FI71501C - FOERFARANDE FOER SEPARERING AV NH3 OCH CO2 UR BLANDNINGAR SOM INNEHAOLLER DESSA. - Google Patents

Description

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^ (45) I t; /; r. t ty (51) K ».lk. * / lnt.a.« B 01 D 53/14 FINLAND —FINLAND (21) Patent application - Patentansökning 773282 (22) Application date - Ansökningsdag 02.11.77 (H) (23) Starting date - Giltighetsdag 02.11.77 (41) Has become public - Blivit offemlig 04,05,78

National Board of Patents and Registration Date of publication and publication.

Patent and registration authorities Ansökan utlagd och utl.skriften publicerad 10.10.86 (86) Kv. application - Int. ansökan (32) (33) (31) Privilege claimed— Begärd priority 03 · 1 1 - 76 O3.ll.76 Netherlands-Nederländerna (NL) 7612163, 7612162 (71) Stamicarbon B.V., P.O. Box 10, Geleen,

Unie van Kunstmestfabrieken b.v., P.O. Box 45, Utrecht, Netherlands-Nederland (NL) (72) Andreas Johannes Biermans, Urmond, Kees Jonckers, Born, Netherlands-Nederland (NL) (74) Berggren Oy Ab (54) Method for separating NH3 and CO2 from mixtures The present invention relates to a process for the separation of substantially pure NH 3 and substantially pure CO 2 from a mixture containing NH 3 and CCl 3 - The phrase "a mixture containing NH 3" means a process for the separation of substantially pure NH 3 and substantially pure CO 2. and CO 2 "means a mixture which may be a binary N 1 / CO 2 mixture, a gas mixture of NH 4, CO 2 and water vapor, or an aqueous solution of NH 3 and CO 2, optionally also containing ammonium carbamate and / or ammonium carbonate in solution. carbonate.

In some chemical processes, gas mixtures of N1 ^ m and CCl2 and optionally H2O are obtained as by-products. For example, in the synthesis of melamine, a gas mixture is obtained from urea which, in addition to melamine, contains NH3 and CO2 in an amount of at least 1.7 tonnes per ton of melamine. In order to recycle this N1 and CO2 after separation from the melamine, for example in the urea synthesis process, it is in most cases necessary to compress the gases to a higher pressure. Compression of the mixture requires special measures to prevent the condensation of NH 4 and CO 2 and the precipitation of the carbamate thus formed, so that the gas mixture is absorbed in the usual way into water or aqueous solution, resulting in ammonium carbamate solution which can be pumped to the urea synthesis process after desorption. and repeated absorption at elevated pressure to increase the concentration of the solution.

2 71501 A disadvantage of this process is that water recycled with NH 4 and CC 4 has an adverse effect on the production of urea from NH 4 and CC 4.

Unless any special measures are taken, it is impossible to remove all NH 3 and CO 2 'separately from the mixtures due to the fact that the binary system NH 3 / CO 2 and their tertiary systems with I 2 O form azeotropes.

Various methods have been proposed to avoid the need to recycle an unreasonable amount of water in conjunction with NH 4 recycled to the urea synthesis process, all of which involve the separation of N 1 -CO 2 mixtures into their constituents. Such processes involve the selective absorption of either ingredient into an aqueous solution. Thus, such a process is one in which NH 2 is absorbed into an aqueous solution of a strong acid ammonium salt (e.g. NH 4 NO 3) under elevated pressure. Another known process involves the selective absorption of CCl 4 by washing the gas mixture with an aqueous solution of an alkanolamine, e.g. monoethanolamine. (One disadvantage of such processes is that the absorbed component must then be removed from the absorbent and cleaned.

It has further been proposed to separate NH 3 and CCl 2 from mixtures of NH 3, CCl 2 and N 2 O 2 by distilling off most of the N 1 O 3 in the first step and then distilling off CC 3 in the second step, which is carried out in the NH 3 -CO 2 ^ O at higher pressure.

Such processes are described in U.S. Patent No. 3,112,177 and British Patent No. 1,129,939.

The U.S. patent describes a process in which, in a first step performed at a system pressure between 1 and 5 atm abs., CC> 2 is separated from a mixture of ΝΗ3 ·· η, CCl 2 and H 2 O, where the NH 3 -P 1 content is quite low , so. smaller than in the N1 ^ / CC ^ azotrope. The liquid base fraction obtained is stripped with a stripping gas, for example methane, at an absolute total pressure of 1 atmosphere. This results in a decrease in system pressure and the escape of NH 4 and a small amount of CO 2, so that a mixture of methane, NH 4 and CO 2 with a total absolute pressure of 1 atmosphere is obtained.

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To remove traces of CO2 contained in the gas mixture, a portion of the mixture is condensed, causing the CO 2 to be absorbed into the liquefied NH 4.

A similar process is described in British Patent 1,129,939. According to this patent, a gas mixture consisting of NH 4 and CO 2 with a higher NH 4 content than in an azeotrope is absorbed into water or an aqueous solution. At normal pressure, NH2 is distilled from the resulting aqueous solution. The remaining solution is then subjected to fractional distillation at a pressure between 5 and 20 atm abs. and heated to remove CO 2.

These two processes are based on the principle that changing the pressure of the system NH 2, CC 2 and H 2 O makes it possible to separate NH 3 at lower pressure and CC 2 at higher pressure. With respect to system pressure, i.e. The sum of the partial pressures ΝΗβΐη, CC ^ tn and H2O in the removal phase Nl ^ rn and the system pressure in the CO2 removal step in both processes must be between about 1: 5 and 1:20, if the separation is to progress smoothly.

These processes have the disadvantage that if the mixture to be treated becomes to be used at an absolute pressure of more than 1 atmosphere, it must first be expanded to an absolute pressure of 1 atmosphere. In addition, gaseous NH2 is released, which has an absolute maximum pressure of 1 atmosphere and possibly a large amount of other gas is also present. If this NH2 has to be further processed, eg in a urea synthesis process, it must be pressed to a higher pressure. The compression energy required for this is considerable and, in addition, the CO2 concentration in them must be kept low to prevent the formation of solid ammonium carbamate in the compressor and high-pressure pipes. Another possibility is the liquefaction of gaseous NH 4 by deep-cooling and then raising the NH 2 to the required pressure. This again requires energy.

