METHOD FOR REMOVING HYDROGEN SULPHIDE FROM GASEOUS MIXTURES CONTAINING HYDROGEN SULPHIDE AND CARBON DIOXIDE

Abstract: A method of removing hydrogen sulphide from a gase­ ous mixture (111) comprising hydrogen sulphide and carbon dioxide, which method comprises the steps of bringing said gaseous mixture into contact with a liquid absorbent in an absorber (112) heating (212, 214) the liquid leaving the absorber to desorb part of the hydrogen sulphide and carbon dioxide dissolved therein, separating said liquid from said desorbed carbon dioxide and hydrogen sulphide in a separator (130) recycling at least part of said desorbed car­ bon dioxide and hydrogen sulphide to said absorber, and expanding (120) said liquid downstream of said separator.

Description: This invention relates to a method for removing hydrogen sulphide from gaseous mixtures comprising hydrogen sulphide and carbon dioxide. Figure 1 of the accompanying drawings shows a known installation for recovering a hydrogen rich stream from synthesis gas. In particular, 4006.4 lb moles/hr of synthesis gas comprising: 1756.5 lb moles hydrogen 22.2 n methane 6.4 " water 249.3 " carbon dioxide 1921.3 " carbon monoxide 41.9 n hydrogen sulphide 7.0 " nitrogen 1.8 n sulphidesulphide at 660 psia and 1220F is heated to 3100F in heat exchanger 2. 5565.4 lb moles/hr of steam at 770 psia and 5140F are added to the synthesis gas at junction 3 and the resulting mixture is subjected to high and low temperature shift reactions in unit 4. The carbonyl sulphide is substantially hydrolysed in unit 5 and the gas leaving unit 5 is cooled to l400F in heat exchanger 6. Substantially all the water which condenses in heat exchanger 6 is separated from the gas in separator 7 and the remaining gas is cooled to 55 0F in heat exchangers 8 and 9. The water which condenses in heat exchangers 8 and 9 is separated from the gas in phase separator 10 and the residual gas at 637 psia and comprising (per hour) 3662.1 lb moles hydrogen 22.2 n methane 1.7 n water 2156.6 N carbon dioxide 15.7 " carbon monoxide 43.7 N hydrogen sulphide 7.0 " nitrogen 1.6 ppm carbonyl sulphide enters adsorber 12 through pipe 11. Substantially all the hydrogen sulphide in the residual g-s is absorbed in absorber 12 and the gas leaving adsorber 1 through pipe 13 contains approximately 0.5. ppm hydrogen sulphide. The gas in pipe 13 is joined by gas from conduit 14 .mprising (per hour) 23.5 lb moles hydrogen 0.2 " methane 0.01. " water 43.8 " carbon dioxide 0.2 " carbon monoxide 0.05 " nitrogen The resulting gas mixture is cooled to 34 F in heat exchanger and is then introduced into the bottom of a second absorber 16 where a substantial portion of the remaining carbon dioxide is removed. The gas stream leaves the second adsorber 16 through conduit 17 at 26 F and passes through heat exchanger 9 which it leaves through conduit 19 at 840F. The gas in conduit 19 Is at 634 psia and contains (per hour) 3651.6 lb moles hydrogen 21.5 N methane 0.3 N water 2.1 N carbon dioxide 15.6 " carbon monoxide 0.1 ppm hydrogen sulphide 6.9 lb moles nitrogen 0.3 ppm carbonyl sulphide The absorbent used in the process described is a mixture of ethers sold under the Trade Name SELEXOL. 200,000 lbs/hr SELEXOL together with 454.3 lb moles water enter second -edsorber 16 through conduit 20 at 650 psia and at 240F Simultaneouslyr 634,000 lbs/hr SELEXOL and 1438.3 lb moles waFer enter second absorber 16 through conduit 21 at 640 psia and 24 0F. As the SELEXOL descends through the second absorber 16 it absorbs carbon dioxide and, in so doing, increases in temperature. The absorbent leaves the bottom of the second absorber 16 through conduit 22 at 340F. The absorbent is then divided into two streams 23 and 24. Stream 24 is cooled to 250F in refrigerator 25 and part is passed through conduit 26 whilst the remainder is passed through conduit 27. Streams 23 and 26 are combined and the remaining stream at 32 F enters absorber 12 at 634 psia. As the stream descends through absorber 12 it absorbs hydrogen sulphide and leaves absorber 12 through conduit 28 at 50 0F and 637 psia. The stream is expanded to 28 psia across valve 29 and in so doing cools to 90F. The expansion releases a substantial quantity of gas from the stream and this is separated from the liquid in separator 30. The gas leaving separator 30 through conduit 31 comprises (per hour) 11.5 lb moles hydrogen 0.3 " methane 573.5 n carbon dioxide 0.1 N carbon