C₂ hydrogenation reaction in ethylene plants

Thread Content

This post was last edited by MTO001 on 2021-8-6 10:17. In the ethylene industry, the carbon dioxide hydrogenation unit plays a crucial role in the quality of ethylene products. In actual operation, the cycle of the carbon dioxide hydrogenation reaction and its control stability are among the key factors for the long-term stable operation of the plant. This article provides a comprehensive analysis of the carbon dioxide hydrogenation unit in steam cracking for ethylene production and methanol-to-olefins technology ； I. Methods for removing acetylene in ethylene plants: 1.1 Solvent absorption method. This method allows for the recovery of acetylene; it is rarely used in currently operating ethylene plants, although the Linde process at Jilin Petrochemical uses solvent absorption to recover some of the acetylene. Absorbents include dimethylformamide (DMF), propanone, and N-methylpyrrolidone (NMP). 1.2 The catalytic hydrogenation method can increase the production of olefins; all current ethylene production plants use this method. The main processes are represented as follows: Process package suppliers, hydrogenation technologies, representative domestic production facilities. Linde – pre-hydrogenation, pre-deethanization, pre-hydrogenation (isothermal reactor); Dushanzi Large Ethylene Plant – KBR; pre-hydrogenation, pre-depropanization, pre-hydrogenation – Lanzhou Petrochemical, Quanzhou Petrochemical; S&W – pre-hydrogenation, pre-depropanization, pre-hydrogenation; Zhejiang Petrochemical (there are many such facilities in China; only one representative is listed) – Lummus. Post-hydrogenation, sequential separation – Zhenhai Refining & Chemical Plant (the small ethylene plants built there use this type of process); Sinopec – pre-hydrogenation, pre-depropanization, pre-hydrogenation – Zhonghan Wuhan Ethylene Plant, S-MTO/DMTO, etc.; post-hydrogenation, pre-deethanization/propanization, post-hydrogenation – Zhongyuan Petrochemical, Shenhua Baotou, etc. II. Advantages and disadvantages of pre-hydrogenation (this article focuses on pre-hydrogenation process technology; currently, all ethylene produced via steam cracking uses the pre-hydrogenation process.) 2.1 Advantages of pre-hydrogenation 2.1.1 Simplified process, lower equipment investment, and reduced energy consumption ； 2.1.2 No external hydrogen supply is required, and no CO is needed to adjust catalyst selectivity ； 2.1.3、Low production volume of green oil, long catalyst lifespan ； 2.1.4 Hydrogenation at over 50% of MAPD to reduce the hydrogenation load on the MAPD hydrogenation system. 2.2 Disadvantages of pre-hydrogenation 2.2.1 When acetylene hydrogenation meets the requirements, MAPD hydrogenation fails to achieve the desired removal levels ； 2.2.2 The control methods are limited, and the temperature is prone to rising. 2.2.3 When heavy feedstocks are used for pyrolysis, the amount of butadiene entering increases, which affects the catalyst’s performance, leading to a shorter lifespan ； 2.2.4 Severe fluctuations in CO can easily lead to alkyne leakage or excessive temperature rise. III. Process flow and principle of pre-hydrogenation: 3.1.1 Linde process: The Dushanzi large ethylene plant uses pre-deethanization pre-hydrogenation technology (isothermal reactor, single reactor unit). The gas from the top of the deethanization tower flows into the reactor (where methanol is used for heat extraction), and then to the demethanization tower. In this reactor, the average conversion rate of acetylene to ethylene is 50%; there is no loss of ethylene, and the catalyst’s lifespan is 5–8 years. 3.1.2、KBR Process: The ethylene project at Quanzhou Petrochemical utilizes pre-propanation pre-hydrogenation technology (adiabatic bed, three beds in series); the conversion rate for MA is 64.7%, the conversion rate for PD is 40%, and the overall conversion rate for MAPD is 40.7%. 3.1.3 The S&W process: Zhejiang Petrochemical, Hengli Petrochemical, Yantai Wanhua, etc., use pre-propane removal followed by pre-hydrogenation technology (adiabatic bed, three beds in series): the conversion rate for MA is 85%, the conversion rate for PD is 50%, and the overall conversion rate for MAPD is 69%. 