ATR based ammonia process and plant

Abstract

A process for producing an ammonia synthesis gas, said process including the steps of: reforming a hydrocarbon feed in a reforming step thereby obtaining a synthesis gas comprising CH.sub.4, CO, CO.sub.2, H.sub.2 and H.sub.2O; and shifting said synthesis gas in a high temperature shift step over a promoted zinc-aluminum oxide based high temperature shift catalyst, wherein the steam/carbon ratio in the reforming step is less than 2.6.

Claims

1. A process for producing an ammonia synthesis gas, said process comprising the steps of: reforming a hydrocarbon feed mixed with oxygen in an autothermal reformer (ATR) thereby obtaining a synthesis gas comprising CH.sub.4, CO, CO.sub.2, H.sub.2 and steam; and shifting said synthesis gas in at least one high temperature shift step over a promoted zinc-aluminum oxide based high temperature shift catalyst, wherein steam is present in the reforming step, wherein a steam/carbon ratio in the reforming step is 1.4-0.3, and wherein at least a required amount of nitrogen for a downstream ammonia synthesis is added into the synthesis gas in an N.sub.2 wash step.

2. The process according to claim 1, wherein the temperature in the high temperature shift step is 300-600 C.

3. The process according to claim 1, wherein the promoted zinc-aluminum oxide based HT shift catalyst comprises in its active form a Zn/Al molar ratio in the range 0.5 to 1.0 and a content of alkali metal in the range 0.4 to 8.0 wt % and a copper content in the range 0-10% based on the weight of oxidized catalyst.

4. The process according to claim 1, wherein a space velocity in the ATR is less than 20,000 Nm.sup.3 C/m.sup.3/h.

5. The process according to claim 1, further comprising a prereforming step.

6. The process according to claim 1, wherein the at least one high temperature shift step is two or more high temperature shift steps in series.

7. The process according to claim 1, further comprising one or more additional shift steps downstream of the high temperature shift step.

8. The process according to claim 7, wherein the one or more additional shift steps are one or more medium temperature shift steps and/or one or more low temperature shift steps.

9. The process according to claim 7, wherein steam is added to the synthesis gas before the one or more additional shift steps downstream of the high temperature shift step.

10. The process according to claim 7, wherein the synthesis gas leaving the one or more additional shift steps downstream of the high temperature shift step is washed with water to reduce a content of methanol formed as a by-product during shifting the synthesis gas.

11. The process according to claim 1, further comprising a CO.sub.2 removal step removing CO.sub.2 from the synthesis gas down to a level less than 400 vppm CO.sub.2.

12. The process according to claim 1, wherein there is no additional steam addition between the reforming step and the at least one high temperature shift step.

13. A process for producing ammonia, wherein an ammonia synthesis gas is achieved by the process according to claim 1.

14. A process for producing an ammonia synthesis gas, said process comprising the steps of: reforming a hydrocarbon feed mixed with oxygen in an autothermal reformer (ATR) thereby obtaining a synthesis gas comprising CH.sub.4, CO, CO.sub.2, H.sub.2 and steam; shifting said synthesis gas in a high temperature shift step over a promoted zinc-aluminum oxide based high temperature shift catalyst; and conducting an N.sub.2 wash of said synthesis gas, adding nitrogen to the washed synthesis gas such that, after addition of nitrogen, the synthesis gas contains hydrogen and nitrogen in a ratio H.sub.2/N.sub.2 of about 3; wherein steam is present in the reforming step; and wherein a steam/carbon ratio in the reforming step is 1.4-0.3.

15. The process according to claim 14, wherein there is no additional steam addition between the reforming step and the high temperature shift step.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

(1) FIG. 1 shows an exemplary process scheme.

(2) FIG. 2 shows an exemplary process scheme.

DETAILED DESCRIPTION

(3) In a first aspect of the present invention is provided a process enabling a process scheme utilizing proven reforming technology operating with low steam/carbon.

(4) In a second aspect of the present invention is provided a process enabling operation of the high temperature (HT) shift downstream the reforming section at the same low steam/carbon ratio as the reforming section.

(5) In a third aspect of the present invention is provided a process scheme without the necessity of a methanation section to remove residual carbon components in the make-up synthesis gas for an ammonia synthesis.

(6) In a fourth aspect of the present invention is provided an overall process layout enabling maximum single line capacity.