The present invention provides a simple and less expensive method for separating substantially pure NH 3 and substantially pure CO 9 from mixtures of NH 4, CO 2 and optionally H 2 O in separate separation steps.

71501

Surprisingly, it has been found that it is possible to recover NH 4 and CO 2 separately from such mixtures without the problems outlined above if the mixture fed to the CO 2 separation step is diluted with water in a weight ratio between water and feed of between 0.2: 1 and 6: 1.

The invention provides a method for separating substantially pure λ 2 and substantially pure CC 2 from a mixture containing NH 4 and CC 2, which mixture is a gas mixture of NH 4 and CC 2 or NH 2. , A gas mixture of CC 2 and water vapor or an aqueous solution of NH 2 and CC 2, which separation is carried out in the NH 4 separation zone by heat and in the CO 2 separation zone by heat. The invention is characterized in that dilution water is fed to the CO 2 separation zone to such an extent that the weight ratio of the dilution water to the feed to the CO 2 separation zone is 0.2-6, and that the NH 2 / CO 2 / H 2 O system pressure in the CO 2 separation zone is at most twice the system pressure NH ^ -erotusvyöhykkeessä.

There are other advantages to adding water to the CO2 separation zone in amounts ranging from 0.5 to 2.5 times the weight of the feed entering it, such as achieving an optimal efficiency separation. Below this range, the difference becomes increasingly difficult to achieve due to the ever-increasing recycling rates required and at less than 0.2 times the difference becomes unsatisfactorily small. Above 2.5 times the value, the temperature at which the CO 2 becomes separated rises, which at the same time poses a risk of corrosion and, moreover, 6 times the weight calculated from the feed, the amount of energy required to remove the added water becomes unsatisfactory.

It should be emphasized that the amount of water added to the CO 2 separation zone according to the invention is in addition to all the water contained in the feed to said CO 2 separation zone, which may in fact consist of NH 3, CO 2 and H 2 O.

The process of the invention may relate to starting mixtures which are rich in NH2 or rich in CO2. By the phrase "high in NH 4" is meant that the NH 4 / CO 2 ratio of the starting mixture is such that, when heated, it is mainly discharged by NH 4, and by the phrase "rich in CO 2" is meant that the NH 4 / CO The ^ ratio is such that when heated, mainly CO2 is released.

Thus, one embodiment of the present invention is a process in which the starting mixture is a NH 4 -rich mixture as defined above, and in which it is passed said mixture to an NH 4 separation zone where gaseous NH 4 is obtained upstream and a constant boiling NH 4 an aqueous solution of rn and CO 2 is obtained as a bottom fraction, and b) said bottom fraction is fed as a feed to said CO 2 separation zone in the presence of water according to the invention, whereby CO 2 is obtained upstream and the boiling aqueous solution of NH 4 and CO 2 is obtained as a bottom fraction.

Another embodiment of the invention is a process in which said mixture is a CO 2 -rich mixture as defined above and in which process a) said mixture is introduced into a CC 4 separation zone in the presence of water according to the invention, whereby CO 2 is obtained upstream and constant boiling NH 2 and CO 2 b) introducing said bottom fraction into a desorption zone to remove substantially all of NH 2 and CO 2 therefrom as a gas mixture with water vapor, and c) introducing said gas mixture into said NH 4 separation zone where gaseous NH 4 is obtained upstream and unchanged a boiling aqueous solution of NH 4 and CO 2 is obtained as the bottom fraction.

According to the invention, water may be added to the CO 2 separation zone, e.g. from one or more points therein, and / or may be added at least in part to the feed to the CO2 separation zone before it enters said zone. The water added itself may be a dilute solution of NH 3 and CO 2 containing at least 90% by weight of water. However, this percentage depends on the system pressure. At low pressures, a higher percentage of water is required.

The ΝΗβΐη separation zone preferably operates at a base temperature of 60-170 ° C and a peak temperature of -35 to + 66 ° C and said CO 2 separation zone operates at a base temperature of 75-200 ° C and a peak temperature of 0-100 ° C, the peak temperatures of said zones being lower than the bottom temperature.

The optimum value of the bottom temperature in the NH 4 separation zone depends on the composition and pressure of the feed entering it. However, if a temperature below 60 ° C is used, either a range of solid reaction products of NH 4, CO 2 and possibly water is achieved, or the feed to the CO 2 separation zone has a composition such that optimal CC 2 S separation is not achieved.

71501

In most cases, the bottom temperature in the NH 4 separation zone is not higher than 170 ° C to keep it and the CC 4 separation zone temperature (which is always slightly higher) below a value that can lead to unpleasant corrosion. For the same reason, the bottom temperature in the separation zone CCl 4 is preferably not more than 200 ° C.

The peak temperature of the NH-^ separation zone is mainly determined by the pressure used. The peak temperature in the CO2 separation zone determines the final purity of the CCl2 achieved.

In general, NH3 contents of 100 ppm or less are easily achieved at peak temperatures of up to 100 ° C.

The temperatures at the peaks of the two separation zones are determined by the amount of any NH 4 return flow and the amount and temperature of dilution and wash water used.

The system pressure in the separation zone CCl is at most twice the system pressure in the separation zone NH 4, and most preferably the system pressures in the two separation zones are substantially the same.

Pressure ratios above 2 can be easily used, but they result in a decrease in the pressure at which NH 3 is released, necessitating a higher capital cost to liquefy NH 3. It also causes energy consumption to rise unreasonably high. For these reasons, it is recommended not to go beyond the maximum pressure ratios described above.

The process according to the invention can advantageously be carried out at equal system pressures in the NH3 separation zone and in the CCH separation zone. The use of pressure ratios that differ by a relatively small factor may be advantageous in terms of, for example, the thermal economy of NH 4 and CO 2 recovery and the size of the heat transfer surface area required.

In general, the absolute value of the system pressure at which the separation is performed is determined by the pressure at which the recovered NH3 is to be used. Preference is given to an absolute system pressure of 5-25 11 7 715 01 atmospheres in the NH4 separation zone, in particular an absolute pressure of 15-22 atmospheres- It is particularly advantageous to work at system pressures where gaseous NH3 can be liquefied by cooling water so that NH4 can be pumped and not no expensive compression step required.

