CA2546181C

CA2546181C - Process for the preparation of urea

Info

Publication number: CA2546181C
Application number: CA2546181A
Authority: CA
Inventors: Poul Erik Hoejlund Nielsen; John Boegild Hansen; Svend Erik Nielsen; Niels Christian Schioedt
Original assignee: Haldor Topsoe AS
Current assignee: Topsoe AS
Priority date: 2003-11-17
Filing date: 2004-11-11
Publication date: 2011-01-11
Anticipated expiration: 2024-11-11
Also published as: CA2546181A1; CN1906156A; JP2007511481A; WO2005049554A1; CN100457724C; KR20060113712A; KR101123052B1; BRPI0406779A; BRPI0406779B1; JP4653108B2

Abstract

A process for the preparation of urea comprising adding methanol to a carbon monoxide containing gas, contacting the carbon monoxide-containing gas, methanol and water in at least one shift step in the presence of a catalyst comprising copper, zinc and aluminium and/or chromium to produce a stream rich in carbon dioxide, transferring the carbon dioxide-rich stream from the at least one shift step to a urea reactor and reacting the carbon dioxide with ammonia to produce urea.

Description

Process for the Preparation of Urea The invention relates to a process for the preparation of urea and provides a facile process for boosting capacity of fuel-based hydrogen plants and ammonia and/or urea plants.
Hydrogen plants as well as ammonia plants can utilise fuels such as natural gas, liquid hydrocarbons or solid fuels like coal or biomass. In these plants, hydrogen production takes place in four consecutive procedures - feed purifica-tion followed by steam reforming (or gasification), water gas shift (WGS) and purification. These procedures are fur-ther described in Kirk-Othmer and Ullman. Ammonia produc-tion is described in depth by Ib Dybkjaer in Ammonia, Ca-talysis and Manufacture, Springer-Verlag, Berlin Heidel-berg, Chapter 6, 1995, Ed. A. Nielsen. Urea production us-ing conventional methods is described in Ullmann's Encyclo-pedia of Industrial Chemistry, 6th Ed. 2002, Wiley-VCH.
The WGS reaction is described in the following equation:
CO + H~0 -~ C0~ + H2 ( 1 ) It is a slightly exothermic reaction used for producing more hydrogen. Known WGS catalysts in industrial high tem-perature shift (HTS) applications are high-temperature catalysts that are chromium-supported and iron-based, and they are sometimes promoted with copper. The operational range for the HTS is typically 340-360°C inlet temperature and with exit temperatures that are approximately 100°C
higher. The operational range of the inlet temperature for low temperature shift (ZTS) catalysts is from 200°C (or 20°C above the dew point of the gas). The inlet temperature should be kept as low as possible. Further details on cata-lysts for shift reactions and operating temperature are given in Catalyst Handboo k, 2. Ed. Manson Publishing Ztd.
England 1996.
In addition to these catalysts, Haldor Topss~e A/S has mar-keted a medium-temperature shift catalyst that is Cu-based and capable of operating a t temperatures up to 310°C. Vari-ous vendors offer sulphur- tolerant catalysts for the gasi-fication-based plants. However, these plants are not widely used for hydrogen producti on.
Methanol is produced on a large scale of more than 30 MM
t/y. Basically methanol is produced in very large plants with capacities of more than 2000 MTPD at places where natural gas is cheap. The production cost for methanol at places with cheap natural gas is estimated to be in the or-der of 60-80 USD/MT.
In the future, it is expected that methanol can be avail-able in large quantities and to a price that on an energy basis might be significant ly lower than the oil price.
In recent years there have been numerous studies of steam reforming of methanol for producing hydrogen and in par-ticular hydrogen for fuel cells. The disadvantage of the steam reforming process is that the heat of reaction has to be supplied through a wall and the equipment as such be-comes cumbersome.
Catalysts for low temperature steam reforming of methanol are copper based or option ally based upon noble metals.