monoxide 19.3 n hydrogen sulphide 0.4 ppm carbonyl sulphide The gas in conduit 31 is compressed to 640 psia in compressor 32, and after being cooled to 1040F in heat exchanger 33 is introduced into absorber 12. The liquid leaving separator 30 comprises (per hour): 455.6 lb moles water 131.9 II carbon dioxide 43.7 " hydrogen sulphide 0.1 ppm carbonyl sulphide 200,000 lbs SELEXOL The liquid is heated to 247 F in heat exchanger 33 and passed through conduit 34 into desorber 35. Part of the liquid leaving the bottom of the desorber 35 is vapourised in heat exchanger 36 and returned to the desorber 35 as reboil whilst the remainder passes through conduit 37 and heat exchanger 33. The liquid leaving heat exchanger 33 through conduit 38 is at 28 psia and is recycled to pump 39 where it is compressed to 650 psia. Make-up SELEXOL and water is introduced into the system by pump 40. Part of the gas leaving the top of absorber 35 is densed in heat exchanger 41 and returned to the desorber : @ @@ separator 42 as reflux. The residual gas comprising 9,0 lb moles water 131.9 " carbon dioxide 43, 7 " hydrogen sulphide passes along line 43 to Claus unit 44 and SCOT unit 45 where .uSstantially all the hydrogen sulphide is removed. The carbon dioxide, water vapour and any residual hydrogen sulphide are passed through an incinerator and vented to atmosphere via conduit 46. The residual stream 27 is expanded to 300 psia in expander 47 and the gas and liquid thus found are separated in separator 48. The gas phase leaves separator 48 through conduit 49 and is compressed to 646 psia in compressor 50. The liquid phase leaves separator 48 through conduit 51 and as expanded to 22 psia across valve 351. The expansion rouses the temperature to drop and the two phase mixture warmed by the expansion is warmed to 310F in heat exchanger 1, before entering phase separator 52 through conduit 53. The gaseous phase leaving separator 52 passes through heat exchanger 9 and leaves the plant at 840F and 18 psia. The aeous phase comprises (per hour) 10.5 lb moles hydrogen 0.7 " methane 1794.4 " carbon dioxide 4,1 " water 0.1 n carboncarbon monoxide 0.1 " nitrogen 0.8 ppm carbonyl sulphide The liquid leaving separator 52 passes through pump 54 and is divided into two streams 55 and 56. Stream 55 is introduced directly into desorber 57 whilst stream 56 is heated in heat exchanger 58 before entering desorber 57 through conduit 59. Air at 800F and 20 psia is introduced into the bottom of desorber 57 through conduit 6n and substantially all the carbon dioxide is desorbed and leaves the desorber 57 through vent 61 together with the air. The liquid leaving the bottom of desorber 57 comprises (per hour) 1438.3 Ib moles water 634,000 lbs SELEXOL and is at 20 psia and 31 F. The liquid is cooled to 24 F in refrigerator 62 and is pumped to 640 psia in pump 63 before being recycled through second absorber 16. Excluding the power requirements for the high and low temperature shift unit 4 and the energy requirements of the apparatus described can be summarized as 4.45 tons/hr steam (1 ton = 2240 lbs) 3250 hp Of the power requirements 1491 hp is required to drive compressor 32. Furthermore the compressor 32 represents a major item of capital expenditure, i.e. about 11.6% of the cost of the plant downstream of phase separator 7. It is an object of at least preferred embodiments of the present invention to provide a plant which will have a lower capital cost and a lower power consumption than that described with reference to Figure 1. It should perhaps be emphasized that the process described with reference to Figure 1 represents a refined state of a process which has been commercially used for about 20 years. The present invention is based on the concept of replacing valve 29 with a heat exchanger to heat the liquid leaving absorber 12 through conduit 28. By applying suffic @ent heat to the liquid sufficient phase separation can occur in separator 30 and compressor 32 can be dispensed with. Our yalculations indicate that there is both an energy and a capital saving to be made by modifying the design in this @anner. It should be noted that such a modification is not @ all it is considered that absorber 12 oper tes below normal ambient temperatur. According to the present invention there is provided a method of removing hydrogen sulphide from a gaseous mixture comprising hydrogen sulphide and carbon dioxide, which method comprises the