by reducing the recirculation ratio. As will be understood from the foregoing, this may involve increasing the difference between the temperature of the input current and the output current. It will also be appreciated from the foregoing that other considerations contradict this increase in the temperature difference across the catalyst bed, with an excessively high starting temperature giving too low methane content in the exhaust gas and, moreover, causing problems with the stability and activity of the catalyst. too low an input temperature tends to destroy the nickel-containing catalyst. It is therefore the object of the invention, in particular, to provide a methanization process of the type indicated, in which the recirculation work is reduced and said temperature difference is increased by lowering the temperature of the input gas to the reactor relative to what was otherwise considered possible.

This is achieved if, according to the invention, the combined input stream is maintained at a temperature between 175 ° C and 250 ° C and the shift catalyst is nickel and iron free, is located directly in front of the methanation catalyst, calculated in the flow direction of the input stream, and is between 10 and 75%, preferably between 30 and 60% of the total catalyst volume '.

While using the aforementioned known methanization with recirculation, an operating temperature of the gas above: 250 ° C is employed, it has been found by the present process to lower this temperature below this value without destroying the nickel catalyst.

The invention should not be bound by any particular theory as to the reason why nickel catalysts for counteracting reactions are normally deactivated at input temperatures below 250-300 ° C, but do not in the present circumstances. However, the following explanation may contribute to the understanding. It is believed that the deactivation of the nickel catalyst that occurs at low input temperature is due to the formation of nickel carbonyl, Ni (CO) 4, by reaction of metallic Ni with gaseous CO. Thus, this nickel carbonyl, which is stable at the considered pressure at temperatures less than 290-300 ° C, can be formed at the inlet of the reactor. When transported with the gas to subsequent warmer zones in the reactor, there is probably a decomposition of the nickel carbonyl and thereby a deposition of nickel which gives rise to crystal growth. The hypothesis is supported by studies of the methanization catalyst from one experiment (inverse experiment 2) showing that there is a transport of nickel from the upper catalyst layer to warmer portions of the reactor. A quantitative analysis of samples taken from the top 30 cm of the catalyst showed a relative decrease in nickel content between 0 and 5%. A similar analysis on samples taken 50-60 cm from the top of the catalyst layer showed a relative increase in nickel content between 2 and 4%. Since the tendency for nickel carbonyl formation U2624 6 to depend on pressure and temperature depends on the concentration of CO present in the gas, the present process is thought to work as follows: The CO-rich gas stays at a low temperature where carbonyl formation over a nickel catalyst, first passed through a shift catalyst whose active metals do not form carbonyl compounds. The slightly exothermic shift process will slightly heat the gas as well as partially convert carbon monoxide.

When the gas is then passed over the methanization catalyst, the tendency for nickel carbonyl formation due to the low temperature rise and lower CO₂ content is substantially removed.

It is most convenient for the shift catalyst and the methanation catalyst to be in the same reactor and so that the former - using downstream gas flow direction, which would be the normal - forms a layer immediately on top of the latter. It provides the simplest device design and ensures the most intimate interaction between the two catalysts in the manner described.

However, in some cases it may be a rebuild of an existing methanisation plant where a methanisation reactor already exists, and such will usually only have the size needed for a given size plant. In this case, it may be convenient for the invention to have the two catalysts disposed in their respective reactors, which together act as an adiabatic working reactor complex, the reactor with the shift catalyst located directly in front of the reactor with the methanization catalyst, calculated in the combined flow stream direction.

Any known methanation catalyst containing nickel can be used as a methanization catalyst. The nickel will usually be on a hair's. This may consist of conventional materials such as alumina, silica, magnesium spinel or mixtures thereof. It is particularly advantageous if the carrier is partially made of zirconia as described in the aforementioned DK patent application 2968/75. The nickel may be particularly advantageously promoted with molybdenum as described in DK Patent Application No. 2388/76. Optionally, part of the nickel may be replaced by cobalt. However, it is not essential for the good effect of the process according to the invention if used with some nickel-containing catalyst.

The shift catalyst may also be a conventional shift catalyst. Such usually contain at least two of the metals Cu, Zn and Cr, optionally 1 form of oxides and optionally on a support. The shift catalyst it contains is of no importance to the beneficial effect of the present process, provided that the shift catalyst does not contain metals forming metal carbonyls.