Some companies, for instance Haldor Topss~e A/S, offer com-mercial products.
U.S. Patent No. 5,221,524 describes a hydrogen production process where a reformed gas is cooled before undergoing a low temperature shift reaction catalysed by a copper cata-lyst with an inlet temperature of 205°C. Liquid methanol is dispersively supplied to the shift converter and uncon-verted methanol is recycled to the methanol supply source and the shift reactor. The catalyst has activity both for low temperature shift conversion of carbon monoxide and the steam reforming reaction of methanol to hydrogen and carbon dioxide. The heat generated from the shift conversion reac-tion is utilised to accelerate the endothermic reaction for methanol decomposition.
U.S. Patent Application No. 2001/0038816 describes a gas generator for generating hydrogen utilising a shift reactor supplied with a reformed gas and water containing small amounts of methanol for frost protection. The gas generator is connected to a fuel cell set-up.
JP Patent Application No. 59203702 describes a hydrogen manufacturing process, whereby methanol and steam are re-acted in a shift reactor and the effluent gas is purified and hydrogen is removed. The remaining gases are combusted and the heat generated is used as a heat source for the methanol decomposition in the shift reactor.
JP Patent Application No. 3254071 describes a process for modifying alcohol and generating hydrogen for a fuel cell.
Natural gas is reacted with air in a methanol modifier and the heat generated is used for conversion of the metha-nol/water mixture.
It is an objective of the rove ntion to provide a process for production of urea by utilising a catalyst capable of operating at a wide range of temperatures.
According to the invention, th a re is provided a process for the preparation of urea as claimed in claim 1.
The process can be carried out by adding methanol to the feed stream to a water gas shift reactor containing a Cu-based catalyst comprising zinc, aluminium and/or chromium and resulting in a catalytic decomposition of the methanol along with the water gas shift reaction. In the isothermal case, the heat released by the exothermic Water Gas Shift Reaction balances the heat use d for the endothermic steam reforming of methanol. The Ben Bible heat in the feed streams may further be used in the process whereby a sig-nificant larger amount of meth anol may be steam reformed.
The catalyst used in the process of the invention is capa-ble of operating both at lower temperatures and at tempera-tures above 350°C.
The catalyst is suitable for a rea production and use of this catalyst provides a boost in the carbon dioxide pro-duction.
Besides this, by using this catalyst in the process the hy-drogen production from the uni t may be boosted up by fac-tors of 1-3. Alternatively the process can be used to de-crease the load on the reforming section. A capacity in-crease of ammonia plants is also provided by applying the process of the invention in such a plant.
The endothermic methanol steam reforming reaction:
CH30H + H~0 -~ 3H2 + C02 ( 2 ) obtains the necessary heat of reaction from the sensible heat in the gas as well as from the latent heat from the WGS reaction. The catalyst utilised in the process of the invention tolerates the maximum inlet temperature and is still active at a much lowe r temperature primarily deter-mined by the desire to keep the outlet methanol conoentra-tion as low as possible (typically in the temperature range from 240-320°C).
Experiments with addition of methanol to iron-based shift catalyst have shown that a significant amount of methane formation takes place on these catalysts. This is also the result of the large scale production of town gas using the Hytanol process developed by Lurgi.
The invention is applicable to a hydrogen plant and a urea plant on any scale. In addition the invention proves to be particularly useful for pea k shaving purposes in gasifica-tion based combined cycle power plant or in fuel proces-sors, e.g. by injecting a (liquid) methanol water mixture after the autothermal reformer.
Fig. 1 illustrates the process of the invention. Synthesis gas 1 is injected into a shift section 2. A stream of methanol 3 and water 4 are also injected into the shift section 2 where the shift step occurs. The methanol stream 3 can be added either in liquid form or in vapour form. The water 4 can be added as vapour. The shift section contains catalyst having activity both for the shift conversion re-action of the carbon monoxide and the steam reforming reac-tion of methanol. The heat required for the endothermic steam reforming reaction of methanol is provided by the heat obtained in the shift conversion reaction. The product is a hydrogen-rich stream 5.
The catalyst suitable for the process contains copper, zinc, aluminium and/or chromium. Using this catalyst re-sults in an increase in capacity and the catalyst is active at both lower temperatures and at temperatures above 350°C.
Addition of methanol and water in vapour form has the ad-vantage that complicated dispersive elements required to distribute liquid methanol in the shift section are avoided. An additional benefit is the high reactant partial pressure created throughout the shift section. Methanol can be added as a single stream, which is an advantage.
The shift section can comprise a single shift step or a combination of shift step s. An embodiment of the invention comprises a process, where at least one shift step is a me-dium-temperature or a high temperature shift step. Another embodiment of the invention comprises a process where the medium or high temperature shift step is followed by a low temperature shift step. Other combinations of shift steps are also possible and are encompassed by the process of the invention.