steps of bringing said gaseous mixture into contact with a liquid absorbent in an absorber, heating the liquid leaving the absorber to desorb part of the hydrogen sulphide and carbon dioxide dissolved therein, separating said liquid from said desorbed carbon dioxide and hydrogen sulphide in a separator, recycling at least part of said desorbed carbon dioxide and hydrogen sulphide to said absorber, and expanding to said liquid downstream of said separator. Preferably, said method further comprises the step c; saturating said absorbent with carbon dioxide prior to introducing said absorbent into said absorber. This has the advantage of preventing further carbon dioxide uptake with its associated heat release taking place in the absorber and upsetting the heat balance therein. Preferably, the temperature of the absorbent in said absorber is maintained between 00F and 500F and more preferebly between 00F and 340F. Advantageously, substantially all the hydrogen sulphide is removed from the gaseous mixture in the absorber and substantially all the carbon dioxide is absorbed in a separate and distinct absorber. The hydrogen sulphide dissolved in the liquid absorbent leaving the heat exchanger is desorbed and is preferably passed through one or more Claus reactors. Whilst the present invention is particularly suited to the purification of synthesis gas (which may or may not be shifted) it is also suitable for the purification of other gaseous mixtures, for example natural gas and synthetic natural gas. For a better understanding of the present Invention reference will now be made, by way of example, to are 2 of the accompanying drawings which is a simplified @lowsheet of a plant in accordance with the present invention .d to Figure 3 which shows a modification to the plant shown in Figure 2.. Referring to Figure 2, 4006.4 1b moles/hr of synthesis @ comprising: 1756.5 lb moles hydrogen 22.2 n methane 6.4 " water 249.3 " carbon dioxide 1921.3 " carbon monoxide 41.9 " hydrogen sulphide 7.0 n nitrogen 1.8 n carbonyl sulphide at 660 psia and 122 F is heated to 310 F in heat exchanger C2. 5565.4 lb moles/hr of steam at 770 psia and 5140F are added to the synthesis gas at junction 103. The resulting mixture is subjected to high and low temperature shift react bs in unit 104 and the carbonyl sulphide is substantially hydrolysed in unit 105. The gas leaving unit 105 is cooled at 140 F in heat exchanger 106 and substantially all the tar which condenses in heat exchanger 106 is separated from ae gas in separator 107. The remaining gas is cooled to rn heat exchangers 108 and 109. The water which con denses in heat exchangers 108 and 109 is separated from the residual gas in phase separator 110 and the residual gas at 7 psia and 550F and compris ing (per hour): 3660.6 lb moles hydrogen 22.2 " methane 26.8 " water 2156.6 " carbon dioxide 15.7 " carbon monoxide 43.7 " hydrogen sulphide 7.0 N nitrogen 1.4 ppm carbonyl sulphide is cooled to 34 F in heat exchanger 202 before being introduced into absorber 112 through pipe 111. Substantially all the hydrogen sulphide in the residual gas is removed in absorber 112 and the gas leaving absorber 112 at 4 F through pipe 113 contains approximately 0.5 ppm hydrogen sulphide. The gas is heated in heat exchanger 202 and is then joined by gas from conduit 114 comprising (per hour): 26.5 lb moles hydrogen 0.2 N methane 27.5 n carbon dioxide 0.3 N carbon monoxide 0.07 N nitrogen 0.2 ppm carbonyl sulphide The water in the. resulting gas mixture is partially condensed in heat exchanger 115 and is then introduced into the bottom of a second absorber 116 where a substantial portion of the remaining carbon dioxide is removed. The gas stream leaves the second absorber 116 through conduit 117 at 26 0F and passes through heat exchanger 109 which it leaves through conduit 119 at 840F. The absorbent used in the process described with reference to Figure 2 is again SELEXOL. 