(7) These and further advantages are achieved by a process for producing an ammonia synthesis gas, said process comprising the steps of: Reforming a hydrocarbon feed in a reforming step thereby obtaining a synthesis gas comprising CH.sub.4, CO, CO.sub.2, H.sub.2 and H.sub.2O Shifting said synthesis gas in a high temperature shift step over a promoted zinc-aluminum oxide based HT shift catalyst, wherein The steam/carbon in the reforming step is less than 2.6.

(8) HT shift is defined as a process step where a synthesis gas containing CO, CO.sub.2, H.sub.2 and H.sub.2O undergoes the shift reaction in the temperature range from 300 C. to 600 C.

(9) In a conventional ammonia plant the standard use of an iron based HT shift catalyst requires a steam/carbon ratio of around 3.0 to avoid iron carbide formation.
5Fe.sub.3O.sub.4+32CO.Math.3Fe.sub.5C.sub.2+26CO.sub.2(1)

(10) Formation of iron carbide will weaken the catalyst pellets and may result in catalyst disintegration and pressure drop increase.

(11) Iron carbide will catalyse Fischer-Tropsch by-product formation
nCO+(n+m/2)H.sub.2.Math.C.sub.nH.sub.m+nH.sub.2O(2)

(12) The Fischer-Tropsch reactions consume hydrogen, whereby the efficiency of the shift section is reduced.

(13) However, according to the present invention a non Fe-catalyst is used, such as a promoted zinc-aluminum oxide based catalyst. For example, the Topse SK-501 Flex HT shift catalyst which enables operation of the reforming section and HT shift section at a steam/carbon ratio down to 0.3.

(14) Thus, the present process operating at a steam/carbon ratio down to 0.3 is in contrast to today's traditional ammonia plants which are based on reforming and/or HT shift sections operating at a steam/carbon ratio of 2.6 or higher. In advantageous embodiments of the process the zinc-aluminum oxide based catalyst in its active form comprises a mixture of zinc aluminum spinel and zinc oxide in combination with an alkali metal selected from the group consisting of Na, K, Rb, Cs and mixtures thereof, and optionally in combination with Cu. The catalyst may have a Zn/Al molar ratio in the range 0.5 to 1.0, a content of alkali metal in the range 0.4 to 8.0 wt % and a copper content in the range 0-10% based on the weight of oxidized catalyst.

(15) The HT shift catalyst used according to the present process is not limited by strict requirements to steam to carbon ratios, which makes it possible to reduce the steam/carbon ratio in the shift section as well as the reforming section.

(16) A steam/carbon ratio of less than 2.6 has several advantages. Reducing the steam/carbon ratio on a general basis leads to reduced feed plus steam flow through the reforming section and the downstream cooling and synthesis gas preparation sections.

(17) A low steam/carbon ratio in the reforming section and the shift section enables higher syngas throughput compared to a high steam/carbon ratio. Nitrogen added via the Nitrogen wash enables higher syngas throughput compared to adding nitrogen in the reforming section. The absence of a methanation section reduces pressure loss and using an inert free gas in the ammonia synthesis section enables higher throughput in the ammonia synthesis section

(18) Reduced mass flow through these sections means smaller equipment and piping sizes. The reduced mass flow also results in reduced production of low temperature calories, which can often not be utilised. This means that there is a potential for both lower CAPEX and OPEX.

(19) The present process may further comprise one or more of the following steps: Shifting the HT shift outlet gas in one or more medium temperature (MT)/low temperature (LT) shift step(s). The MT/LT shift step(s) can optionally be performed at a higher steam/carbon ratio than the HT shift to limit byproduct formation such as methanol. Optionally remove methanol from the MT/LT shift outlet gas in a water wash Remove CO.sub.2 from the MT/LT shift outlet gas/water wash outlet gas down to a level lower than 500 ppm preferably down to below 20 ppm. Remove residual CO.sub.2 and H.sub.2O from the gas leaving the CO.sub.2 removal section in a molecular sieve dryer section. Remove CH.sub.4, CO and inerts such as Ar and He from the gas leaving the molecular sieve dryer section in a nitrogen wash section and adjust the N.sub.2/H.sub.2 ratio to approximate 3 as needed for the ammonia synthesis. Convert the adjusted outlet gas from the nitrogen wash to ammonia in an inert free ammonia synthesis section.

(20) In preferred embodiments the reforming step comprises at least an autothermal reformer (ATR).