It should be noted that in the process described in British Patent 1,129,939, CO 2 is washed in a scrubber tower with water to remove the last traces of NH3 in order to prevent precipitation which would have an adverse effect on the process. This results in the formation of an aqueous solution of NH 3 which is passed to a distillation tower to release CC 3. The amounts of water used in this process are very small and insufficient to achieve the effects of this invention.

As indicated above, either NH 3 or CO 2 is first separated depending on the NH 3 / CO 2 molar ratio of the mixture to be treated. However, if the mixture to be treated according to the invention is in the form of a very dilute aqueous solution of NH 3 and CO 2, it is advisable to remove substantially all or part of NH 3 and CO 2 in the first step to obtain a gas mixture and separate the resulting gas mixture into NH 4 and CO 2 according to the invention.

The invention can be applied particularly advantageously to the separation of a gas mixture consisting essentially of NH 3, CO 2 and IVO and obtained by decomposing unconverted ammonium carbamate γ at 1-25 kg / cm 2 in absolute pressure in the urea production process from NH 3 jacc> 2. The separation can then be performed by the following steps: a) passing said further gas mixture to a CO 2 separation column having a temperature gradient from thereon to the top, feeding water or an aqueous solution from one or more points along said separation column in an amount to form NH3 and CO2 an aqueous solution containing 65 to 96% by weight of water and a CO 2 gas stream substantially free of NH 3 and water and removing said solution from the bottom and said CO 2 stream from the top of said separation column, b) passing said NH 3 and An aqueous solution of CO 2 in a desorption zone where NH 3 and CO 2 are desorbed to provide a water stream substantially free of NH 3 and CO 2 and a second gas stream containing NH 3, residual CO 2 and water, and 71501 c) introducing said second gas stream to an NH 3 separation zone, wherein CC 2 and NH 2 are separated into an aqueous effluent and NH 2 is removed as a separate gas stream from said NH 3 separation zone and is discharged directly or indirectly into said synthesis zone, which steps a, b and c are carried out at an absolute pressure between about 1-25 kg / cm 2.

The invention is shown continuously in the accompanying drawings, in which Figure 1 is a schematic representation of a process according to the invention in which the separation zones of NH 3 and CO 2 are maintained at substantially the same pressure suitable for separating NH 3 and CO 2 from mixtures of azeotropic NH and Fig. 2 is a schematic representation of the process of Fig. 1 with the change that the separation zones for NH 3 and CO 2 operate at different pressures, Fig. 3 is a schematic representation of a process according to the invention in which NH 3, CO 2 and possibly H 2 O mixtures with a composition on the CCl 2 -containing side of the azeotrope can be separated, and Fig. 4 is a schematic representation of a process according to the invention suitable for separating NH 3 and CCl 2 from dilute solutions, Fig. 5 is a schematic representation of a process for the preparation of urea to which the invention applies.

Referring to Fig. 1, a mixture of NH 4 -containing NH 4, CO 2 = n and f 2, either as an aqueous solution or as a gas mixture, is fed through a pipe 1 and a pump 2 to a Nl 2 separation zone 3 in the form of a concentration column.

At the top of the column 3, gaseous NH 3 is discharged through a pipe 4 and is condensed in a cooling condenser 5 and an uncondensed gas mixture of NH 3 and other gases leaves the condenser upstream. Said other gases are derived from air, oxygen or an oxygen-releasing substance which is fed to the apparatus as a passivating agent to prevent unpleasant corrosion in the material from which the plant is made. A part of the passivation gas is fed through the compressor 6 and the pipes 7 and 8 to the NF 2 concentrator 3 and a part through the pipe 9 to the desorption device 10.

From the overflow exiting the condenser 5 through line 36, a stream of NH 3 is washed in a scrubber 11 by washing in scrubber 11 with water fed through conduit 34 to form an N1 solution, which is passed through pump 12 to recirculating condenser 13 and recycled to scrubber 11 through recycle 14. to the concentration column 3 via line 15, if necessary after depressurization.

The other gases separated in the scrubber 11 are discharged through the pipe 16 and, if desired, completely or partially discharged into the outside air through the pipe 19 and / or fed through the pipe 13 to the CO 2 concentration column 18.

A portion of the NH 3 liquefied in the condenser 5 is returned via line 20 to the NH 3 concentration column 3 as a reflux. The remainder of the liquefied NH 4 is discharged through line 37. The heat required for concentration is fed to the bottom of the column by means of a steam coil 35.

From the bottom of the NH 3 concentration column 3, the aqueous solution of NH 3 and CO 2 is discharged through line 21 and the solution is passed to the CO 2 concentration column 18 via line 21. The CO2 concentration column 18 operates at substantially the same system pressure as the NH 3 concentration column 3. The bottom stream from the desorption column 10 mentioned later is passed through the pump 22 and pipe 23 to the separation column 18 as a diluent. In order to improve the heat distribution, this liquid flow is first dissipated at the bottom of the CO 2 concentration column 18 and in the condenser 26. The rest of the heat required for concentration in the column 18 is supplied by heating coils 24 containing, for example, steam.

A portion of the liquid stream from the pipe 23 is passed through a pipe 25 which is cooled by water in a condenser 27 and discharged through a pipe 28. The wash water stream is fed to the CO 2 concentration column 18 via line 29 to remove as much NH3 as possible from the CO2.

The upstream flow in the tube 30 from the CO 2 concentration column 18 is a gas consisting of CO 3 and possibly other gases as described above, but containing essentially no NH 3.

The bottom stream from the CO 2 concentration column 18 is a dilute aqueous solution of NH 3 and CO 2 and is passed through line 31 to a desorption device 10 where almost all of NH 3 and 00 3 is expelled by heating 71501 10, for example with steam fed through heating coils 32.

The liquid from which almost all NH 3 and CO 2 have been stripped is passed via line 23 to a concentration column <302: η and the gas mixture of NH 3, CO 2 and H 2 O formed in the desorption device 10 is passed via line 33 to a concentration column 3 of NH 3 and CO 2.

