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Wen Xiaozhong (Hunan Chemical Industry Vocational and Technical College) As is well known, in the urea production process based on complete water solution circulation, although the technology is mature and reliable and requires low investment, it results in high consumption. The proper operation of the medium-pressure absorption tower is not only one of the key factors affecting safe and stable operation, but also crucial for achieving high yields with low consumption. Therefore, when comprehensively optimizing urea production, we should first strive to identify the bottleneck issues affecting production, carefully optimize the innovative operations of the medium-pressure absorption system, and conduct thorough scientific and rational analysis and evaluation in order to achieve high yields with low consumption. 1 Description of the medium-pressure absorption process: The distilled gas coming out of the first column (at about 125°C) is sent to the heat recovery section at the bottom of the first-stage evaporator, where it mixes with dimethyl liquid and undergoes a reaction that releases heat; this heat is then used for the evaporation of urine. The gas-liquid mixture exits the first-stage evaporation heater (at about 115°C) and enters an external absorption cooler, where it is cooled with soft water (to about 110°C). It then proceeds to the bubbling section at the bottom of the medium-pressure absorption tower, where it undergoes bubbling absorption in conjunction with recycled ammonia and ammonia water. The unabsorbed ammonia and carbon dioxide rise to the purification section, where they are further washed and absorbed using concentrated ammonia water. The gas emerging from the top of the absorption tower enters two ammonia coolers in series; these coolers condense the gaseous ammonia. The non-condensable gases proceed to the inert gas scrubbing tower, where they are absorbed using dilute ammonia water as the absorbent. The unabsorbed inert gases go to the tail scrubber tower, where any remaining gaseous ammonia is recovered before being released. 2 Calculation of the liquid-phase components in the first-stage external cooler: Using 1 ton of urea as a basis for calculation, it is assumed that all of the CO2 in the gas from the first stage of decomposition is absorbed before the first-stage external cooler; in this case, CO2 will exist in the form of ammonium carbonate. Approximately 364.762 kg of CO2 is present in the gas from Stage 1; assuming that all of it is absorbed to form ammonium methoxide before entering the external absorption cooler, 1116.678 kg of ammonia in the gas from Stage 1 participates in the condensation and absorption reaction. The liquid-phase composition of the external cooler after absorption is shown in Table 1; among its components, CO2 accounts for 41.35% and NH3 accounts for approximately 38.02% (the actual value in operation should be lower than this). The melting point of this solution is around 90–100°C. http://pub2.hi2000.com/upload1/0712171539394615.jpg 2NH3+CO2→NH4COONH2 (1—1) 34 44 78 Based on equation (1—1) and Table 1, assuming that all the CO2 in the decomposition gas and dimethyl liquid reacts completely, the amount of ammonium methylate produced is: 78×51.987÷44+78×364.762÷44=738.782kg 3 Heat balance parameters of the external cooler http://pub2.hi2000.com/upload1/0712171540276027.jpg In Table 2, the heat of formation of ammonium methylate is 1160835 kJ/t·ur; 438270 kJ/t·ur is lost through the external section of the heat utilization unit, and 430181 kJ/t·ur is lost through the external cooler, accounting for approximately 75% of the total heat of formation of ammonium methylate. The heat removed from the primary external cooler is transferred away through a cooling cycle that uses soft water; the flow rate of water is approximately 2 m3/t·ur, and the temperature difference of the soft water entering and leaving the primary external cooler is around 50°C. 4 Operation of traditional medium-pressure absorption systems. General theory suggests that the CO2 in the gas phase from the decomposition stage is completely absorbed to form methammonium, which exists in the liquid phase of the external absorber cooler ； There have also been studies that treat the external cooler as a liquid phase, thereby ignoring the existence of the gas phase. Based on the previous calculations and the aforementioned theory, it is estimated that in the external cooler, about 41.25% of the CO2 in the liquid phase exists in the form of ammonium formate; the melting point of this compound is approximately 90–100°C. These conditions determine the operating parameters of conventional medium-pressure absorption towers. (1) Control of the first-stage external