(21) As the requirements to the steam/carbon ratio in the HT shift step by the present process is significantly reduced compared to known technologies it is possible by the present invention to reduce the steam/carbon ratio through the front end to, e.g., 0.6 or as low as possible dependent on the possible shift solutions. An advantage of a low steam/carbon ratio to the ATR and in the overall process is that smaller equipment is required in the front-end due to the lower total mass flow through the plant.

(22) The carbon feed for the ATR is mixed with oxygen and additional steam in the ATR, and a combination of at least two types of reactions takes place. These two reactions are combustion and steam reforming.

(23) Combustion Zone:
2H.sub.2+O.sub.2.Math.2H.sub.2O+heat(3)
CH.sub.4+3/2O.sub.2.Math.CO+2H.sub.2O+heat(4)
Thermal and Catalytic Zone:
CH.sub.4+H.sub.2O+heat.Math.CO+3H.sub.2(5)
CO+H.sub.2O.Math.CO.sub.2+H.sub.2+heat(6)

(24) The combustion of methane to carbon monoxide and water (4) is a highly exothermic process. Excess methane may be present at the combustion zone exit after all oxygen has been converted.

(25) The thermal zone is part of the combustion chamber where further conversion of the hydrocarbons proceed by homogenous gas phase reactions, mostly (5) and (6). The endothermic steam reforming of methane (5) consumes a large part of the heat developed in the combustion zone.

(26) Following the combustion chamber there may be a fixed catalyst bed, the catalytic zone, in which the final hydrocarbon conversion takes place through heterogeneous catalytic reactions. At the exit of the catalytic zone, the synthesis gas preferably is close to equilibrium with respect to reactions (5) and (6).

(27) The steam/carbon ratio in the reforming section may be 2.6-0.1, 2.4-0.1, 2-0.2, 1.5-0.3, 1.4-0.4, such as 1.2, 1.0 or 0.6.

(28) The steam/carbon ratio is defined as the ratio of all steam added to the reforming section upstream of the HT shift section (i.e. steam which may have been added to the reforming section via the feedgas, oxygen feed, by addition to burners etc.) and the hydrocarbons in the feedgas to the reforming section on a molar basis.

(29) Thus, according to the present invention it is possible to run the process with no additional steam addition between the reforming step(s) and the high temperature shift step.

(30) In advantageous embodiments, the space velocity in the ATR is low, such as less than 20,000 Nm.sup.3 C/m.sup.3/h, preferably less than 12,000 Nm.sup.3 C/m.sup.3/h and most preferably less 7000 Nm.sup.3 C/m.sup.3/h. The space velocity can be defined as the volumetric carbon flow per catalyst volume and is thus independent of the conversion in the catalyst zone.

(31) In preferred embodiments, the temperature in the HT shift step is in the range 300-600 C., such as 360-470 C. This means that, according to the present process, it is possible to run a high temperature shift reaction on a feed with a much lower steam/carbon ratio than is possible by known processes. For example, the high temperature shift inlet temperature may be 300-400 C., such as 350-380 C.

(32) Preferably, a prereformer is provided as part of the reforming section upstream for example an ATR. In the prereformer all higher hydrocarbons can be converted to carbon oxides and methane, but also for light hydrocarbons the prereformer is advantageous. Providing the prereformer may have several advantages including reducing the required O.sub.2 consumption in the ATR and allowing higher inlet temperatures to the ATR since cracking risk by preheating is minimized. Thereby the ignition conditions are reached. Furthermore, the prereformer may provide an efficient sulphur guard resulting in a practically sulphur free feed gas entering the ATR and the downstream system. The prereforming step may be carried out at temperatures between 300-650 C., preferably 390-480 C.

(33) In various embodiments a fired heater is used for preheating the natural gas feed, the prereformer and ATR feed and for steam superheating. The necessary heat may be generated by burning a mixture of natural gas, waste gas (from the N.sub.2 wash), off gas (from the inert vent gas separator) and flash gas (from the CO.sub.2 removal section).

(34) The low steam/carbon ratio may result in a lower than optimal shift conversion which means that it in some embodiments may be advantageous to provide one or more additional shift steps. The one or more additional shift steps may include a MT shift and/or a LT shift and/or a HT shift. Generally speaking, the more converted CO in the shift steps the more gained H.sub.2 and the smaller front end required.

(35) This is also seen from the exothermic shift reaction given below
CO+H.sub.2O.Math.CO.sub.2+H.sub.2+heat(7)

(36) Steam may optionally be added after the HT shift step such as before one or more following MT or LT shift and/or HT shift steps in order to maximize performance of said following HT, MT and/or LT shift steps.