The synthesis gas stream 1 can be obtained from various sources for example a steam ref ormed gas, a secondary re-former, an auto thermal reforme r or a partial oxidation unit such as an oil or coal gasifier.
A particular embodiment of the invention comprises the pro-cess where a hydrocarbon stream and steam are first pre-reformed to obtain methane and then steam reformed to ob-tain a gas containing carbon monoxide before entering the shift step. After the shift reaction the hydrogen produced is separated and unconverted methanol is recycled to the pre-reformer.
Besides methanol, other similar species like methyl for-miate, formaldehyde or formic acid may be used.
The invention is also applicabl a in an ammonia or urea plant of any scale. Methanol ma y be used as fuel substitute or for boosting the capacity of the plant.
In the conventional ammonia plant, nitrogen is supplied as air to the secondary reformer i n a balanced amount so that the H2/N2 ratio is close to 3 before the gas enters the am-monia synthesis loop. Addition of methanol to the shift section in the loop increases t he amount of hydrogen pro-duced. The H2/N2 ratio can be maintained at 3 by increasing the amount of air added to the secondary reformer. This will require a decrease of the firing in the primary re-former.

Methanol is stoichiometric with respect to urea:
CH30H + H~0 -> 3H~+ C02 ( 2 ) 3H2 + N~ -> 2NH3 ( 3 ) 2NH3 + COz --> ( NH2 ) 2C0 + H20 ( 4 ) Synthesis gas arising from steam reforming of light natural gas has a deficit in C02. Addition of a large amount of methanol requires no firing in the primary reformer i.e.
firing becomes superfluous. Carbon dioxide produced during the process (reaction (2)) may be used in the ammonia plant for additional urea production (reactions (3 and 4)). In the process of the invention, urea is produced by reacting ammonia and carbon dioxi de according to reaction (4) using conventional methods. Thus, methanol can be used to in-crease the fuel flexibility of an ammonia plant and simul-taneously supply C0~ for urea production.
Partial oxidation based ammonia preparation based on addi-tion of hydrogen and carbon dioxide can be supplied in a similar manner.
The advantages of the process of the invention are illus-trated in the following examples.
EXAMPLES
The following catalysts from Haldor Topsoe A/S have been used in the examples:

Catalyst A: SK201-2 - a high- temperature shift catalyst comprising oxides of copper, iron and chro-mium.
Catalyst B: MK101 - methanol synthesis catalysts com-prising oxides of copper, zinc and aluminum.
Catalyst C: MK121 - methanol synthesis catalysts com-prising oxides of copper, zinc and aluminum.
The reactions all take place at pressures of 0-10 Mpa g, preferably at 2-6 Mpa g. The men tinned pressures are values above atmospheric pressure as indicated.
Example 1 is a comparative example, which serves to demon-strate that catalysts such as catalyst A are not suited for the production of hydrogen from methanol cracking. Examples 2-13 serve to demonstrate the scope of the present inven-tion using copper-based catalyst s. In these examples, it is demonstrated how hydrogen production, according to the pro-cess of the invention, may be improved significantly and with extremely high efficiency. Examples 14-18 are compara-tive examples demonstrating the performance of the cata-lysts under normal water gas shi ft conditions. Catalyst C
is used in these examples.
Example 1 (Comparative) g of catalyst A is activated by means of steam and a dry gas containing 15% C0, 10 o COz and 75 o H2. It is further tested at 380°C at a dry gas flow of 50 N1/h and a steam flow of 45 Nl/h at a pressure of 2.3 Mpa. After 70 hours the CO concentration in the dry exit gas is 3.70. Further addition of 0.5 N1/h of methanol causes the CO exit concen-tration to increase to 4.Oo and the exit CH4 concentration to increase from 20ppm to 1000ppm. Furthermore, the water condensed after the reacto r contained a significant amount of unconverted methanol corresponding to approximately 500 of the methanol added. Whe n the methanol was removed the CH4 formation decreased to 25ppm and the CO formation to 3.90.
The result clearly shows that this catalyst is unsuitable for catalytic methanol decomposition into hydrogen and car-bon oxides.
Example 2 15.2 g of catalyst B is reduced in diluted hydrogen (1-5 volo) at 185°C at a pressu re of 0.1 MPa and the synthesis gas being comprised of 43.10 hydrogen, 14.30 carbon monox-ide, 11.10 carbon dioxide and 31.50 nitrogen is introduced.
The pressure is increased to 2.5 MPa and the temperature is raised to 235°C. A solution of 19.630 wt/wt methanol in wa-ter is evaporated and co-fed with the synthesis gas. The dry gas flow is 100 N1/h, whereas the liquid flow is 41.6 g/h corresponding to a ste am flow of 41.6 Nl/h and a metha-nol flow of 5.7 N1/h. The exit gas is analysed after con-densation of residual steam and methanol. At these condi-tions the CO exit concentration amounts to 0.900 and the CO~ exit concentration is 21.70 and the dry flow gas flow is increased to 130 Nl/h. No CH9 is observed at any time, the detection limit being approximately 1 ppm.
At these conditions, the a xit temperature is measured to be 242°C immediately after th a catalyst bed and the liquid flow exit in the reactor is 20.8 g/h with a methanol con-centration of 8.140 wt/wt. The methanol exit flow is thus 1.18 Nl/h. This corresponds to a methanol conversion C(M):
C (M) _ ( (methanol flowi"let - methanol floweXit) /methanol flow inlet)*100a - 79.30.
The carbon monoxide conversion is calculated as C(CO):
C (CO) _ ( (CO flowinlet - CO flowexit) /CO flow inlet) *100 0 =
91.80.
The productivity of hydrogen is calculated as Prod(H2):
Prod(H2) - (hydrogen flowexit - hydrogen flowinlet) /mass of catalyst = 1700 N1 H2/kg/h.
Carbon mass balance, C(in)/C(ex), is found to be 1.02. The results are summarised in Table 1.
Examples 3-7 As Example 2 except for variations in temperature, dry gas flow and liquid flow as according to Table 1. The catalyst is the same batch as used in Example 2. Analysis of the condensable part of the exit gas of Example 7 reveals a concentration of ethanol of 10 ppm wt/wt. No higher alco-hols, methane or any other hydrocarbons are observed in any of Examples 3-7. The selectivity of methanol conversion to carbon oxides and hydrogen is thus 1000 within the accuracy of the experiments.
Example 8 15.1 g of catalyst C is reduced in dry diluted hydrogen (1-volo) at 185°C at a pressure of 0.1 Mpa and the synthesis gas being comprised of 43.10 hydrogen, 14.30 carbon monox-ide, 11.10 carbon dioxide and 31.50 nitrogen is introduced.
The pressure is increased to 2.5 MPa and the temperature is raised to 216°C. A sot ution of 22.370 wt/wt methanol in wa-ter is evaporated and co-fed with the synthesis gas. The dry gas flow is 50 Nl / h, whereas the liquid flow is 16.0 g/h corresponding to a steam flow of 15.5 Nl/h and a metha-nol flow of 2.5 N1/h. The exit gas is analysed after con-densation of residual steam and methanol. At these condi-tions the CO exit concentration amounts to 0.640 and the C02 exit concentration is 22.30 and the dry flow gas flow is increased to 63 Nl / h. No CHQ is observed at any time, the detection limit being approximately 1 ppm. At these conditions, the exit temperature is measured to be 219°C
immediately after the catalyst bed and the liquid flow exit the reactor is 18.7 g/ h with a methanol concentration of 11.26 % wt/wt. The met hanol exit flow is thus 1.47 N1/h.
The conversions are calculated as above with C(M) - 56.90 and C(CO) - 94.30. The productivity of hydrogen is Prod(H2) - 749 Nl H2/g/h. Carbon mass balance is found to be 1.00.
The results of methano 1-boosted shift over catalyst C are summarised in Table 2.