100,000 lbs/hr SELEXOL together with 60 lb moles water enter second absorber 116 through conduit 120 at 640 psia and 240F. Simultaneously, 734,000 lbs/hr SELEXOL and 714.9 lb moles of water enter second absorber 116 through conduit 121 at 640 psia and 240F As the SELEXOL descends through the second absorber 116 it absorbs carbon dioxide and, in so doing, increases in temperature. The absorbent leaves the bottom of the second absorber 116 through conduit 122 at 34 0F and is then cooled to zs F in refrigerator 203. The cooled liquid is divided inte two streams 204 and 127. Stream 204 is expanded to 420 in in expander 205. The liquid leaving expander 205 is introduced into a column 206 where it is saturated with carton dioxide. The saturated liquid leaves column 206 through pipe 207 and, after being cooled to 50F in heat exchanger 208 and 00F in refrigerator 209 is pumped to 635 via by pump 210 before entering absorber 112. As the liquid descends through absorber 112 it absorbs hydrogen sulphide @@d in so doing is warmed to 34 F as it leaves absorber 112 conduit conduit 128. The liquid is pumped to 640 psia by 2Il and is then heated to 160 F in heat exchanger 212. e resulting two phase mixture is separated in separator The The gas leaving separator 130 through conduit 131 comprises (per hour): 4.3 lb moles hydrogen 0.2 n methane 0, 4 " water 178 " carbon dioxide 7.8 " hydrogen sulphide 0.3 ppm carbonyl'sulphide The gas in conduit 131 is joined by gas from conduit 213 6 and the combined stream is cooled to 95 F in heat exchanger @@3 before being introduced into the feed stream immediately upstream of separ ator 110. The liquid leaving separator 130 heated to 2680F in heat exchanger 214 and the resultant to phase mixture is separated in separator 215. The gas ase leaving separator 215 passes through conduit 213 whilst the liquid phase is expanded through expander 129 to 28 psia. the expanded liquid is heated in heat exchanger 133 and is - ased through conduit 134 into desorber 135, Part of the liquid leaving the bottom of the absorber 135 is vaporized in heat exchanger 136 and returned to the desorber as reboil whilst the remainder passes through heat exchanger 133. The liquid leaving the heat exchanger 133 through conduit 138 is at 26 psia and 3020F. Heat is extracted from this liquid in heat exchanger 214 and the liquid is then heated to 1950F in heat exchanger 216. The hot liquid is passed through heat exchanger 212 which it leaves at 440F. The liquid is then cooled to 240 in refrigerator 217 before being pumped to 640 psia by pump 139. Part of the water in the gas leaving the top of absorber 135 is condensed in heat exchanger 141 and returned to desorber 135 via separator 142 as reflux. The residual gas comprising (per hour): 7.4 lb moles water 96 " carbon dioxide 43.7 " hydrogen sulphide is passed along line 143 to Claus unit 144. The tail gasos leaving the Claus unit 144 are passed through SCOT unit 1 where substantially all the hydrogen sulphide is removed The carbon dioxide, water vapour and any residual hydrogen sulphide are passed through an incinerator and vented to atmosphere via conduit 146. The residual stream 127 is expanded to 300 psia in expander 147 and the gas and liquid thus found are separated in separator 148. The gas phase leaves the separator 148 through conduit 149 and is compressed to 634 psia in compressor 150. The liquid phase leaves separator 148 through conduit 151 and is expanded to 146 psia across valve 451. The expansion causes the temperature to drop to 220F and the two phase mixture formed by the expansion is passed to separator 226. The gas phase leaving separator 226 comprises almost entirely carbon dioxide with small traces of hydrogen, carbon monoxide and nitrogen. The gas is passed through conduit 227 to compressor 228 where its pressure is raised to 426 psia. The gas is then cooled in heat exchanger 229 before being introduced into column 206 where it saturates the SELEXOL with carbon dioxide. The liquid leaving separator 226 through conduit 230 is expanded to 26 psia across value 231 and is then warmed to 280F in heat exchangers 208 and 115. The two phase mixture leaving heat exchanger 208 trough conduit 153 is separated in separator 152. The gaseous phase leaving separator 152 passes through heat exchanger 109 and leaves the plant at 84 F and 22 psia. The liquid leaving separator 152 passes through pump 154 and is fed directly into desorber 157. Air is introduced into the bottom of desorber through conduit 160 and substantially all the carbon dioxide is absorbed and leaves the desorber 157 through vent 161 together with the air. The liquid leaving the bottom of desorber 157 comprises (per hour): 714.9 lb moles water 734,000 lbs SELEXOL and is at 22 psia and 26 0F. The liquid is cooled to 240F in refrigerator 162 and is pumped to 640 psia in pump 163 before teing recycled through second absorber 116. Excluding the power requirements for the