(37) Having two or more HT shift steps in series (such as a HT shift step comprising two or more shift reactors in series e.g. with the possibility for cooling and/or steam addition in between) may be advantageous as it may provide increased shift conversion at high temperature which gives a possible reduction in required shift catalyst volume and therefore a possible reduction in capex. Furthermore, high temperature reduces the formation of methanol, a typical shift step byproduct.

(38) Preferably, the MT and LT shift steps may be carried out over promoted copper/zink/aluminia catalysts. For example the low temperature shift catalyst type may be LK-821-2, which is characterized by high activity, high strength, and high tolerance towards sulphur poisoning. A top layer of a special catalyst may be installed to catch possible chlorine in the gas and to prevent liquid droplets from reaching the shift catalyst.

(39) The MT shift step may be carried out at temperatures at 190-360 C.

(40) The LT shift step may be carried out at temperatures at T.sub.dew+15-290 C., such as, 200-280 C. For example the low temperature shift inlet temperature is from T.sub.dew+15-250 C., such as 190-210 C.

(41) Reducing the steam/carbon ratio leads to reduced dew point of the process gas, which means that the inlet temperature to the MT and/or LT shift steps can be lowered. A lower inlet temperature can mean lower CO slippage from the outlet of the shift reactors.

(42) It is well known that MT/LT shift catalysts are prone to produce methanol as a byproduct. Such byproduct formation can be reduced by increasing the steam/carbon ratio. The CO.sub.2 wash following the MT/LT shifts requires heat for regeneration of the CO.sub.2 absorption solution. This heat is normally provided as sensible heat from the process gas but this is not always enough. Typically an addition steam fired reboiler provides the missing heat. Optionally adding steam to the process gas can replace this additional steam fired reboiler and simultaneously ensures reduction of byproduct formation in the MT/LT shifts section.

(43) The methanol formed by the MT/LT shift catalyst can optionally be removed from the synthesis gas in a water wash to be placed upstream of the CO.sub.2 removal step or on the CO.sub.2 product stream.

(44) In many advantageous embodiments a CO.sub.2 removal step may be carried out after/downstream the one or more shift steps. In standard design the CO.sub.2 content is 500 vppm in the treated gas.

(45) In preferred embodiments a CO.sub.2 removal step may be used to bring the CO.sub.2 content down to less than 400 vppm CO.sub.2, such as below 100 vppm or in some preferred embodiments down to 20 vppm or below.

(46) The process may further comprise a washing step, preferably a N.sub.2 wash. The N.sub.2 wash may serve several purposes, such as purification of the syngas, as well as to add the stoichiometric required nitrogen for a downstream ammonia synthesis.

(47) The nitrogen for the N.sub.2 wash unit (NWU) may be supplied by an air separation unit (ASU) which separates atmospheric air into its primary components nitrogen and oxygen. The oxygen is used in the ATR and the nitrogen in the NWU.

(48) After the one or more shift sections and CO.sub.2 removal unit, the gas may contain residual CO and CO.sub.2 together with small amounts of CH.sub.4, Ar, He and H.sub.2O.

(49) CO.sub.2 and H.sub.2O are preferably removed before the N.sub.2 wash because they otherwise would freeze at the low operating temperature of the N.sub.2 wash. This may, for example, be done by adsorption in a molecular sieve dryer consisting of at least two vessels one in operation while the other is being regenerated. Nitrogen may be used as dry gas for regeneration.

(50) In the NWU the syngas is washed by liquid nitrogen in a column where CH.sub.4, Ar, He and CO are removed. The purified syngas preferably contains only ppm levels of Ar and CH.sub.4.

(51) The waste gas containing the impurities together with some lost nitrogen may advantageously be used as fuel in the fired heater.

(52) After the NWU, nitrogen gas may be added to the process stream in order to adjust the N.sub.2 content to a preferred ratio H.sub.2/N.sub.2 ratio of 3 in the make-up stream to the ammonia synthesis loop.

(53) Because the purified syngas now only contains H.sub.2 and N.sub.2 in the correct stoichiometric ratio for ammonia synthesis, plus ppm levels of Ar and CH.sub.4, the ammonia synthesis section can be considered inert free.

(54) An ammonia synthesis loop is defined as inert free when it is not required to purge gas from the loop because the build-up of inerts is negligible without such purge.

Example

(55) The below positions refer to the inlet of the units.

