Thread Content
The methanol-to-olefins process mainly consists of two steps. First, crude methanol is produced from natural gas, a process that has been industrialized ； Then methanol is converted into olefins, mainly ethylene and propylene. Different processes result in varying ratios of ethylene to propylene. UOP/Hydro’s Methanol-to-Olefins (MTO) process was developed based on Mobil’s Methanol-to-Gasoline (MTG) technology. This MTO process boasts great flexibility; it allows for adjustments to the production levels of ethylene and propylene by changing the operating conditions of the reactor, in response to changes in market demand. The ratio of ethylene to propylene produced in the product can be adjusted within the range of 0.77–1.33. 1 Catalyst developments: UOP/Hydro has made breakthrough progress by developing a new catalyst, MTO-100, based on the SAPO-34 catalyst. The SAPO-34 catalyst is a silicoaluminate phosphate molecular sieve that exhibits high selectivity for the conversion of methanol into ethylene and propylene. The new catalyst MTO—100 exhibits shape-selectivity, with its acidic sites and strength being controllable, **which improves the selectivity for conversion to ethylene and propylene, enabling a selectivity of up to 80% for these compounds. The SAPO series belongs to catalytic materials with high versatility. Although its thermal stability differs from that of zeolites, its chemical properties and crystal structure are very similar to those of zeolite materials, featuring uniform porosity, a crystalline molecular structure, adjustable acidity, shape-selective catalytic activity, and acid exchange capacity. Its greatest improvement lies in smaller pores, as well as controllable acidity levels and strength. Although the improved SAPO-34 is an ideal catalytic material for the MTO process, it is still not the best choice for fluidized-bed reactors. SAPO-34 must be combined with a series of specially selected binders. The choice of adhesive is extremely important; it must be able to enhance the catalyst’s activity without affecting its selectivity. The American company Nexant Chemical Systems believes that the use of treated silica and alumina as binders can achieve certain levels of porosity, acidity, and strength. The porosity of the binder is important; it must allow methanol and MTO products to enter and exit SAPO-34 rapidly. This catalyst is prepared in a similar manner to FCC catalysts, through spray drying. 2 Process Progress: The MTO process design of UOP/Hydro is similar to that of Mobil; since there are few heavy by-products that need to be separated and treated, the separation system is relatively simple. The raw material used in this process is crude methanol; therefore, there is no need to produce AA-grade methanol (with a purity of 99.85%) through distillation, which reduces the capital investment required for upstream methanol production facilities. However, crude methanol cannot be sold for other purposes, which limits the flexibility of methanol equipment. To maintain a stable temperature and yield more easily, the MTO process uses fluidized-bed reactors, with an operating temperature of 350–525°C and an operating pressure of 0.1–0.3 Mh. The severity of the MTO process can be controlled by yield, temperature, pressure, and catalyst circulation rate. Temperature determines the thermodynamic operation, while production capacity determines the contact time. Meanwhile, the conversion rate and selectivity change with pressure. The production performance of the MTO process of UOP/Hydro is shown in Table 1. Combining UOP/Hydro’s MTO process with the ATOFI-NA/UOP olefin cracking (OC) process provides greater flexibility. The OC process can utilize C4, a by-product of the MTO process, to convert it into ethylene and propylene (primarily propylene). A typical product balance is shown in Table 2. Combining the OC unit with the MTO unit can achieve a light hydrocarbon yield of 90%. This is very important for manufacturers who only need olefins and polyolefins, as well as for MTO plants in remote areas. Hydro now has a demonstration unit built at its production facility in Norway, which uses a fluidized bed reactor and a continuous fluidized bed regenerator. Since 1995, this demonstration unit has been operating periodically, and according to data provided by UOP, it has achieved a long-term methanol conversion rate of 99% along with stable product selectivity. To date, there have been no reports of large-scale industrial installations in operation. During the reaction process, carbon deposits that affect the catalytic efficiency are formed, and they must be removed by combustion. Using air as the combustion medium, the coking process takes place in a catalyst fluidized bed