3.1.4 IFP Company’s technology involves the mixed-phase hydrogenation of pyrolysis gas; the materials exiting the gas and liquid phase dryers are combined and fed into the reactor. This technology is not used in China. 3.2 Principle of the pre-hydrogenation reaction (no further explanation is provided for the basic principles; please look them up on your own). 3.2.1 When the ethylene concentration is high, it is much more likely that acetylene will be converted into ethane rather than just ethylene ； 3.2.2 The rate of hydrogenation of ethylene to ethane is 10-100 times faster than that of hydrogenation of acetylene to ethylene ； 3.2.3 Acetylene adsorbed on the catalyst surface can be easily hydrogenated to form unsaturated hydrocarbons such as 1,3-butadiene. 3.2.4 CO can improve selectivity and inhibit catalyst activity. CO undergoes formylation reactions with H2 and unsaturated compounds (alkenes, alkynes, diolefins) to produce aldehydes and alcohols, or CO polymerizes with alkenes to form polyketones, thereby facilitating the formation of green oil. IV. Catalysts for Acetylene Hydrogenation 4.1 Catalyst Types 4.1.1 Non-palladium-based Catalysts These catalysts have low hydrogenation activity, require high reaction temperatures (120–260°C), exhibit poor selectivity, and result in significant losses of ethylene (1–3% mol). 4.1.2 Palladium-based catalysts: These catalysts exhibit high hydrogenation activity, operate at low reaction temperatures (40–90°C), offer good selectivity, and result in low ethylene loss (0.2–0.5% mol); in most cases, they lead to an increase in ethylene production. Adsorption capacity of palladium catalysts: Acetylene > CO > Butadiene > MA > PD > Ethylene, Propylene > CO2 4.2; Catalyst manufacturers: 4.2.1 Phillips Petroleum Company, which authorizes German company Südchem to produce G-83A/G-83C (catalysts containing Pd and Pd-coating/Al2O3, respectively) and C-36 (Ni-coating/Al2O3 catalyst); 4.2.2 Clariant catalysts, produced by German company Leuna: 7741, 7741B catalysts (Pd-based) ； German CRI KL-7741B-R type (type 1), a Pd-based catalyst supported on special Al2O3. 4.2.3 Sinopec Beihua Research Institute’s BC-H-20C and BC-H-21B; 4.2.4 Johnson Mathhey Company in the UK, palladium-rare earth bimetallic selective hydrogenation catalysts. 4.2.5 Shell catalysts; 4.2.6 Lanhua Institute catalysts. V. Factors affecting the acetylene hydrogenation process (pre-hydrogenation process). The optimal stable range for the acetylene hydrogenation reaction: it is the difference between the deacetylenation temperature in the acetylene hydrogenation reactor and the critical temperature at which selectivity drops by 1/3; this value represents the optimal stable range. The lower the CO content and the lower the hydrogenation temperature, the smaller the optimal safe area. Composition of green oil: Analysis shows that green oil consists of 90% aliphatic dienes, along with a mixture of C4-C20 unsaturated hydrocarbon products resulting from the reaction of olefins and alkanes, accounting for 10%. Boiling point: 120–400°C. 5.1、Hydrogenation reactor inlet temperature ； 5.2 CO concentration: CO can inhibit catalyst activity and improve catalyst selectivity; the formation of CO can be suppressed by adjusting the DMDS injection amount in the cracking furnace. 5.3 H2 concentration: It affects over-hydration; when the hydrogen content increases significantly, the catalyst’s selectivity declines, side reactions increase, losses of ethylene and propylene rise, and the risk of temperature spikes increases. 5.4 Acetylene concentration: 5.5 MAPD concentration: Affects the temperature rise in the bed. 5.6 Ethylene concentration: The higher the ethylene concentration, the greater the risk of loss. 5.7 Catalyst performance: Different catalyst manufacturers exhibit varying tolerance to CO, with significant variations in the allowable adjustment range as well as in the optimal stable operating range. 5.8 Air velocity: At low load or when the anti-surge valve is open, adjusting the anti-surge flow rate can limit the minimum air velocity; long residence time and excessive hydrogenation can lead to temperature spikes, as well as flow deviation ； Excessive air velocity, too short residence time, inadequate hydrogenation, and prone to acetylene leakage. 