cooler: Since the melting point of the components in the first-stage external cooler is between 90 and 100°C, it is necessary to maintain an operating temperature about 10°C higher than this melting point; therefore, the temperature of the gas-liquid mixture entering the first-stage tower should be controlled at 100–110°C. Due to the relatively high melting point, in normal production (operation), the temperature of the gas-liquid mixture entering the absorption tower is rarely kept below 100–110°C; this is one of the conditions that require optimization of the operating procedures, and it also constitutes the basic premise for the optimization discussed in this paper. (2) Control of soft water entering the primary external cooler: If the amount of soft water entering the primary external cooler is 2.0 m3/t·ur, it is necessary to ensure that the temperature difference between the soft water entering and leaving the external cooler is 50°C. To prevent localized crystallization inside the external cooler upon suction, the outlet water temperature is maintained at 95–100°C, which results in an inlet water temperature of 45–50°C. This reveals the objective contradiction that \"the water temperature in the external cooler is much lower than the melting point of the solution inside it, which may lead to local crystallization of the solution within the external cooler.\" This is also the second condition for optimization in this paper. (3) Theoretically, controlling the temperature in the bubbling section (bottom) of the medium-pressure absorption tower shows that both NH3 and CO2 release heat during dissolution and the formation of ammonium methide within the absorption tower; therefore, a lower temperature is favorable for absorption. Moreover, a high temperature in the bubbling section increases the load on the polishing section of the absorption tower, and may even lead to excessive CO2 levels in the gas phase at the top of the tower, worsening operational conditions and affecting production. Therefore, to ensure a normal gas phase temperature, the process specifies that the temperature in the bubbling section of the absorption tower should be maintained within the range of 80–90°C. Traditional conservative practices are widespread in production; in fact, even under high load conditions, the temperature of the bubbling section is kept below 80°C to ensure system stability. Therefore, it is recommended to maintain the temperature at the bottom of the absorption tower between 90 and 95°C; this is the third condition for optimizing production as discussed in this paper. (4) Control of the horizontal balance in a single absorber: The traditional horizontal balance for a single absorber is maintained using about 260.697 kg/t·ur (0.32 m3/t·ur) of dimethyl ether, and 196.101 kg/t·ur (0.16 m3/t·ur) of ammonia water. Under the above conditions, a single urea production unit operates stably when its production capacity is between 400 and 450 t/d, but faces operational difficulties when the capacity reaches 500 to 550 t/d. Its main manifestations are: a high absorption load on the absorption tower, a large amount of ammonia required for reflux, increased load on the ammonia cooler, a large volume of off-gas from the first stage, and rising ammonia consumption. 5 Comprehensive optimization and operational analysis of the medium-pressure absorption system: Through analysis, the actual content of CO2 present in the liquid phase of the first absorption cooler in the form of ammonium methoxide was determined, and the control methods for the horizontal balance in the first absorption tower were adjusted. This allowed the heat load entering the tower to be shifted forward, thereby improving the stability of the tower and creating conditions for high production with low consumption. (1) Determination of the total amount of CO2 present in the form of ammonium carbamate. Under normal conditions, 60% of the total CO2 in the feed stream can be absorbed before the first absorption cooler; thus, 40% of the total CO2 remains in the gas phase inside the first absorption cooler. If 65% of the total CO2 in the feed stream is absorbed before the first absorption cooler and exists in the form of ammonium carbamate, then the following reaction occurs: 2NH3 + CO2 → NH4COONH2. The weights involved are 34, 44, 78, 101.375 + 183.21; 51.987 + 237.095; 92.159 + 420.305. In this case, the amount of CO2 present in the liquid phase of the first absorption cooler in the form of ammonium carbamate is 298.081 kg/t·ur, which corresponds to 36.976%. According to the phase diagram, the melting point of this component is approximately 70–80°C (see Table 3), and this value lies at the edge of the methylammonium phase region in the phase diagram. It can be