Referring to Figure 2, in which the same portions as in Figure 1 are denoted by the same numerals, the CO 2 concentration column operates at a system pressure higher than the system pressure in the NH 4 concentration tonne. Compressor A and pump B are installed in pipes 17 and 21, respectively, to increase the pressure of the gas and liquid flows. Further, the pipe 31 includes a reduction valve C which reduces the pressure of the bottom stream coming from the CO 2 separation column 18. The desorption column 10 operates at a system pressure that is substantially the same as the system pressure in the NH 3 concentration column. Another suitable way is to carry out the desorption at the same system pressure as where the CO2 concentration column is.

Referring to Fig. 3, in which the same parts as in Fig. 1 are denoted by the same numerals, a CO 2 -rich ΝΗ 3 ·· η mixture of CO 2 and H 2 O is fed through a tube 41 and a pump 42 to a CO 2 concentration column 18.

According to this embodiment of the invention, it is possible to separate substantially pure CO 2 and substantially pure NH 4 from mixtures having a composition on the CO 2 -containing side of the azeotrope. With the exception of the position of the first feed stream, this Fig. 3 is identical to Fig. 1.

Referring to Figure 4, in which the same portions as in Figure 1 are denoted by the same numerals, a dilute aqueous solution of NH 3 and CO 2 is fed through line 51 and pump 52 to desorption column 10. Except for the position of this first feed stream, Figure 4 is identical to Figure 1.

Referring to Fig. 5, in the urea synthesis reactor 101, a urea synthesis solution is formed at a pressure of 110 to 250 kg / cm 2 and a temperature of 165 to 220 ° C. This solution is fed to the stripper 103 via a pipe 102 which can operate at a synthetic pressure or lower pressure and is brought into countercurrent contact - by adding heat - 11 715 01 with fresh CCl 3, which is fed through a pipe 104. A gas mixture consisting of Noi, CC 4 and i 2 Or and expelled in a stripper is passed through line 105 to a condenser 106, preferably operating at synthetic pressure, to which a certain amount of solution formed in the bottom of the urea reactor 101 is also fed through lines 107 and 108. by means of an ejector 109 for raising the sealing temperature so that the sealing heat is released at the highest possible temperature level. This heat is used to produce steam with an absolute pressure of 1-6 kg / cm, which can be used elsewhere in the process. The ejector 109 is operated by NH 4, which is fed from the storage tank 111 through the pump 110. Fresh NH3 is fed through line 112. In the cooler 106, complete or partial condensation and absorption of the gas mixture fed through the pipe 105 takes place. The carbamate solution thus obtained and the remainder of the gas mixture, if any, are passed to the urea reactor 101 via line 113 together with the carbamate solution obtained by condensing NH 4, CO and containing inert components brought in by fresh CO2, NH3 and possibly air or oxygen used for passivation of the equipment and pipes. This sealing takes place in an exhaust gas cooler 115, which is connected to a pipe 113 or to the lower part of the urea reactor 101 by means of a pipe 116. The gas mixture, which is not condensed in the purge gas cooler, is optionally passed through line 117 to an absorber (not shown) which operates at a lower pressure to recover the NH 4 still present. However, it is also possible to feed the solution obtained in the exhaust gas cooler 115 to the cooler 106 through the ejector 109.

The stripped urea synthesis solution is discharged from the stripper 103 and, when expanded to a pressure of 2-5 kg / cm 2 in the reduction valve 118, is passed through line 119 to a carbamate decomposer 120 where the carbamate still present is decomposed by virtually complete heating. The gas mixture exiting the heating zone of the carbamate decomposer 120 together with the urea solution is separated from the solution in the separator 121 and then passed to a plant for the separation of NH 3, CO 2 and H 2 O, which is described in more detail below. The aqueous urea solution obtained in the separator 121 of the carbamate decomposing device 120 is expanded in the expansion vessel 122 to normal pressure, at which some 12 715 01 of dissolved NH 4 and water vapor are removed, which are discharged through the pipe 123. Through line 124, the aqueous urea solution is passed to a compartment, which as a whole is numbered 125, where it is concentrated in a manner known per se by evaporation or crystallized or can be treated in another known manner. The final product is removed through line 126.

The gas mixture separated from the urea solution in the separator 121 is passed via a pipe 127 to a plant for the separation of NH 4, CO 2 and H 2 O, which plant mainly comprises a CO 2 separation column 128, a desorption column 129 and a ΝΗ ^: η separation column 130. These columns operate at virtually the same pressure between 1-25 kg / cm abs. A particular advantage is obtained if the separation system operates at an absolute pressure between 1 and 6 kg / cm, because in this case the solution fed to the desorption column 129 can be stripped with the steam generated in the condenser 106. Preferably, the separation system operates at the same pressure as the carbamate decomposer 120 and separator 121. At the bottom of column 128, the gas mixture is contacted with an amount of water such that substantially all of NH 3 and H 2 O is absorbed or condensed and CC> 2 with very little NH 3 and H 2 O rises. to the top of the column. For this purpose, water or a water solution, e.g. the process concentrate from compartment 125, is fed through line 131 and further aqueous solutions from the subsequent separation steps are recycled through lines 132 and 133. The gas mixture rising at the top of column 128 is washed with water fed through line 134 to remove small amounts of NH 4 still present therein.

This water, together with the NH 3 absorbed therein, also enters the bottom of the column. The total weight of water fed to the bottom of the column 128 through the tubes 131, 132, 133 and 134 is such that the solution discharged from the bottom of the columns contains 65-96% by weight of water, i. 5-20 times the weight of the gas mixture fed through line 127. Usually, the optimum amount of water is 80-95% by weight, and thus the amount of water to be added is preferably 8-15 times the weight of the gas mixture. By using the column in this way, the removal of CO2 and NH3 from the gas mixture separated in the separator 121 can be carried out with an attractively low total energy consumption in the separation compartment. The temperature at the bottom of the column is between 70-190 ° C while the temperature at the top of the column is between 30-70 ° C.

From the top of the column 128, CCl 3 is discharged through a pipe 135 containing only a few percent water and no more than traces of NH 4.