5.9 Quality of catalyst bed packing: The quality of the catalyst bed packing has a significant impact on the performance of the catalyst. For example, channeling can lead to excessive local resistance and an increase in the temperature of certain areas within the bed, which may result in excessively high temperatures in those specific points ； For example, if the amount of ceramic balls used is too large, the heat capacity is high, and there is significant lag in regulation ； Catalyst loading can cause crushing, which may lead to blockages in the heat exchanger, among other issues. 5.10, Butadiene: Excessive levels can cause catalyst coking and deactivation, as well as an increase in bed temperature and pressure difference. 5.11 Toxins: hydrogen sulfide, arsenic, carbonyl sulfide, Hg, etc., which cause inactivation due to poisoning. The requirement for thiols and thioethers in hydrogenation catalysts is generally less than 25 PPb (mol%). 5.12 Control schemes for the inlet of the hydrogenation reactor and between its sections: The control schemes vary depending on the specific process package, and the choice of such schemes affects the ease of operation as well as the effectiveness of automatic control. 5.13 Bed pressure difference: Reactor layout and coking blockage, etc. VI. Design and Operation of the Acetylene Hydrogenation Unit 6.1. Taking a 1 million tons/year ethylene plant as an example 6.1.1. Design operating parameters (the composition varies depending on the cracking feed; one feed ratio is used as an example): Component A: Inlet mol% = 0.51, Outlet mol% = 0.18; Conversion (%) = 1.62. Component B: Inlet mol% = 0.40, Outlet mol% = 0.29; Conversion (%) = 1.62, Total conversion (%) = 4.8. Component C: Inlet mol% = 0.46, Outlet mol% = 0.41; Conversion (%) = 4.26, Total conversion (%) = 4.7. Acetylene: Inlet concentration = 0.67%, Outlet concentration = 0.34%; ppm = 0.025; Conversion (%) = 99%, Total conversion (%) = 100%. Ethylene: Inlet concentration = 34.84%, Outlet concentration = 35.23%; Conversion (%) = 1.12, Total conversion (%) = 2.44%. Ethane: Inlet concentration = 8.37%, Outlet concentration = 8.57%; Conversion (%) = 2.39, Total conversion (%) = 5.26%. Hydrogen: Inlet concentration = 19.16%, Outlet concentration = 18.65%; Conversion (%) = 2.66, Total conversion (%) = 6.05%. Propylene: Inlet concentration = 12.18%, Outlet concentration = 12.38%; Conversion (%) = 1.64, Total conversion (%) = 4.51%. Propane: Inlet concentration = 1.42%, Outlet concentration = 1.44%; Conversion (%) = 1.41, Total conversion (%) = 3.52%. 6.1.2. Actual operating offline analysis data (different cracking feeds result in different compositions; one offline sample is used as an example). 6.1.3. Comparison between the designed pressure differences and actual operating pressure differences for various equipment in the hydrogenation reaction system. 6.1.4. Load distribution among the coolers between the protection beds: Cooler at the outlet of Bed A: 28%; Bypass of the cooler at the outlet of Bed A: 72%. Cooler at the outlet of Bed B: 28%; Bypass of the cooler at the outlet of Bed B: 72%. 6.2. Commissioning and daily operation of the acetylene hydrogenation reactor 6.2.1. Commissioning of the acetylene hydrogenation reactor Due to differences in catalyst performance among different manufacturers, as well as variations in their tolerance to CO and their stable operating ranges, only one domestic catalyst is used as an example here. Among the parameters provided by the process package design, it is required that the butadiene content in the feed to the acetylene hydrogenation reactor be less than 0.3% wt (0.4% mol); for more conservative operation, this value should be less than 0.1% wt. When domestic catalysts are used, the required level for butadiene is generally no higher than 500 ppm (V). (1) Before driving the system for the first time or when the catalyst activity in the reactor is high, CO must be injected into the reactor bed layers; it is necessary to inject it into each of the three bed layers in the reactor. After CO saturation, the reactor needs to be left for a period of time. If there is a pressure leak, the leak must be checked and then the pressure restored with nitrogen. If feed gas