Table 1 Example 2 3 4 5 6 7 inlet Temp (C) 235 235 273 273 311 3l2 exit Temp (C) 242 237 275 275 312 309 Inlet dry flow 100 50 100 50 100 100 (Nl/h) inlet liqui d flow41.6 18.8 41.7 17.8 41.5 60.0 (g/h) inlet steam flow 42 19 42 18 42 60 (N1/h) inlet MeOH flow 5.7 2.6 5.7 2.4 5.7 8.2 (N1/h) exit dry flow 130 66 137 67 137 l48 (Nl/h) exit liquid flow 20.8 7.9 19.5 9.4 17.0 27.6 (g/h) [MeOH] exit 8 . 8 . 3 . 2 . 1 . 1 ( a wt/wt) 14 26 58 03 03 .

[CO]exit (mole 0.90 0.66 1.20 1.30 1.79 1.20 o) C(M) (%) 79.3 82.3 91.5 94.6 97.8 97.0 C(CO) (%) 91.8 93.8 88.4 87.7 82.7 87.5 Prod(HZ) (N1 /kg/h)1700 940 2080 970 2090 2640 ~ din) /f hex) 1 . 0 . 0 . 0 . 0 . 0 02 99 98 98 98 .

Example 9 This experiment is similar to Example 8 except for varia-tion in dry gas flow and liquid flow as shown in Table 2.
The selectivity of methanol conversion to carbon oxides and hydrogen is 1000.

Example 10 The catalyst a sed in Examples 8-9 is left on stream for 120 hours at an in 1 et temperature of 313°C, a dry gas flow of 100 Nl/h, a liquid flow of 60 g/h, a pressure of 2.5 MPa and with feed compositions as in Examples 8-9. The selec-tivity of meth anol conversion to carbon oxides and hydrogen is 1000. The exit concentration of carbon monoxide is con-stant at 1.25~0.050 in this period. After the 120 hours pe-riod the condensate was analysed again with the results given in Table 2.
Examples 11-13 These experiments are similar to Example 10 except for variations in temperature, dry gas flow and liquid flow as shown in Table 2.
Examples 14-17 (Comparative) These experiments are similar to Examples 10-13 except that methanol is excluded from the liquid feed. The results catalyst C without methanol addition are shown in Table 3.