This gas is purified. The bottom product of column 128 contains virtually all of the NH 4 present in the gas mixture fed through line 127, inseparable CO 2, and water. This solution is passed through line 136 to a desorption column 129, where substantially all of NH 3 and CC 2 are desorbed, e.g., by fresh steam fed through line 137. The bottom product of the desorption column 129 is water, which now contains an amount of NH 3 and CO 2 below the detectable limit in the effluent streams in accordance with environmental control regulations and is discharged through a pipe 138 to which a pipe 132 for a CO 2 separation column 128 is connected. Some of the bottom product can be passed to a CO 2 separation column 128 via line 134 as wash water. The remainder is discharged through line 139 and may be returned in whole or in part to compartment 125.

The upper product of the desorption column 129, which has a temperature between 105 and 175 ° C and which, in addition to CO 2 and water, contains practically all the NH 3 separated in the separator 121 and also a certain amount of circulating NH 3, is passed to the NH 4 separation column 131 via a pipe 140 , through a condenser 141 and a pipe 142, in which column CO 2 and water are removed by means of water fed through a pipe 143 and liquid NH 3 fed through a pipe 144. The substantially pure NH 4 thus obtained is passed through a pipe 145 to a cooling plant 145 where it is sealed. Next, the liquid NH3 continues through line 147 to the NH3 storage tank 111. The bottom solution formed in the NH3 separation column consisting of NH3, CO2 and water is returned to the CO2 separation column 128 via line 133.

In the process described above, no water is returned to the urea synthesis from the low pressure step, i. part of the process after the reduction valve 118. As a result, a significantly higher degree of conversion is achieved in the urea reactor, so that a synthesis solution containing less carbamate is obtained. Because less carbamate needs to be decomposed in the stripper and stripping goes better due to the lower water content of the synthesis solution, a smaller amount of steam is needed in the stripper. The consumption of steam required for NH3 / CO2 / H2O separation is compensated by a reduction in the amount of steam in the treatment of the urea solution due to the lower amount of water to be evaporated. Since no solution but only gaseous NH 4 is recycled from the low pressure stage, the operation of the plant is simpler. Operational reliability is increased because vulnerable carbamate pumps are no longer needed.

14 715 01

The following practical examples of the invention are presented. In Examples 1-4, all percentages are by weight unless otherwise indicated.

Example 1

In an apparatus such as that shown in Figure 1, substantially pure NH 2 and substantially pure CO 2 were separated from a mixture of NH 3, CO 2 and water. At an absolute system pressure of 18 atm and a temperature of 73 ° C, an aqueous solution of NH 3 and CO 2 consisting of 33.4% NH 4, 18.2% CO 2 and 48.4% water was fed NH to the bottom of the concentration column 3 at a rate of 51,972 kg / h. 635 kg / h of air was fed through the compressor 6, of which 248 kg / h was fed to the NH 3 concentration column 3 and 387 kg / h to the desorption device 10.

A gas mixture containing 50.4% NH 4, 17.7% CO 2 and 30.9% H 2 O and up to 1.0% other gases at a temperature of 162 ° C from desorption apparatus 10 was also fed to NH 3. to the concentration column at a rate of 39,304 kg / h. A gas mixture containing 99.4% NH 2 and 0.6% other gases was discharged from the top of this column at a rate of 65,217 kg / h. With cooling water, part of this gas mixture was liquefied in condenser 5. A reflux stream comprising 44,630 kg / h of this liquefied mixture was passed to column 3. A gas mixture containing 88.9% NH 4 and 11.1% other gases was removed from condenser 5 at a speed of 3474 kg / h. This was washed in a scrubber with 11 4496 kg / h of water. The heat was removed from scrubber 11 in a recycle condenser 13. The solution containing 38.7% NH 3 and 61.3% H 2 O was recycled back to the NH 3 concentration column at 7335 kg / h. The temperature at the top of this column was 46 ° C. Through pipes 16 and 17, other gases were introduced to the CO 2 concentration column 18 at a rate of 635 kg / h.

On the basis of the NH 3 concentration column 3, a liquid containing 25.4% NH 4, 21.0% CO 2 and 53.6% water and a temperature of 131 ° C was passed through line 21 to a CO 2 concentration column 18. at a speed of 78,027 kg / h.

This column, which also had an absolute pressure of 18 atm, flowed through line 25 with a stream of 73,299 kg / h of diluent containing water and traces of NH 3 and CO 2 and a liquid temperature of 206 ° C as it exited the desorption apparatus. . The liquid transferred some of its heat to the bottom of the CC> 2 concentration column 18 and the remaining heat was recovered in the condenser 26.

15 71 501

The liquid was removed from the desorption device 10 at a rate of 108,909 kg / h, which meant that 35,610 kg / h of water was removed, possibly by freezing in a condenser 27. This water could be used, for example, to absorb NH 4 and CO 2.

Wash water was added to the top of the CO 2 concentration column at a rate of 5959 kg / h to wash off the last traces of NH 4. With steam, the bottom temperature of the CO 2 concentration column 18 was maintained at 158 ° C. The peak temperature was 35 ° C. A gas mixture containing 93.7% CCl 4 and 6.3 l of other gases and less than 100 ppm NH 3 was removed from the column at a rate of 10,094 kg / h.

From the bottom of column 18, a solution containing 81.9% H 2 O, 13.4% NH 3 and 4.7% H 2 O and a temperature of 158 ° C was passed to desorption apparatus 10 at a rate of 147,826 kg / h.

With the aid of steam, substantially all of the NH 3 and CCl 2 were stripped from the solution in a desorption apparatus so that water containing only traces of NH 3 and CCl 4 was removed at a rate of 108,909 kg / h. The temperature at the top of the desorption apparatus was 161.8 ° C.

Example 2

In an apparatus such as that shown in Figure 2, substantially pure NH 3 and substantially pure CCl 4 were separated from a mixture of NH 3, CCl 4 and H 2 O.