It should be mentioned that it is known per se to place a shift catalyst and a methanization catalyst in the same reactor, German Patent Publication No. 2432887. It describes a process for producing a methane-rich natural gas substitute gas from a coal, tar or Residual oils by gasification produced crude gas, after purification and leaching to a content of CO 2 of less than 2% the vaporized gas was adjusted to a volume ratio of water vapor: carbon monoxide of 0.55: 1 to 1: 1 and conducted this gas at an inlet temperature of 300 -500 ° C through a reaction zone containing a shift catalyst and a methanization catalyst in the volume ratio of approx. 1: 4 to 1:10. '

Thus, what happens in the known method is something; other than by the present process, namely, placing the shift reaction and the methanation reaction to the same reactor, and substantially saving space, in particular reducing the reactor volumes themselves, while utilizing the water generated by the methanation reaction. (see preceding reaction equations (2) and (3)) scan raw material in the shift reaction (equation (1)).

By contrast, by the known method, the inlet temperature of the methanization reactor is not lowered and the recirculation ratio must therefore be high. It should be emphasized that in the present process a shift reactor is not avoided if the process is part of a process for the production of SNG by gasification of coal, tar or heavy oil, purification, shift reaction and methane leaching, the starting material for the In the present method, a gas containing hydrogen and carbon monoxide in the ratio of 3: 1 and, for example, hair, has been subjected to the shift reaction. ? German Publication No. 2432887 mentions that as a shift catalyst, a mixture of iron oxide and chlorine oxide can be used, and scat methanation catalyst nickel on a water-resistant support material. In the example, however, there is only one catalyst, namely a methanization catalyst of 50 wt% nickel on a magnesium spinel support which has acted as both shift and methanation catalyst. Both iron and nickel form metal carbonyls, but this has hardly happened in the known process since the temperature of the reactor, even at the inlet, has been too high for the difficulties explained above to occur; as mentioned, the writing requires an inlet temperature of 300-500 ° C; the example mentions only exit temperatures which are 460 ° C.

The process according to the invention will be explained in more detail below with reference to the drawing, which shows a principle diagram for the preparation of a methane-rich gas from coal, oil or tar by gasification, purification, shift reaction and methanisation. As will be understood, the invention relates only to the latter step.

Raw material such as coal or a heavy oil fraction, e.g., a heavy residual oil or fuel oil, is passed through a conduit or other feed means to a gasification zone 12 where the raw material is gasified under high pressure. The gas formed is passed through a conduit 13 to a purification stage 14, thence through a conduit 15 to a shift reactor 16 and thence through a conduit 17 to a washing zone 18. In the washing zones 14 and 18, catalyst poisoning and the like are removed. and leaching of carbon dioxide; one of them, preferably zone 14, may optionally be omitted. From the wash zone 18, the gas enters a conduit 19 with a volume ratio of Hg to CO of substantially 3: 1; the gas here may contain alternating amounts of water vapor and very small amounts of carbon dioxide and methane. The gas is passed through conduit 19 to a methanization reactor 20 having an inlet temperature here of 175-250 ° C. The methanization reactor is of a conventional type per se, but while in conventional methanization reactions a methanization reactor contains only one methanization catalyst, usually with nickel as the main component of the catalytically active material, the present reactor contains at the bottom a layer 22 with a methanization catalyst and immediately on top of it. a layer 21 with a shift catalyst. It is hereby assumed that the direction of the gas flow as shown by arrows is from top to bottom; if the flow direction is the opposite, the location of the two catalysts must be changed, the shift catalyst having to be at the reactor entrance. The two catalysts are separated by a boundary layer 23. This may or may not be marked in terms of apparatus, for example by means of a wire mesh. The cassette leaves the metization reactor 20 through an output nip 24 with a temperature between 500 and 700 ° C, preferably at most <S00 ° C. It is cooled in a cooling means 25 and then divided into two streams, an output stream 26 which is passed on for further processing or use, and a recirculation stream 27 which is returned to the inlet joint IS, where it is mixed with fresh synthesis gas from the shaker 16 and washing stream). 18. As explained above, the recirculation flow must be as small as possible and the inlet temperature of the recirculation stream and the flow of the fresh synthesis gas inlet conduit 19 must be as low as possible within the specified limits; as also explained, these two parameters are related.