regenerator. During the reaction cycle, some catalyst particles break apart. These powders can be removed from the material exiting the fluidized bed using a suitable set of multi-stage cyclone separators. The exhaust gas discharged from the regenerator can have its heat recovered by the furnace, while the catalyst particles in the exhaust gas are removed through electrostatic precipitation. The mixture leaving the reactor passes through a specially designed feed/exit heat exchanger before entering the separator. Inside the separator, the vast majority of the water and unreacted methanol are removed. Hydrocarbons are separated from the oxygenated compounds stream through distillation, while the oxygenated compounds are removed in the compression section. A multi-stage compressor is used to liquefy hydrocarbon mixtures, removing residual catalyst powder from them. The hydrocarbon material containing 50% ethylene is sent to a separation system; since the material does not contain acetylene or other heavier components, the separation system is simpler than that of a steam cracking unit. Typically, separation systems include an ethane removal tower, a methane removal tower, a propane removal tower, and a butane removal tower, while the design of the cold box is simplified. Distillation towers produce polymeric-grade ethylene and propylene. Other products include fuel gas, light fuels containing ethane and LPG, C5, and C4 containing a small amount of 1,3-butadiene. Although a small amount of acetylene exists theoretically, a deacetylene reactor was not included in the original design. Because small amounts of acetylene, methylacetylene, and methylenepropadiene (MAPD) are within the allowable limits for polymeric-grade ethylene and propylene. If the levels of acetylene and MAPD exceed the upper limits, selective hydrogenation is required; if the hydrogen production in the MTO unit is low, a certain amount of hydrogen must be supplied. 3 Economic evaluation Below is a comparison of the costs of producing ethylene using the MTO process in the United States and the Middle East. The natural gas prices in these two regions represent two extremes. In the Middle East, it is 0.75 dollars per pound, while in the United States it is 5.31 dollars per pound. Raw material costs include the cost of large-scale methanol production and a 10% return on investment. 3.1 United States: In 2003, the production cost of methanol in the United States, plus a 10% profit margin, was 223 dollars per ton, while the market price of methanol was 227 dollars per ton. The cash cost of producing ethylene using UOP/Hydro’s MTO process, including the cost of methanol plus a 10% profit margin, amounts to $850 per ton. The cost plus a 10% return on investment amounts to approximately $1,050 per ton, which is significantly higher than the average U.S. ethylene contract price in 2003 of $628 per ton. Table 3 compares the MTO process with the steam cracking process. The cost of raw materials alone is $100/ton higher than the average U.S. ethylene contract price in 2003, indicating that the MTO process for producing ethylene is not feasible in the United States. Natural gas prices are expected to drop significantly in 2013, but they still will not reduce production costs markedly (see Table 4). If the price of natural gas drops to $2 per pound, the MTO route for producing ethylene becomes viable, yielding an 10% return on investment. This indicates that the feasibility of the MTO process for producing ethylene depends to a large extent on the price of natural gas. 3.2 Middle East: In 2003, the price of natural gas was 0.75 dollars per pound. The cash cost of producing ethylene using the UOP/Hydro MTO process was approximately 62 dollars per ton; adding a 10% return on investment brought the total cost to 280 dollars per ton. Table 5 compares the MTO process and the steam cracking process. Compared with the traditional steam cracking process, the total cost of producing propylene using the MTO process includes cash costs, depreciation, and a 10% return on investment. Due to the lower capital requirements of the MTO process, it has lower costs and higher profits compared to condensate cracking units. Middle Eastern natural gas has price advantages, making it feasible to produce ethylene using MTO technology in the Middle East. It is estimated that by 2013, the production costs of ethylene using MTO technology, as well as those of ethylene produced via traditional steam cracking using ethane or condensate as feedstocks, are shown in Table 6. As many large companies are looking for investment opportunities, natural gas prices are expected to rise to $1 per pound by 2013, and it will no longer be as attractive as it was in 2003. However, profits can still be obtained