is introduced, the unit must be started immediately. Because of prolonged exposure, acetylene will coking on the surface of the catalyst. However, in actual operation, CO was not injected into the bed because the catalyst’s activity was not very high ； Some catalysts have high activity; 2000 ppm (v) of CO needs to be injected at startup. However, the catalysts used do not have very high tolerance for CO – when the CO concentration exceeds 1000 ppm (v), the catalyst’s activity drops significantly, making it easy to experience acetylene leakage. (2) Since the system is not operating at full capacity in the initial stage of commissioning, the anti-surge valve is kept at a certain opening degree, ensuring that the minimum flow rate remains above 340 t/h (the catalyst supplier requires a minimum flow rate of no less than 280 t/h). The feed rate – the higher the flow rate, the safer the startup process – needs to be increased gradually as the reactor is brought online, layer by layer. (3) When driving, start by raising the temperature in accordance with the requirements specified by the temperature rise chart, gradually increasing the temperature of bed A until the initial temperature (T1) is reached; at this temperature, a positive temperature difference appears. Continue to increase the inlet temperature to bring the adiabatic total temperature difference of bed A to around 10°C. Then, the inlet temperature of bed B is raised to T1 (i.e., the inlet temperature when bed 1 has a temperature increase), and it is then slowly increased to the initial temperature T2 to create a positive temperature difference. Increase the inlet temperature until the acetylene content at the outlet of the second bed layer reaches 100 ppm. Increase the inlet temperature of bed C further to T2 (i.e., the inlet temperature when bed B experiences a temperature rise), and then slowly raise the temperature to T3 at which a positive temperature difference appears. Increase the inlet temperature so that the residual acetylene content at the outlet is less than 1 ppm. (4) During actual operation, the three bed layers can simultaneously increase the inlet temperature; however, it is necessary to closely monitor changes in the bed layer temperatures. When a temperature difference arises between the bed layers, the rate of increase in the inlet temperature must be controlled to prevent excessive temperature rises. If the temperature rise of the bed increases too rapidly (it is possible to closely monitor the trend of this temperature rise as well as the changes in the highest temperature within the bed), there is a risk of overheating; therefore, the inlet temperature of the bed should be reduced promptly to prevent overheating-related interlocks from occurring and causing losses. (5) Once the acetylene content at the reactor outlet meets the specified standards, the acetylene conversion rates for each bed layer can be adjusted based on actual operating conditions – the designed conversion rates being 49%, 47%, and 3% – to 50–70%, 30–40%, and 5–10%, respectively. In the initial stage of operation, the temperature rise in each bed layer can be controlled at 12–14°C, 10–12°C, and 3–5°C, respectively. (The previous hydrogenation reactor generally requires 2–3 hours to achieve satisfactory hydrogenation results.) Note: According to relevant data, for every increase of 50 ppm in CO, the inlet temperature rises by 1.5°C to compensate for the loss in catalyst performance (this rule does not apply during the initial phase of operation). 6.2.2 Abnormality handling of the acetylene hydrogenation reactor 6.2.2.1 Emergency shutdown handling of the pyrolysis gas compressor When the pyrolysis gas compressor shuts down emergency, the reactor’s emergency isolation interlock is activated; the inlet and outlet isolation valves of the reactor close while the bypass valves open. At the same time, the reactor is depressurized. Once the pressure drops to a level suitable for nitrogen purging, large amounts of nitrogen are used to purge the reactor bed (at this point, attention should be paid to the flare system and the nitrogen pipeline system). After the nitrogen purging is completed, to save time when the reactor is put into use next time, the bed can be pressurized in advance using pure ethylene. 