At an absolute system pressure of 15 atmospheres and a temperature of 73 ° C, an aqueous solution of NH 3 and CO 2 consisting of 33.4% NH 3, 18.2% CO 2 and 48.4% water was fed at a rate of 51,972 kg / h to the bottom of NH3 concentration column 3. Air was fed through the compressor 6 at a rate of 635 kg / h, of which 248 kg / h went to the NH 3 concentration column 3 and 387 kg / h to the desorption unit 10. The NH 3 concentration column 3 was further fed with 37,943 kg / h gas mixture containing 49, 1% NH 3, 15.5% CO 2 / 34.4% H 2 O and 1.0% other gases at a temperature of 158 ° C, which mixture came from the desorption apparatus 10.

A gas mixture containing 99.0% NH 3 and 1.0% other gases was discharged from the top of this column at a rate of 60,888 kg / h. By deep cooling, a portion of this gas mixture was liquefied in condenser 5. A stream of 40,056 kg / h of this liquefied mixture was passed to column 3 as a reflux mixture. A gas mixture containing 81.7% 16 71501 NH 3 and 18.3% other gases exited the condenser 5 at a rate of 3474 kg / h. This mixture was washed in a scrubber with 11,500 kg / h of water.

The heat was removed from the scrubber 11 via a recycle cooler 13. The solution containing 36.2% NH 3 and 63.8% H 2 O was returned to the NH 4 concentration column at a rate of 7842 kg / h. The temperature at the top of the column was 39 ° C.

A stream of 635 kg / h of inert gas was passed through lines 16 and 17 and a compressor Δ to a CO 2 concentration column 18 operating at 25 atmospheres absolute system pressure.

A stream of 77,172 kg / h of liquid with a composition of 24.1% NH 3, 19.9% CO 2 and 56.0% H 2 O and a temperature of 128 ° was passed through line 20 and pump B. C, CC> 2 to the concentration column 18.

This column 18, which also had an absolute pressure of 25 atmospheres, was supplied via line 25 with a stream of 48,915 kg / h of diluent consisting of water with traces of NH 4 and CO 2 and a liquid temperature as it exited the desorption apparatus of 197 ° C.

The liquid transferred some of its heat to the bottom of the CO 2 concentration column 18 and the remaining heat was removed in a condenser 26.

The liquid was discharged from the desorption unit at a rate of 85,031 kg / h, which meant that 36,116 kg / h of water was discharged by cooling in a condenser 27, which could be used, for example, to absorb ΝΗβΐη and CO 2.

Wash water was added to the top of the CO 2 concentration column at a rate of 5959 kg / h to wash off the last traces of NH 4 as much as possible. By means of steam, the bottom temperature in the CO 2 concentration column 18 was maintained at 170 ° C. The peak temperature was 35 ° C. A gas mixture containing 3.7% CCl 3 and 6.3% other gases and less than 100 ppm NH 2 exited the top of the column at a rate of 10,094 kg / h.

From the bottom of column 18, a solution containing 80.0% H 2 O, 15.2% NH 3 and 4.8% CO 2 and a temperature of 170 ° C was passed to desorption apparatus 10 after being pressurized. reduced with reduction valve C.

17 71501 With the aid of steam, substantially all of the NH 4 and CO 2 were stripped from the solution in a desorption apparatus so that water containing only traces of NH 4 and ECU was discharged at a rate of 85,011 kg / h. The temperature at the top of the despiration device was 158 ° C.

Example 3

In an apparatus such as that shown in Figure 3, substantially pure NH 3 and substantially pure CCl 4 were separated from an aqueous solution of NH 3 and CCl 4 on the CO 2 side of the azeotrope. At an absolute system pressure of 18 atmospheres and a temperature of 130 ° C, an aqueous solution of NH 4 and CO 2 consisting of 20% NH 4, 20% CO 2 and 60% water was fed to a CC 18 concentration column 18 at 50 ° C. 000 kg / h.

On the basis of the NH 3 concentration column 3, a liquid containing 26.0% NH 4, 22.1% CO 2 and 51.9% H 2 O and a temperature of 130 ° C was passed through line 21 to the CCl 2 concentration column. 18 at a speed of 27,405 kg / h.

A flow of 54,990 kg / h of diluent consisting of water and traces of NH 3 and CC 2 with a liquid temperature of 206 ° C as it exited the desorption apparatus 10 entered this column 18 via line 25. The liquid transferred some of its heat to the bottom of the CC> 2 concentration column 18 and the remaining heat was removed in a condenser 26.

The liquid was removed from the desorption device 10 at a rate of 95,278 kg / h, which meant that 40,288 kg / h of water was discharged, possibly with cooling in the condenser 27. This water could be used, for example, to absorb NH 3 and CO 2.

Wash water was added to the top of the CO2 concentration column at 6347 kg / h to wash off the last traces of NR3. With steam, the bottom temperature in the CCl 2 concentration column 18 was maintained at 160 ° C. The peak temperature was 35 ° C. A gas mixture containing 94.6% CCl 3 and 5.9% other gases and less than 100 ppm NH 3 exited this column at a rate of 10,694 kg / h.

From the bottom of column 18, a solution containing 82.0% H 2 O, 13.3% NH 3 and 4.7% CO 2 and a temperature of 160 ° C was passed to desorption apparatus 10 at a rate of 128,683 kg / B.

18 71 501 With the aid of steam, substantially all of the NH 3 and CC 2 were stripped from the solution in a desorption apparatus so that water containing only traces of NH 3 and CC 4 was discharged at a rate of 95,278 kg / h. The temperature at the top of the desorption apparatus was 161.8 ° C. 635 kg / h of air was fed through the compressor 6, of which 248 kg / h was fed to the NH 3 concentration column 3 and 387 kg / h to the desorption device 10.

A gas mixture containing 50.7% NH3, 17.9% CO2> 30.4% H2O and up to 1.0% other gases from desorption unit 10 was fed to an NH3 concentration column at a rate of 33,405 kg /B.

A gas mixture containing 99.0% NH 4 and 1.0% other gases was discharged from the top of this column at a rate of 62,360 kg / h. With cooling water, part of this gas mixture was liquefied in condenser 5. A reflux stream of 49,071 kg / h of liquefied mixture was passed to column 3. A gas mixture containing 80.7% NH 3 and 19.2% other gases exited condenser 5 at 3289 kg /B. This was washed in a scrubber with 11,000 kg / h of water. The heat was removed from scrubber 11 by recirculating condenser 13. A solution containing 39.9% NH 3 and 60.1% H 2 O was recycled to the NH 3 concentration column at 6654 kg / h. The temperature at the top of this column was 50 ° C. Through tubes 15 and 16, other gases were introduced to the CC> 2 concentration column 18 at a rate of 635 kg / h.