(56) TABLE-US-00001 TABLE 1 Flows and operating conditions Position Pos. 1 Pos. 2 Pos. 3 Pos. 4 Pos. 5 Pos. 6 Pos. 7/8 Pos. 9 Temperature, 450 650 360 205 75 20 15 32 C. Pressure, 41.3 39.3 35.7 34.8 33.3 32.4 31.4/186 5.0 kg/cm.sup.2 g Flow, 187,432 198,532 391,946 391,946 391,946 270,351 334,226 164,416 Nm.sup.3/h
Main Components

(57) TABLE-US-00002 TABLE 2 Stream compositions, Positions refer to FIG. 1 Composition, Mole % Pos. 1 Pos. 2 Pos. 3 Pos. 4 Pos. 5 Pos. 6 Pos. 7/8 Pos. 9 Ar 0.07 0.07 0.07 0.10 C.sub.3+ 2.23 C.sub.2H.sub.6 2.58 CH.sub.4 40.02 47.47 0.95 0.95 0.95 1.37 CO 0.38 0.09 20.08 10.07 3.49 5.04 CO.sub.2 2.58 5.50 5.85 15.86 22.44 20 ppm H.sub.2 1.49 4.93 47.81 57.82 64.40 92.96 75.00 N.sub.2 0.66 0.63 0.32 0.32 0.32 0.46 25.00 NH.sub.3 100.00 H.sub.2O 50.04 41.38 24.93 14.92 8.34 0.07

(58) Prereformer: Tin/Tout: 450/449 C. (T=1 C.)

(59) Steam/carbon ratio, S/C=0.9 inlet of the prereformer

(60) ATR:

(61) The process gas enters the ATR at 650 C. and the temperature of the oxygen is around 260 C.

(62) Steam/carbon ratio, S/C=1.0 as per definition in the description

(63) The process gas leaves the reforming section at about 1025 C. through a refractory lined outlet section and transfer line to the waste heat boilers in the process gas cooling section.

(64) Shift Section:

(65) HT: Tin/Tout: 360/469 C. (T=109 C.)

(66) LT: Tin/Tout: 205/280 C. (T=75 C.)

(67) After reforming, about 26.7 vol % CO is present in the gas (dry basis). In the high temperature CO converter the CO content is reduced to approximately 11.8 vol %, and the temperature increases from 360 C. to 469 C. The heat content of the effluent from the high temperature CO converter is recovered in a waste heat boiler and in a boiler feed water preheater.

(68) The process gas is thereby cooled to 205 C. and passed on to the low temperature CO converter in which the CO content is reduced to approximately 3.8 vol %, while the temperature increases to 280 C.

(69) CO.sub.2 Removal Section

(70) The CO.sub.2 content in the outlet stream from shift section is reduced to 20 ppm. All Methanol in the synthesis gas going to the CO.sub.2 removal section will leave this section with the process condensate and the CO.sub.2 product stream. A water wash on the synthesis gas going to the CO.sub.2 removal section, see FIG. 2, or on the CO.sub.2 product stream can minimize the Methanol content in the CO.sub.2 product stream

(71) N.sub.2 Wash Section

(72) The first step in this section is a quantitative removal of CO.sub.2 and H.sub.2O in a molecular sieve dryer. The next step is a N.sub.2 liquid wash removing components other than H.sub.2 and N.sub.2 down to ppm level. The third step is to adjust the H.sub.2/N.sub.2 ratio to approximate 3 using gaseous nitrogen.

(73) Syngas Compressor:

(74) The synthesis gas is compressed from 31.4 to 185.5 kg/cm.sup.2g in the centrifugal type two-casing synthesis gas compressor. Part of the last casing forms the recirculation compressor in the synthesis loop.

(75) Inert free loop: The loop can be defined as inert when no purge gas system is required.

(76) The small amounts of inert gases entering the loop with the make-up synthesis gas will accumulate in the loop until the amount of inert gases dissolved in the liquid ammonia exit the let-down vessel equals the amount entering the loop. Off gas from the let-down vessel is recycled back to the synthesis gas compressor.

(77) The recycled inert level is dependent on the level of inerts dissolved in the liquid ammonia leaving the ammonia separator and the let-down vessel.

(78) If required, the level of inert gas in the loop can be reduced by an intermittent purge of a small gas stream.

(79) In this example the inert level in the purified gas leaving the N.sub.2 wash is 17 ppm Ar, in the make-up gas 53 ppm Ar (after addition of the off gas recycle stream from the let-down vessel) and 0.30% Ar inlet the converter.