The method according to the invention is further elucidated by an experiment (hereinafter referred to as experiment # 3), which is compared with two comparative experiments (1 and 2). Three trials were conducted as long-term trials in a pilot plant there; only included a methanization reactor of basically the same kind as the reactor 20 in the drawing, as well as refurements, etc. No gasification, purification and shift conversion were thus carried out. The reactor used was, moreover, for the purpose of the probe being equipped with a thermocouple located congruent with the reactor in the denries axis. The thermocouple inserted at the bottom of the reactor runs along the axis all the way to the top of the reactor. In such tpmolonaje, ternometers can be placed and thus determine the temperature profile in the reactor.

The reactor dimensions were:

Material AISX 316 * Length 2800 mm

Inner diameter 50 isn

Outer diameter 63 mm

Thermal pocket outer diameter · 8 mro 2

Free cross-sectional area 191Q mm x Standard steel alloy according to the American Iron and Steel Institute containing, inter alia, chrome, nickel and molybdenum.

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The reactor was operated adiabatically in all three experiments, ie. without heat exchange with the surroundings. In addition to insulation, the reactor was wrapped with an electric heater, with the help of which was compensated for the heat loss which is impossible to avoid in practice.

A total of three different catalysts were used during the experiments.

For methanization, two different catalysts designated M1 and M2 were used. Both catalysts are conventional nickel catalysts with .25% W / W nickel on a ceramic support. M2 differs from Ml only by containing 2% W / W molybdenum. The content of molybdenum in the catalyst M2 does not affect the test results and has no bearing on the invention. As mentioned, it is not linked to the use of a particular type of methanization or shift catalyst, only the former contains nickel and the latter does not.

The shift catalyst used. S, was a conventional Cu-Zn-Cr catalyst.

The synthesis stream was prepared for practical reasons by cracking methanol over a catalyst followed by the addition of hydrogen to a hydrogen-carbon monoxide ratio of 3.

Contrary to what one would do in normal operation, the recirculation flow was, for practical reasons, driven back to the egas flow by a compressor, designed to operate at temperatures below approx. 100 ° C. For this reason, the output stream from the reactor was cooled to below 100 ° C as opposed to what would be done in normal operation, thereby condensing out all the water vapor.

After separating the water, the dry starting stream is divided into a recycle stream and a product stream, after which heating to the desired temperature is carried out. To the recycle stream, after compression, the calculated amount of water it would have contained had it not been necessary to remove it prior to compression was added. An amount of water corresponding to the difference between condensed water and water added to the recycle stream is thought to be added to the product stream. In this way, the real operating conditions that would have prevailed if a compressor had not been used as mentioned above were simulated.

Table 1 shows operating parameters for the three experiments. The experiments, as seen in the table, were very lengthy. Over such a long period of time, it is not possible to keep the individual temperatures, the composition and volume of the gas streams as well as the pressure constant. Small oscillations around the desired values in the thigh cannot be avoided. Therefore, at regular intervals, measurements of the various temperatures, pressures and gas flow volume velocities were made in each experiment. The gas streams were also analyzed and their contents of various components determined. Thus, there are three such measurements listed in Table 1. The individual measurements are selected because they are estimated to represent an average of measurements taken during each trial.

It can be seen from the table that the compositions of the recycle stream and the product stream are not the same as they should theoretically be. This is partly due to measurement uncertainty and partly that it is not possible to calculate exactly how much water vapor is to be added in the simulation after compression in the simulation.

The catalyst used in experiment # 1 was Ml. The amount of catalyst in the reactor was 3.3 liters, corresponding to 1730 mm of the reactor being filled with Ml. The experiment could be conducted in it, shown for a long time virtually without change.

In experiment # 2, a combination of the two methanation catalysts was used. A total of 1.15 liters of Ml and 2.16 liters of M2 were used. The catalysts were distributed such that from the entrance to the reactor the first 500 mm ml, the next 1130 mm m2 and the last 100 mm again were ml.

In experiment # 3, a combination of the shift catalyst S and M2 was used. The amounts of catalyst used were 0.95 liters and 2.85 liters, respectively. The position in the reactor was calculated from the entrance so that the first 500 mm was S and the subsequent 1500 mm was M2. As a result, the ratio of shift to methanation catalyst was 1: 3 on a volume basis. Calculated on a weight basis, the ratio is 1: 3.5, with one liter of each catalyst weighing 1040 g and 1228 g in reduced state, respectively.