inferred that appropriately increasing the amount of liquid ammonia before the primary external cooler not only helps to shift the absorption load of gaseous CO2 forward, but also facilitates the component reactions to proceed into the ammonium formate phase region. http://pub2.hi2000.com/upload1/0712171541186605.jpg The melting point of the liquid phase component in the external cooler is determined to be 70–80°C, and the operating temperature can be set 10–20°C above this melting point. To this end, the temperature of the gas-liquid mixture exiting the cooler is controlled at 90–95°C. Once the melting point of the liquid phase component in the primary cooler is determined, the temperature of the soft water flowing in and out of the cooler should be kept above the melting point of the medium present in the cooler; therefore, the water temperature entering and leaving the primary cooler can be controlled at 70–80°C, which solves the problem of localized crystallization in the cooler. However, this requires that the external cooler to which the soft water is fed use a control method of \"high flow rate and small temperature difference\"; the flow rate increases from 2.0 m3/t·ur to 8.0 m3/t·ur, with the inlet water temperature ranging from 70–75°C and the outlet water temperature ranging from 80–85°C. By adjusting the water temperature to a suitable level, the temperature of the gas-liquid mixture in the external cooler is maintained at 90–95°C, which not only ensures an earlier onset of absorption load but also reduces the thermal load on the absorption tower. (2) New equilibrium for the amounts of ammonia water and dimethyl ether used in the single absorber. In traditional control of the level in a single absorber, the amount of water returned to the second stage is approximately: 260.697 kg/t·ur for dimethyl ether, and 196.101 kg/t·ur for ammonia water. Under high-load conditions, the CO2 conversion rate in the synthesis tower decreases, while the decomposition and absorption loads increase; simultaneously, the load on the secondary absorption stage also rises, which is manifested by an increase in the CO2 content in the two-cycle one-cooling (dimethyl ether) stream. Based on the analysis, the above situation can have its operating conditions improved by adjusting the amounts of dimethyl liquid and ammonia water fed into the absorption tower. The actual material composition under adjusted high-load conditions is shown in Table 4. http://pub2.hi2000.com/upload1/0712171541547873.jpg After adjustment, the ammonia reflux at the top of the absorption tower is controlled at 265 kg/t·ur (or 0.45 m3/t·ur), while no ammonia reflux occurs at the bottom. By achieving an ammonia concentration of 88–89% on tray 3#, the corresponding boiling temperature is around 50–55°C, while the temperature of the vapor stream exiting the first absorption tower is only 43–46°C. Due to the reduced water volume on tray 3#, only 265 kg of ammonia in top reflux is required to maintain the optimal temperature and concentration at various points in the washing section. Meanwhile, the decrease in the amount of ammonia solution creates the conditions for increasing the amount of dimethyl liquid used. The increased use of dimethyl ether leads to a reduction in the amount of the second-stage component, thereby stabilizing the operation of this stage. Meanwhile, the decrease in CO2 levels in the first cooling stage facilitates a shift in the load of the medium-pressure absorption tower. (3) In the new equilibrium of the methylammonium solution components within a absorption tower, under traditional practices aimed at maintaining stability, the temperature in the bubbling section of the absorption tower is often kept around 80–85°C; at this temperature, the CO2 concentration in the methylammonium solution is approximately 29–32%. Provided that system stability is maintained, the temperature in the bubbling section of the first absorption tower can be appropriately increased to 90–95°C, and operation at a higher liquid level can be employed; meanwhile, the amount of ammonia used for reflux can be reduced in a reasonable manner to increase the concentration of methylamine. At this point, the CO2 content in the methylamine is around 32–35%, which helps to reduce the H2O/CO2 ratio in the methylamine. This improves the conversion rate in the synthesis tower, thereby increasing production, reducing consumption, and achieving the goal of optimizing operations. By strictly following this approach to achieve innovation while maintaining stability and striving for excellence, ammonia consumption per ton of urea can be reduced by 2–3 kg, and the conversion rate in the synthesis tower can increase by about 1%, resulting in a cost reduction of 5–6 yuan per ton of urea.