METHOD AND DEVICE FOR CARRYING OUT A WATER-GAS SHIFT REACTOR
20210246021 · 2021-08-12
Assignee
Inventors
Cpc classification
Y02P20/52
GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
C10K3/04
CHEMISTRY; METALLURGY
C01B3/48
CHEMISTRY; METALLURGY
C01B2203/0233
CHEMISTRY; METALLURGY
International classification
Abstract
A process for performing the water gas shift reaction wherein raw synthesis gas is reacted in the presence of steam and at least one water gas shift catalyst to convert carbon monoxide into carbon dioxide and to form hydrogen. The raw synthesis gas is initially passed through at least one unit for high-temperature CO conversion and subsequently, downstream thereof, passed through at least one unit for low-temperature CO conversion. After passing through the at least one unit for high-temperature CO conversion the synthesis gas stream is divided into at least two substreams. The first substream is passed through a first unit for low-temperature CO conversion and the second substream is passed through a second unit for low-temperature CO conversion, wherein both units for low-temperature CO conversion are arranged in parallel relative to one another.
Claims
1.-17. (canceled)
18. A process for performing the water gas shift reaction, comprising: reacting raw synthesis gas in the presence of steam and at least one water gas shift catalyst to convert carbon monoxide into carbon dioxide and to form hydrogen, wherein the raw synthesis gas is initially passed through at least one unit for high-temperature CO conversion and subsequently, downstream thereof, passed through at least two units for low-temperature CO conversion, wherein after passing through the at least one unit for high-temperature CO conversion the synthesis gas stream is divided into at least two substreams, wherein the first substream is passed through a first unit for low-temperature CO conversion and the second substream is passed through a second unit for low-temperature CO conversion, wherein both units for low-temperature CO conversion are arranged in parallel relative to one another.
19. The process of claim 18 wherein, after flowing through the at least one unit for high-temperature CO conversion, the synthesis gas stream is divided into two substreams which each comprise a proportion of 40% by volume to 60% by volume of the total synthesis gas stream after the high-temperature CO conversion.
20. The process of claim 18 wherein, after flowing through the at least one unit for high-temperature CO conversion, the synthesis gas stream is divided into two substreams of unequal size, wherein a first substream comprises a proportion of 95% by volume to 65% by volume of the total synthesis gas stream after the high-temperature CO conversion and a second substream comprises a proportion of 5% by volume to 35% by volume of the total synthesis gas stream after the high-temperature CO conversion.
21. The process of claim 18 wherein the entry temperature of the raw synthesis gas into the high-temperature CO conversion is in the range of 350-400° C.
22. The process of claim 18 wherein the maximum exit temperature of the synthesis gas stream from the high-temperature CO conversion is in the range of 430-450° C.
23. The process of claim 18 wherein after exiting the high-temperature CO conversion and before entering the two units for low-temperature CO conversion the synthesis gas stream is cooled by means of a cooling unit so that the entry temperature upon entry into the two units for low-temperature CO conversion is in the range from 180° C. to 220° C.
24. The process of claim 18 wherein the maximum exit temperature of the synthesis gas stream from the two units for low-temperature CO conversion is in the range from 220 to 240° C.
25. The process of claim 18 wherein the entry concentration of carbon monoxide in the raw synthesis gas upon entry into the high-temperature CO conversion is in the range from 10 mol % to 16 mol % and the exit concentration of carbon monoxide in the synthesis gas stream upon exiting the high-temperature CO conversion (2) and upon entering the low-temperature CO conversion is in each case in the range from 3 mol % to 5 mol %.
26. The process of claim 18 wherein the exit concentration of carbon monoxide in the synthesis gas stream upon exiting the low-temperature CO conversion is less than 0.8 mol %.
27. The process of claim 18 wherein the water gas shift reaction in the high-temperature CO conversion is carried out in the presence of an iron-containing catalyst which contains not only iron but also chromium and copper, where the latter are present in the catalyst in smaller amounts than iron.
28. The process of claim 18 wherein the water gas shift reaction in the low-temperature CO conversion is carried out in the presence of a copper-containing catalyst which contains not only copper but also zinc and aluminum, where the latter are present in the catalyst in smaller amounts than copper.
29. An apparatus for performing the water gas shift reaction, comprising: at least one high-temperature CO conversion unit, at least one first and one second low-temperature CO conversion unit connected downstream of the high-temperature CO conversion unit in the flow path, at least one cooling unit configured to cool the product gas mixture withdrawn from the high-temperature CO conversion unit, wherein provided in the conduit system downstream of the cooling unit is a means for dividing the product gas mixture into two substreams, and wherein the first low-temperature CO conversion unit and the second low-temperature CO conversion unit are arranged in parallel and downstream of the means for dividing such that each first and second low-temperature CO conversion unit is traversable by only one of the two substreams.