6.2.2.2 Handling of Abnormalities in Daily Reactor Operation (1) Acetylene leakage at the reactor outlet: a) Increased flow rate into the reactor – When the feed volume increases, it is necessary to raise the inlet temperature of the bed, based on the bed temperature and the changes in acetylene levels as measured through online analysis. Each increase of 0.2–0.5°C, or an increase of around 2°C, must be adjusted strictly based on the temperature rise in the bed layer. For better operability, it is possible to use the trend of the highest temperature in the bed layer as a reference. b. The reactor inlet temperature drops. Increase the inlet temperature, identify the causes affecting it, and find corresponding solutions. In actual production, the most common problems include control valve failures, improper setting of PID parameters in the control loop leading to poor tracking performance, significant fluctuations in steam pressure, and issues with the transportation of steam condensate. c. An increase in the CO content in the feed requires coordination with the cracking unit to increase the sulfur injection volume in order to gradually reduce CO production. When lowering the CO level, it is important to keep it at no more than 200 ppm per adjustment, to avoid issues such as alkyne gas leakage or excessive temperature rise due to too rapid a decrease in CO levels. If CO cannot be reduced, the bed temperature distribution can be optimized further based on the CO level. d. The carbon tetrahydrocarbon content in the feed is too high; this can be controlled by adjusting the high-pressure depropanizer in order to regulate the carbon tetrahydrocarbon level at the top of the tower. e. If there are toxic substances in the feed, analyze their source and reduce or eliminate them; if the levels of sulfur, arsenic, and mercury are too high, analyze the composition at the outlet of the arsenic-removal protection bed to prevent catalyst deactivation. The toxins in the pyrolysis gas generally originate from raw materials or off-gases returned to the ethylene plant from other units; therefore, it is necessary to regularly monitor the composition of these components and any changes in their concentrations. (2) Excessive hydrogenation leads to loss of olefins. The main purpose of the acetylene hydrogenation reactor is to remove acetylene from the pyrolysis gas through hydrogenation; this does not mean that the acetylene content needs to be reduced to 0 PPm. It is preferable for the pyrolysis gas to contain at least 0.1 PPm of acetylene after hydrogenation, otherwise the hydrogenation process would be too aggressive. The root cause of this situation may be excessive reactor temperature, reduced load, decreased CO content, or oxygen presence in the feedstock. In this case, if proper temperature adjustment is no longer possible, the reactor temperature must be reduced. Under certain conditions, the sulfur injection amount can also be reduced, thereby increasing the CO content in the feed. Reduce the reactor inlet temperature in steps of 2°C, at a rate of 0.2–0.5°C per step, thereby lowering the temperature of the reactor. If the acetylene content at the reactor outlet drops below the detectable range or the temperature at the outlet of the reactor bed rises rapidly, immediate action must be taken. If the selectivity of the hydrogenation process decreases significantly, the CO content in the feed material should at least be increased until the loss of ethylene becomes negligible. (3) Device lifting load: The load variation is within the range of 1000 Nm3, and the temperature must be decreased or increased slowly by approximately 0.5–2°C. If the load decreases and it is not possible to fully optimize the operating results by adjusting the temperature settings, then only one additional step is required to increase the carbon monoxide content. The lower the current or expected load, the lower the required temperature or the higher the CO content. If the feed flow experiences strong but brief fluctuations in the upstream or downstream sections, it is not advisable to change the reactor’s operating parameters, as this generally does not result in a significant amount of acetylene leakage.