Example 4

In an apparatus such as that shown in Figure 4, substantially pure NH 4 and substantially pure CO 2 were separated from their dilute aqueous solution at 18 atmospheres absolute system pressure.

At a rate of 50,000 kg / h NH 3 and an aqueous solution of CC> 2 with a composition of 82.0% H 2 O, 4.7% CC> 2 and 13.3% NH 3 were fed to a desorption apparatus 10. From the bottom of the column] 3 the solution was passed , having approximately the same composition and a temperature of 159.6 ° C, to the desorption apparatus 10 at a rate of 38,166 kg / h.

With the aid of steam, substantially all of the liquid in the desorption apparatus 10 was stripped with NH 3 and CC 2 in the desorption apparatus so that water containing only traces of NH 3 and CC 4 was discharged at a rate of 65,279 kg / h. The temperature at the top of the desorption apparatus was 162 ° C. 635 kg / h of air was fed through the compressor 6, of which 400 kg / h was fed to the NH 3 concentration column 3 and 235 kg / h to the desorption device 10.

Il 19 715 01

A gas mixture containing 50.7% NH 4, 17.9% CO 2, 30.4% i 2 O and up to 1.0% other gases from desorption apparatus 10 was fed to the NH 3 concentration column at a rate of 23,522 kg /B. A gas mixture containing 98.4% NH 3 and 1.6% other gases was discharged from the top of this column at a rate of 39,099 kg / h.

With cooling water, part of this gas mixture was liquefied in condenser 5. A reflux stream of 29,160 kg / h of liquefied mixture was passed to column 3. A gas mixture containing 80.7% NH 4 and 19.3% other gases was removed from condenser 5. at a speed of 3289 kg / h. This was washed in a scrubber with 11 4000 kg / h of water. The heat was removed from scrubber 11 in a recycle cooler 13. A solution containing 39.9% NH 4 and 60.1% H 2 O was recycled to the NH 4 concentration column at a rate of 6654 kg / h. The temperature at the top of this column was 50.6 ° C. Through pipes 15 and 16, other gases were passed to the CO 2 concentration column 18 at a rate of 635 kg / h.

On the basis of the NH 3 concentration column 3, a liquid containing 25.1% NH 3, 20.5% CO 2 and 54.4% H 2 O and having a temperature of 132 ° C through line 21 was passed to a CC 2 concentration column 18 at a rate of 18 20,237 kg / h.

This column 18 was fed via line 25 with a stream of 18,801 kg / h of diluent consisting of water and traces of NH 4 and CO 2, the liquid temperature of which as it exited the desorption apparatus 10 was 206 ° C. The liquid transferred some of its heat to the bottom of the CO 2 concentration column 18 and the remaining heat was removed in a condenser 26.

The liquid was discharged from the desorption device 10 at a rate of 65,279 kg / h, which meant that 46,478 kg / h of water was discharged, possibly with cooling in the condenser 27. This water could be used, for example, to absorb NH 3 and CO 2.

Wash water was added to the top of the CO 2 concentration column at a rate of 1491 kg / h to wash off the last traces of NH3. With steam, the bottom temperature of the CO 2 concentration column 18 was maintained at 159.6 ° C. The peak temperature was 42 ° C. A gas mixture containing 80.4% CO2 and 19.6% other gases and up to 100 ppm NHgra was removed from this column at a rate of 2995 kg / h.

20 71 501

Example 5 The amounts reported in this example are in kmol / h. 2317 kmol of NH 4 was fed through line 112 and 1143 kmol of CC 2 in line 104 to a plant as shown and described above in Figure 4 to produce 1019 kmol / h of urea in aqueous solution. The pressure in the urea reactor 101, the stripper 103, the condenser 106 and the exhaust gas cooler 115 was about 140 kg / cm 3. In the stripper, 103 2542 kmol of NH 4, 551 km of CCl and 132 kmol of CH 2 were expelled from the synthesis leach from urea reactor 101 with fresh CCl 2 with heating at 215 ° C. In the condenser 106, the gas mixture discharged through line 105 was partially condensed by a mixture - which was fed through line 108 - with fresh NH 4 and a solution drawn from the urea reactor. From the gas / liquid mixture thus formed and the carbamate solution formed in the exhaust gas cooler 115 containing 497 kmol of NH 3, 169 kmol of CO 2 and 33 kmol of H 2 O, a synthesis solution was formed in the urea reactor 101 at an average temperature of 183 ° C, which solution contained In addition to 1061 kmol of urea, 2683 kmol of NH3, 627 kmol of CC> 2 and 1191 kmol of P2O.

The CCl sn conversion therefore increased to 62.8%.

A solution of 1019 kmol of urea, 225 kmol of NH 4, 118 kmol of CCl and 1019 kmol of CH 2 O was discharged from a stripper 103, which solution was heated to 124 ° C in a carbamate decomposer 120 after the mixture was expanded to 2. 5 kg / cm. A gas mixture was formed containing 151 kmol of NH 3, 98 kmol of CCl 3 and 159 kmol of CH 2 O, which was separated in the separator from the remaining urea solution 121, which in addition to 1019 kmol of urea and 860 kmol of I 2 O still contained 74 kmol. NH 3 and 20 kmol CO 2 = a. NH 3 and CO 2 were largely removed in the expansion tank 122 to give an aqueous urea solution.