30. The apparatus of claim 29 wherein said apparatus comprises an exit conduit from the high-temperature CO conversion unit that divides into a first conduit for a first substream which leads to the first low-temperature CO conversion unit and a second conduit for a second substream which leads to the second low-temperature CO conversion unit.
31. The apparatus of claim 29 the second low-temperature CO conversion unit has only a fraction of the size and/or the plant capacity of the first low-temperature CO conversion unit and/or the second low-temperature CO conversion unit contains only a fraction of the catalyst amount present in the first low-temperature CO conversion unit.
32. The apparatus of claim 31 wherein the fraction of the size and/or the plant capacity and/or the catalyst amount of the second low-temperature CO conversion unit relative to the first low-temperature CO conversion unit is in the range from 0.15 times to 0.4 times.
33. The apparatus of claim 29 wherein said apparatus comprises at least one second cooling unit which, in the flow path, is arranged downstream of the first low-temperature CO conversion unit and does not have the product gas mixture exiting the second low-temperature CO conversion unit flow through it and/or said apparatus comprises at least one third cooling unit for cooling the product gas mixture which, in the flow path, is arranged downstream of the second low-temperature CO conversion unit and does not have the product gas mixture exiting the first low-temperature CO conversion unit flow through it or comprises a common cooling unit both for the product gas mixture exiting the first low-temperature CO conversion unit and for the product gas mixture exiting the second low-temperature CO conversion unit provided downstream of these the first and second low-temperature CO conversion units.
Description
BRIEF DESCRIPTION OF THE FIGURES
[0057]
EMBODIMENTS OF THE INVENTION
[0058] The variant of an apparatus 1 according to the invention shown in
[0059] The present invention is more particularly elucidated hereinbelow with reference to specific examples.
[0060] In the water gas shift reaction a distinction is made between three types which are classified and described according to the reaction temperatures prevailing in each case, namely high-temperature shift (HT shift or HTS for short),
[0061] medium-temperature shift (MT shift for short) and
[0062] low-temperature shift (LT shift or LTS for short), wherein medium-temperature shift has no particular significance for the process according to the present invention.
[0063] The following operating conditions are contemplated by way of example for the two types of shift reaction relevant here: [0064] High-temperature shift [0065] entry temperature 350-400° C. (370° C. in typical design) [0066] exit temperature 400-440° C. (limited, to prevent sintering of the catalyst) [0067] entry concentration of CO 13 mol % (excluding water) [0068] exit concentration of CO 4 mol-% [0069] catalyst: KATALCO 71-5 from Johnson Matthey Catalyst [0070] composition: Fe (88%), Cr (9%), Cu(3%) [0071] Low-temperature shift [0072] entry temperature 180-220° C. (200° C. in typical design) [0073] exit temperature 200-230° C. (limited, to prevent sintering of the catalyst) [0074] entry concentration of CO 3-5 mol % (excl. water) [0075] exit concentration of CO 0.2-0.5 mol % [0076] catalyst: KATALCO 83-3 from Johnson Matthey [0077] composition: Cu (51%), Zn (31%), Al (18%) [0078] The medium-temperature shift also used in the prior art is preferably run with the following operating conditions: [0079] entry temperature 200-230° C. [0080] exit temperature 300-350° C. (limited, to prevent sintering of the catalyst) [0081] entry concentration of CO for example 9 mol-% (excl. water) [0082] exit concentration of CO for example 0.5 mol % mol-%.
COMPARATIVE EXAMPLE
[0083] In this first example the capacity of an existing plant for performing the water gas shift reaction is to be expanded. This exemplary variant of the invention is hereinbelow referred to as case 0. Capacity expansion is to be carried out according to the following parameters. [0084] Original capacity: 3300 t/d of ammonia [0085] Capacity expansion: 20% ->3960 t/d (tons per day) of ammonia [0086] HT shift: 82.6 m.sup.3 of KATALCO 71-5 (see above) catalyst [0087] LT shift: 139 m.sup.3 of KATALCO 83-3 (see above) catalyst [0088] Calculated lifetime 5 years
TABLE-US-00001 CO.sub.2 CO Dry gas H.sub.2O Temperature mol % (dry) kmol/h ° C. HTS inlet 6.96 13.4 18757 8894 370 HTS outlet 15.18 3.38 20574 7077 442 LTS inlet 15.18 3.38 20574 7077 200 LTS outlet 17.69 0.33 21199 6451 226
Capacity Expansion
[0089] It would in principle be conceivable to increase the capacity of the plant to a certain extent without performing alterations to the reactors for the CO shift. The increased catalyst space velocity, i.e. the amount of gas that is treated by an amount of catalyst per hour, would increase, thus shortening the contact time of the gas with the catalyst and leading to an increase in the exit concentration of carbon monoxide due to a greater distance from equilibrium. This is not particularly restrictive with fresh catalyst but markedly reduces the lifetime of the catalyst.