high and low temperature shift unit 104 and the carbonyl sulphide hydrolysis unit 105 the energy requirements of the apparatus described can be summarized as 2.96 tons/hr steam (1 ton = 2240 lbs) 2190 hp It will be seen that the energy consumption of Figure 2 is 33.48* lower than the prior art described with reference bo Figure 1 and the horse power requirements are reduced by Furthermore, our calculations indicate that the stalled capital cost of the installation shown in Figure 2 should be 18.67% lower than the prior art described with reference to Figure 1. Various modifications to the installation shown in figure 2 are currently contemplated. Thus, if the feed mixture has a very low hydrogen sulphide content it may be advantageous to modify the treatment of the liquid leaving the bottom of the absorber 112 as shown in Figure 3. Parts in Figure 3 which have similar counter-parts in Figure 2 are dentified by the same reference numeral as used in Figure 2. In particular, the liquid which leaves the bottom of absorber 112 through conduit 128 is pumped through heat exchanger 212 0 where it is heated to 160OF. The resulting two phase mixture is separated in separator 130 and the gas phase returned to the absorber feed via cooler 433. The liquid phase leaves separator 130 through conduit 301 and is mixed with gas from conduit 302. The resulting mixture is flashed in vessel 303 and the gaseous phase returned to the absorber feed via conduit 304 and cooler 433 The liquid phase leaves separator 303 through conduit 305 and is heated to 2680F in heat exchanger 306 before being flashed to 304 psia across valve 307. The two phase mixture leaving valve 307 is separated in separator 308 and the gaseous phase is compressed in compressor 309 and passed through conduit 302. The liquid leaving separator 308 is passed through heat exchanger 133 and fed to desorber 135 The liquid leaving the bottom of the desorber is passed through heat exchanger 133 and is then conducted to heat exchanger 306 via conduit 138. The liquid leaving heat exchanger 306 is heated in heat exchanger 310 and is then passed through heat exchanger 212. The liquid leaving the heat exchanger 212 is then cooled in heat exchanger 217 and pumped to the top of the second absorber 116 by pump 139. The modification shown in Figure 2 requires more power than the installation shown in Figure 2 and is somewhat more expensive. It is still, however, considerably cheaper and more efficient than the arrangement shown in Figure 2. It is believed that the modification will be necessary only in unusual circumstances where the level of hydrogen sulphide in the feed streams is particularly low.

Claims: CLAIMS 1. A method of removing hydrogen sulphide from a gaseous mixture comprising hydrogen sulphide and carbon dioxide, which method comprises the steps of bringing said gaseous mixture into contact with a liquid absorbent in an absorber, and characterized in that it comprises the steps of heating the liquid leaving the absorber to desorb part of the hydrogen sulphide and carbon dioxide dissolved therein, separating said liquid from said desorbed carbon dioxide and hydrogen sulphide in a separator, recycling at least part of said desorbed carbon dioxide and hydrogen sulphide to said absorber, and expanding said liquid downstream of said separator. 2. A method according to Claim 1, characterized in that said method further comprises the step of saturating said absorbent with carbon dioxide prior to introducing said absorbent into said absorber. 3. A method according to Claim 1 or 2, characterized in that the temperature of the absorbent in said absorber is maintained between 00F and 500F. 4. A method according to Claim 3, characterized in that the temperature of the absorbent is maintained at between OOF and 340F.' 5. A method according to any preceding claim, characterized in that substantially all the hydrogen sulphide is removed from the gaseous mixture in the absorber and substantially all the carbon dioxide is absorbed in a separate and distinct absorber. 6. A method according to any preceding claim, characterized in that the hydrogen sulphide dissolved in the liquid absorb ent leaving the heat exchanger is desorbed and is passed through one or more Claus reactors.

Inventors: ANTONAS, IOANNIS ANTONIOS
Applicants: AIR PRODUCTS AND CHEMICALS, INC.
Application year: 1979
Main classification: B01D53
IPC codes: C10K1, B01D53, C07C7, C01B3