During experiment # 2, it was observed that the deactivation of the methanization catalyst began in the wood flake between the input gas to the reactor and the catalyst and, over time, spread through the catalyst layer.

The reason that the deactivation of the methanization catalyst is not reflected in reduced methane content in the product stream is that the reactor contains methanization catalyst in excess.

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As a result, the amount of active catalyst was constantly so high that the methanation reactions (2) and (3) could reach equilibrium.

As a measure of the deactivation of the catalyst at the end of the experiment, the average diameter of the catalyst nickel crystals was used, as crystal growth in these experiments can be taken as an expression of deactivation.

No deactivation of the catalyst was observed during Experiment # 1 beyond what normally occurs. Upon opening the reactor after completion of the experiment, it was observed that the catalyst appeared to be in good condition. Closer examination of the top 50 cm layer revealed that the average nickel crystal diameter was 200 Å. The nickel crystal diameter of the unused catalyst varies between 140 and 180 Å.

Despite the shorter test time, after the end of experiment # 2, a deactivation of the catalyst was observed. Thus, upon inspection of the catalyst, a general discoloration as well as a degradation of parts of the upper catalyst layer were noticed. An analysis of the upper 50 cm long catalyst layer showed that the nickel crystal diameter varied between 10,000 and 20,000 Å. .

After Experiment # 3, both the shift and the methanation catalyst were inspected. Both the shift catalyst and the methanation catalyst had normal appearance. An analysis of the top 50 cm long layer of the methanization catalyst showed that the average nickel crystal diameter was 200 Å.

Table 1

Attempt _: _.__ 1 _2_ 3

Synthesis gas flow, Nm3 / h 12.65 13.28 12.17% H2 74.6 73.0 72.9% CK4 0.2.0.4 0.3 I CO 24.5 25.8 25.7% CO 2 O, 7 0.8 1.1

Input flow, Nm3 / h 50.87 49.53 38.77% H2 35.3 33.5 36.5% CH4 29.1 29.6 27.9% CO 6.8 7.6 7.7 14262Λ 13

Table 1 (continued)

Test _ 1 2_ 3% CO 2 2.8 3.4 2.9% H 2 O 26.0 25.9 25.0

Recycle flow, Nm3 / h 38.22 36.25 26.6% H2 22.1 18.7 19.1% CH4 39.0 41.1 40.1% CO 0.7 0.8 0.6% CO 2 3 , 4.5 4.5 3.7% H2 O 34.5 35.0 36.5

Product flow, Nm3 / h 7.39 7.05 6.84% H2 22.2 19.2 19.1% CH4 39.2 42.3. 40.0% CO 0.7 0.8 0.6% CO2 3.6 4.6 3.7% H2 O 34.2 33.0 36.5

Catalyst Ml M1 + M2 S + M2

The temperature of the input current, ° C 300 245 226

The output current temperature, ° C 598 601 593

Pressure, kg / cm 2 g 31.2 30.5

Recycle ratio Nm3 / Nm3 3.0 2.7 2.2

Duration of trial, h 3000 1600 1175

Time from start of experiment to measurement, h_800_739 136

From Table 1, it is seen that the temperature of the input gas stream in Experiment 1 was 300 ° C and the recirculation ratio approx. 3. These are quite conventional parameters and, as seen, the process can go undisturbed for a very long time. The disadvantage of this is initially explained, that such a high recycling ratio is uneconomical. The effort to reduce the input temperature by decreasing the recirculation ratio somewhat with a conventional methanization catalyst which in Experiment 2 failed in so far as the just explained destruction of the catalyst occurred. In other words, the experiment could not have been carried out for a significantly longer time as Experiment 1 could. In Experiment 3, the inlet temperature 14262Λ 14 was further reduced and the recirculation ratio to just over 2, and despite this, thanks to the presence of the shift catalyst, the experiment could be conducted without destroying the catalyst.

Γ Experiment # 3 also measured temperature and gas composition in the transition layer between the shift and methanation catalyst. The measurements were made at the same time as those in Table 1, and the result was as follows:

Temperature 28 ° C

Gas composition:% H2 41.6% CH4 27.9% CO 2.6% CO2 8.0% H2 O 19.9

These figures support the hypothesis on the effect of the shift catalyst in the methanization reactor listed above.