[0090] The lifetime of shift catalysts is limited by progressive poisoning by chlorides and other catalyst poisons. The amount of catalyst is therefore chosen such that after the specified lifetime the desired exit concentration may still be achieved by adding to the catalyst amount required therefor an amount which is inactive at the end of this lifetime due to poisoning. Catalyst employed beyond this required amount is available as a reserve and increases the lifetime of the bed. The poisoning therefore has the effect that fresh catalyst allows exit concentrations markedly below the design value since more than the required amount of catalyst is available and said concentration exponentially approaches the design value towards the end of the lifetime. In practice the catalyst is operated until the pressure drop associated with the poisoning renders operation uneconomic or the elevated exit concentration of carbon monoxide results in losses of hydrogen high enough to render operation uneconomic.
[0091] A capacity expansion by 20% to 3960 t/d where the lifetime of the catalyst is to be maintained therefore requires the provision of additional catalyst. Since an expansion of ammonia production results in a largely linear increase in the process gas stream through the CO shift, the stream at the entry to the HTS would be about 526271 kg/h after the expansion. After a lifetime of 5 years the setup of a parallel LTS reactor comprising 34.8 m.sup.3 of catalyst results in the same exit concentration of 0.33 mol % (dry) of carbon monoxide as in the main tract when 20% of the total gas stream are passed through the parallel secondary tract. The pressure drop over the parallel tract is dependent on reactor geometry but in this example should be precisely equal to that over the main tract. This case is hereinbelow referred to as case 1.
[0092] Especially in the course of a revamp, but in some cases also in the course of planning a new plant, divergence from this concept may be advantageous. If the low-temperature shift reaction is performed in two parallel reactors this may have positive effects on: [0093] plant capacity [0094] exit concentration [0095] pressure drop [0096] service life of the catalyst
[0097] The objectives according to which these aspects are desired in the present case determines the size of the parallel reactor and the amount and distribution of the catalyst.
[0098] Sticking with the above example of a 20% capacity increase the following three cases may be distinguished.
EXAMPLE 1 (Inventive)
[0099] Parallel low-temperature shift with 16.7% of the total stream and the amount of catalyst required according to the configuration.
[0100] The plant capacity is increased by treatment of the additional gas exclusively in the parallel reactor. The required amount of catalyst is more than 20% of the previous volume since elevated flow causes the exit concentration of the high-temperature shift to fall and additional catalyst is required to compensate this. Based on the above example in which 34.8 m.sup.3 of catalyst are required to achieve a 20% production increase, only 139 m.sup.3*0.2=27.8 m.sup.3 are required to achieve the exit concentration and the remaining 7 m.sup.3 are required for compensation.
[0101] In this case the exit concentration is just as high as in the original plant since the amount of the additional catalyst corresponds to the additional amount of gas. c.sub.CO=0.33 mol % (dry)
[0102] The pressure drop over the parallel unit depends on the reactor geometry, i.e. on the height of the catalyst bed. Since the pressure drop over the existing tract is unchanged an optimization of the new tract to below this value is unnecessary since both tracts should have the same pressure drop to achieve the desired flow ratios. In the present example the pressure drop Δp=0.59 bar.
[0103] The service life of the catalyst is not affected since the amount of employed catalyst corresponds to the amount of additional gas. The expected lifetime remains at 5 years.
EXAMPLE 2 (Inventive)
[0104] Parallel low-temperature shift with 35.2% of the total stream and more than the required amount of catalyst according to the configuration.
[0105] This variant is based on providing more than the amount of catalyst strictly necessary according to the configuration and partial relocation of process gas into the parallel tract. This reduces the catalyst space velocity and the increased contact time of the gas with the catalyst allows the reaction to better approach equilibrium.
[0106] Since the plant capacity is determined by the amount of the process gas flowing through it this variant has only indirect advantages over example 1.
[0107] The increased amount of catalyst allows a lower exit concentration to be achieved since the reaction is brought closer to its equilibrium. This results in reduced losses of hydrogen for reaction with the remaining carbon monoxide. Increasing the amount of catalyst to 75.6 m.sup.3 reduces the exit concentration of the low-temperature shift to 0.23 mol %. In the present example in the further process this saves 104 kmol/h of hydrogen which would have been necessary to react with the difference in carbon monoxide. This amount of hydrogen could in turn be used to produce 69 kmol/h of ammonia, or an amount of 28.3 t/d.