Table 2 Example No. 8 9 10 11 l2 13 Inlet Temp. (C) 216 216 313 3l3 275 236 Exit Temp. (C) 219 224 310 314 279 244 Inlet dry flow 50 100 100 100 100 l00 (N1/h) Inlet liquid flow18.7 60 60 41.9 39.8 41.7 (g/h) Tnlet steam flow 18 58 58 40 38 40 (N1/h) Inlet MeOH flow 2.9 9.4 9.4 6.6 6.2 6.5 (N1/h) Exit dry flow 63 l31 148 139 139 134 (N1/h) Exit liquid flow 16.0 39.6 31.9 20.3 19.3 21,4 (g/h) [MeOH]e,~it(%wt/w)11.26 14.77 1.52 1.29 3.45 10.87 [CO]exit (moleo) 0.64 0.95 1.23 1.86 1.34 1.11 C(M)(o) 56.9 56.4 96.4 97.2 92.5 75.1 C(CO)(%) 94.3 91.2 87.2 81.8 86.9 89.5 Prod(H2) (N1/kglh)750 1700 2550 2140 2180 1920 C(in)/C(ex> 1.00 1.03 1.04 1.02 1.01 1.03 Table 3 Example No. l4 15 16 l7 Inlet Temp. (C) 236 274 312 313 Exit Temp. (C) 253 289 325 327 Inlet dry flow (N1/h) 100 100 100 100 Inlet liquid flow (g/h) 31.8 31.8 31.8 46.2 Inlet steam flow (N1/h) 40 40 40 57 Inlet MeOH flow (Nl/h) 0 0 0 0 Exit dry flow (N1/h) 116 116 115 116 Exit liquid flow (Nl/h) - - - -[MeOH] exit (% wt/wt) - - - -[CO] exit (mole %) 0.88 1.13 1.62 1.15 C(M) (%) _ _ _ _ C(CO) (o) 92.9 90.8 87.0 90.8 Prod (H2) (N1/kg/h) 1060 1040 1000 1040 C(in)/C(ex) 1.03 1.03 1.03 1.03 The above examples demonstrate that hydrogen production may be significantly improved by addition of methanol to a syn-thesis gas and exposing the resulting mixture to a catalyst containing copper. Thus, when 15 g of the catalyst MK121 is exposed to synthesis gas at an inlet temperature of 313°C' at a dry gas flow of 100 Nl/h, a steam flow of 57 Nl/h and 25 bar pressure, the hydrogen production amounts to 1040 Nl/kg/h(Example 17). In this example the exit temperature is 327°C and the CO concentration is 1.150. With the same catalyst, addition of 9.4 Nl/h methanol to the feed but otherwise the same conditions of operation, the hydrogen productivity increases to 2550 N1/kg/h (Example 10). In this example the exit temperature is 310°C and the CO con-centration is 1.230.
Example 18 This example describes the benefit of adding methanol to a natural gas based ammonia plant for increasing the urea production.
In many situations the balance between hydrogen and carbon dioxide does not fully make the requirement for urea pro-duction due to a shortage in carbon dioxide. The process of the invention can be used for new grassroots plants as well as for exiting plants.
This example is illustrated by the process shown in Fig. 2.
Methanol from the storage tank 1 is pumped to the methanol preheater 2, where the methanol is evaporated. Methanol is mixed with the gas stream 3 from the secondary reformer (after cooling) and sent to the shift reactor 4. In reactor 4, which is loaded with a catalyst containing copper, zinc, aluminium and/or chromium, the water gas shift reaction (reaction 1) as well as methanol decomposition (reaction 2) take place.
The exit gas from shift reactor 4 contains more carbon di-oxide than the exit gas from a conventional shift reactor process. Table 4 shows the concentrations of the various components present in the gas stream at three different po-sitions indicated in Fig. 2.
Table 4 Pos. 1 2 3 Comp. Nm3/h Mole Nm /h Mole % Nm /h Mole o o HZ 103323 54.19 103323 53.38 133229 59.62 NZ 47596 24.97 47596 24.59 47596 21.30 Co 26024 13.65 26024 13.44 4743 2.12 Coy 12595 6.61 12595 6.51 36751 16.44 Ar 575 0.30 575 0.30 574 0.26 CHQ 541 0.28 541 0.28 541 0.24 MeoH - - 2913 1.51 39 0.02 Total 190654 - 193567 - 223473 -Dry Total 279125 - 282038 - 287788 -Table 5 shoran the production figures achieved by adding 100 MTPD methanol upstream of the shift reactor in a 1500 MTPD
ammonia plant used for urea production. The amount of ammo-nia produced is reduced due to the formation of urea. As can be seen the urea production is increased by 191 MTPD by adding 100 MTPD methanol.

Table 5 Component Conventional ProcessMeoH addition Feed Gas (Nm /h) 38260 38260 MeoH (MTPD) - 100 Ammonia Prod. (MTPD)161 151 Urea Prod. (MTPD) 2366 2557

Claims

1. A process for the preparation of urea comprising adding methanol to a carbon monoxide containing gas, con-tacting the carbon monoxide-containing gas, methanol and water in at least one shift step in the presence of a cata-lyst comprising copper, zinc and aluminium and/or chromium to produce a stream rich in carbon dioxide, transferring the carbon dioxide-rich stream from the at least one shift step to a urea reactor and reacting the carbon dioxide with.
ammonia to produce urea.

2. A process according to claim 1, wherein methanol and water are in vapour form.

3. A process according to claim 1, wherein methanol and water are in liquid form.

4. A process according to claim 1 and 2 or 3, wherein the at least one shift step is a Medium Temperature or High Temperature shift step.

5. A process according to claim 4, wherein the Medium Temperature or High Temperature shift step is followed by a Low Temperature shift step.

6. Process according to claims 1, 2, 4 and 5, wherein the carbon monoxide-containing gas is obtained from reform-ing and/or partial oxidation of a hydrocarbon feed.

7. Process according to claim 6, wherein the hydrocar-bon feed is pre-reformed before the reforming step.

8. Process according to claim 7, wherein unreacted methanol is separated from the shift step effluent and re-cycled to the pre-reforming step.

9. Process according to claim 4, wherein the shift in-let temperature is at least 280°C and the pressure is 0-10 Mpa g, preferably 2-6 Mpa g.