The gas mixture separated in separator 121 was passed to the bottom of a CO 2 separation column 128 where a pressure of 2.5 kg / cm was maintained and treated with a process concentrate at 90 ° C from the last treatment compartment containing 1117 kmol of H 2 O. , 5 kmol of NH3 and 2.25 kmol of CCl2. Further, column 128 was fed via line 132 with 3700 kmol of H 2 O, which contained traces of them, through line 133 with 739 kmol of H 2 O, 298 kmol of NH 4 and 134 kmol of CO 2 and through line 134 with 157 kmol of H 2 O. The temperatures of the solutions fed through lines 132 and 133 were 88 and 85 ° C, respectively, and the temperatures of the wash water fed through line 34 and 71501 21 were 40 ° C. The gas mixture leaving the top of column 128 contained 143 kmol of CO 2 and 3 kmol of H 2 O. A stock solution at 106 ° C containing 5939 kmol of H 2 O, 461 kmol of NH 4 and 92 kmol of CO 2 was passed to a desorption column 129. To this column was also fed a gas mixture recovered from the 115 peak products cooled by the exhaust gas, which contained 156.5 kmol of NH 4, 39.2 kmol of CCl 2 and 371 kmol of H 2 O, and steam at 138 ° C was passed to the bottom of the desorption column 129. From the bottom product of this column, which contained 6972 kmol of H 2 O, which still had residual NH 3 and CO 2, the already mentioned part was fed through pipes 132 and 134 to a CO 2 separation column 128 and the remaining part was partly passed as washing water to the NH 3 separation column. 130 via pipe 143 and the remainder was used elsewhere in the process or discharged into the sewer.

The gas mixture discharged from the top 129 of the desorption column at 110 ° C and containing 618 kmol of NH 2, 134 kmol of CO 2 and 709 kmol of H 2 O was separated - after cooling in an 87 ° C condenser 141 - NH 3 in a separation column 130 by washing with 29 kmol of water at 40 ° C and 170 kmol of NH 4 at -15 ° C to give 490 kmol of pure NH 2.

Claims (17)

A process for separating substantially pure NH 2 and substantially pure CO 2 from a composition containing NH 2 and CO 2, which composition is a gas mixture of NH 2 and CC> 2 or a gas mixture of NH 2, CC> 2 and water vapor. or an aqueous solution of NH 2 and CO 2, which separation is carried out in an NH 2 separation zone by the influence of heat and in a CO 2 between the dilution water and the inflow into the CO 2 separation zone is 0.2-6, and that the NH 2 / CO 2 / H2 O system pressure in the CO 2 separation zone is at most twice the system pressure in the NH 2 separation zone. Process according to Claim 1, characterized in that said amount of water is 0.5-2.5 times the weight of the feed into the CO 2 separation zone. Process according to claim 1 or 2, characterized in that the water is a solution of NH 2 and CC> 2 and contains at least 90% by weight of water. A process according to any one of claims 1-3, wherein said composition is an NH 2-rich composition, a) said composition is introduced into the NH 2 separation zone, in which gaseous NH and CO 2 is obtained as bottom flow, and b) said residual flow is transferred as flow to the CO 2 separation zone, in which CO 2 is obtained as peak flow and a constant boiling aqueous solution of NH Process according to claim 4, characterized in that said aqueous bottom flow solution obtained from the CO 2 separation zone is treated in a desorption zone to remove substantially all of NH is brought to the CO 2 separation zone. 6. A process according to claim 4, characterized in that said aqueous solution obtained from the CO 2 separation zone is treated to remove substantially all NH 2 and CC 2 from there in the form of a gas mixture with water vapor. at least a portion of said gas mixture is introduced into the NH 2 separation zone. Process according to any one of claims 1-3, wherein said composition is a CC> 2-rich composition, characterized in that a) said composition is introduced into the CO 2 separation zone, in which CC> 2 is obtained as a peak flow and a constant boiling aqueous solution of NH 2 and CO which gaseous NH₂ is obtained as peak flow and a constant boiling aqueous solution of NH₂ and CO₂ is obtained as bottom flow. Method according to claim 7, characterized in that said aqueous bottom flow is introduced into the CC> 2 separation zone. Method according to any one of claims 1-8, characterized in that the NH 2 separation zone is operated at a bottom temperature of 60-170 ° C and a peak temperature of -35 - + 66 ° C, and that the CO 2 separation zone is operated. with a bottom temperature of +75 - + 200 ° C and a peak temperature of from 0 - + 100 ° C, the peak temperature for said zones being lower than the bottom temperature. Method according to any of claims 1-9, characterized in that the CO 2 separation zone is operated at one. NH 2 / CO 2 / H 2 O system pressure, which is substantially equal to the system pressure in the NH 2 separation zone. Method according to any of claims 1-10, characterized in that said NH NH separation is carried out at an absolute NH₂ / CC₂ / t₂O system pressure in the range of 5-25 atmospheres. Method according to claim 11, characterized in that the NH 2 separation is carried out at an absolute NH 2 / CC 2/10 system pressure of 15-22 atmospheres. IN! 71 501 27 The process of claim 1, wherein said composition is a gas mixture consisting essentially of NH 2, CO 2 and E 2 from NH 3 and CO2 t characterized in that a) said further gas blast is introduced into a CO 2 separation column with a temperature gradient from the bottom to the top, water or an aqueous solution is introduced at one or more places along the CO an aqueous solution of NH 2 and CO 2 containing 65-96% by weight of water and a gas stream of CO 2 which is substantially free from NH 3, and water and said solution are removed from the bottom of the separating column and said stream of CO ) said aqueous solution of NH ^ and CO₂ is introduced into a desorption zone in which NH₂ and CO₂ are desorbed to obtain a stream of water which is substantially free of NH₂ and CO₂, and a second gas stream contains , NH 2, residual CO 2 and water, and c) said second gas stream is fed to an NH 2 separation zone, in which CO 2 and H 2 or indirectly to said synthesis zone, wherein the CO 2 separation step, desorption step and NH 2 separation step are carried out at an absolute pressure within the range of 1-25 kg / cm 2. Process according to Claim 13, characterized in that in step a) water is added in such an amount that an aqueous solution of NH 2 and CO 2 containing from 80 to 95% by weight water is formed. Method according to claim 13 or 14, characterized in that the CO 2 separation column, the desorption zone and the NH 2 separation zone are operated at an absolute pressure within the range 1-6 kg / cm 2. Process according to Claim 13 or 15, characterized in that the CO 2 separation column, the desorption zone and the NH 2 separation zone are operated at substantially the same pressure. Method according to any of claims 13-16, characterized in that the aqueous effluent from the NH4 separation zone is fed back to the CC> 2 separation zone.