[0108] Due to the elevated amount of catalyst compared to example 1 the design of the reactor must be chosen with care to avoid an additional pressure drop. If this is constructionally no longer possible a slightly higher pressure drop will be established in the parallel tract.
[0109] The elevated amount of catalyst increases the lifetime of the two beds to 6.8 years until the desired exit concentration of 0.3 mol % (dry) is exceeded.
EXAMPLE 3 (Inventive)
[0110] Two parallel low-temperature shift apparatuses with division of the synthesis gas stream in the same ratio into two substreams, passage of respectively 50% of the total stream through each of the two parallel low-temperature CO shift apparatuses and use of the amount of catalyst required according to the configuration.
[0111] This variant is based on the concept that process gas passes through the parallel tract instead of through the existing tract. The maximum effect is achieved when the distribution between the streams is precisely 50:50. The amount of catalyst must be at least high enough to ensure that it can treat half of the total amount of process gas. An appropriate amount of catalyst is withdrawn from the existing tract. In the present example both reactors would therefore be filled with 86.9 m.sup.3 of catalyst respectively.
[0112] This variant has no direct effect on plant capacity since only a spatial relocation of the reaction is carried out. However, the reduction in pressure drop allows for increasing of the maximum possible capacity.
[0113] As in the two previous cases any possible reduction in the exit concentration of carbon monoxide is dependent on an additionally introduced amount of catalyst. This is possible but would run counter to the fundamental concept of this case.
[0114] The reduction in the amount of catalyst and the accompanying reduction in the bed height in the existing tract markedly reduces the pressure drop over the low-temperature shift. In this example the pressure drop over the catalyst bed is reduced from 0.59 bar at 139 m.sup.3 to 0.11 bar at 86.9 m.sup.3 per bed. This reduction in the pressure drop saves energy when operating the upstream compressors which in turn has a positive effect on the economy of the plant. Since in the case of capacity increases the overall pressure drop over the plant may approach the maximum pressure increase of the upstream compressors and these thus limit the maximum capacity of the plant, a setup as in this case can increase the maximum plant capacity by reducing the pressure drop. In this case the maximum capacity is reached at 120.4% of the nominal capacity since the pressure profile becomes limiting. By reducing the pressure drop by setting up the low-temperature shift as in this case, the maximum capacity of the plant may be increased to 122.4% of the nominal capacity which corresponds to an amount of ammonia of 63.2 t/d. Not only the achievable exit concentration but also the lifetime of the low-temperature shift are affected by additional amounts of catalyst. This is possible but would run counter to the general idea of this case.
[0115] The following table summarizes the three examples described hereinabove and the comparative example:
TABLE-US-00002 Flow distribution Catalyst c.sub.CO Pressure drop Lifetime [%] [m.sup.3] [mol % (dry)] [bar] [y] Variant LTS 1 LTS 2 LTS 1 LTS 2 LTS 1 LTS 2 LTS 1 LTS 2 LTS 1 LTS 2 0 83.3 0 139 0.33 0.59 5 1 83.3 16.7 139 34.8 0.33 0.33 0.59 0.59 5 5 2 64.8 35.2 139 75.6 0.23 0.23 0.4 0.4 6.8 6.8 3 50 50 86.9 86.9 0.33 0.33 0.11 0.11 5 5
[0116] The results of the examples reported above show that in the context of the present invention it is advantageous for example in the construction of new plants or in the course of revamps to provide existing plants with an additional parallel low-temperature shift apparatus in which 0-50% of the total catalyst amount is employed and in which 0-50% of the total synthesis gas stream is treated.
[0117] Sought in the context of the invention are capacity increases of up to 100% of the previous capacity, preferably of at least about 10%, for example 10% to 13%, in order to achieve an increase in ammonia production. However, in advantageous cases capacity increases of for example up to 20% are quite realistic.
[0118] It is a further object of the invention to reduce the pressure drop by relocation of the process gas from the primary tract, in some cases even with reduction of the catalyst amount in the primary reactor, to 0.1 to 0.5 bar, preferably 0.1 to 0.4 bar, depending on the application in order thus to improve the energy efficiency of the process.
[0119] It is a further objective of the invention to reduce the exit concentration of carbon monoxide from 0.05 to 0.5 mol % (dry), preferably to about 0.2 mol % (dry), which results in an increase in the ammonia yield.
[0120] It is a further objective of the invention to extend the lifetime of the plant, until the maximum allowed exit concentration is achieved, by up to 5 years, preferably by at least about 2 years, in order thus to reduce catalyst costs.
[0121] Appropriate provision of catalyst and distribution of the gas stream makes it possible in the context of the present invention to optimize the important operating parameters plant capacity, pressure drop and lifetime for the individual usage case, with positive side effects on ammonia production and the energy balance of the process.