Monitoring and treatment to maximise light cycle oil
Fouling control, simulation software and tower scanning can improve operations and reliability when maximising middle distillate and propylene production
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Current global trends in product demand favour the increasing production of diesel fuel over motor gasoline because the demand growth for middle distillate fuels has exceeded the demand growth for gasoline for some time.2 In fact, the US is now a net exporter of refined petroleum products (mostly middle distillates) for the first time since 1949.3 Furthermore, the demand for gasoline from refineries has declined due to increases in the average fuel economy and the use of renewable fuels.5
Decline in gasoline demand5 in the US resulted in a reduction in refinery capacity, particularly in PADD I (East Coast), which increased cost sensitivities throughout the US refining market. US demand for propylene continues to rise at a rapid pace because steam crackers are switching to ethane feeds, and propylene and diesel continue to command a price premium over most other refined products.8 These changes in product demand mix and market over-capacity have resulted in refiners focusing on increasing middle distillate production with minimal additional capital and operational costs.
The question many refiners are now facing is how can existing refinery assets be used to economically increase diesel and propylene production. Refiners have many options to increase middle distillate and propylene production. Unfortunately, many of the options have significant negative consequences or additional capital costs. Changes in operation are often constrained by existing equipment limitations, and changing the product mix can reveal new bottlenecks.2
Capital investment options
Crude distillation units and delayed coker units have some flexibility to increase distillate production, but are constrained by feed slate and mechanical limitations. Most US refineries only have two to five trays between the flash zone and the distillate draw, which results in less precise distillate cuts. In contrast, European refineries that are historically optimised for diesel production may have 10 to 14 trays in the same zone to ensure good separation. Some US refiners are now adding distillate draws on vacuum towers to maximise distillate production and limit the amount of distillate sent to the FCC unit, which is chiefly designed for the production of gasoline and will convert middle distillate-range feed to lighter products.2
Most US refiners have invested in a FCC unit as their main conversion technology. Changes in operation, catalyst and feed can optimise the FCC unit for distillate production. Increasing diesel in a FCC unit is challenging, in part because FCC light cycle oil (LCO) has limited value as a blend stock for diesel fuel due to its aromatic, sulphurous character and because it requires further hydrotreating before blending.6
Construction of new hydrocrackers, which provide increased flexibility over FCC units and better-quality gasoline and distillate products, is ongoing but slow due to relatively high capital costs and the resulting increased hydrogen demand. Tighter sulphur specifications along with the increased processing of higher-sulphur feeds and the rise in production of high-sulphur intermediate products such as LCO are increasing the demand on hydrodesulphurisation units (HDS).4 Moreover, increasing olefin production at the FCC unit increases the fouling potential for downstream HDS units, which can be controlled with appropriate antifoulant additives. HDS capacity is often a bottleneck for the refinery, particularly when optimising for distillate and propylene production. Therefore, there is a greater potential for HDS preheat fouling and a larger associated economic penalty.
Changing cut points
The most common strategy for increasing diesel production is to adjust cut points between the naphtha and middle distillate draws in the crude, FCC and coker units. Often, changes in the cut points are accomplished by reducing fractionator tower top temperatures. This reduction will increase ammonium chloride (NH4Cl) salting potential. Figure 1 shows a plot of ammonium chloride salt deposition isotherms from a typical FCC unit on the Gulf Coast, which was generated using Nalco Pathfinder simulation software. The plot demonstrates that the lower the temperature, the less ammonia and chloride are necessary to create salt. Similarly, Figure 2 shows a plot of overhead temperature versus salt formation temperature and salting potential from a typical Gulf Coast FCC unit, but with constant ammonia and chloride levels. This plot shows that a 6°F reduction in overhead temperatures can result in nearly three times the salting potential. Changes in tower pressure, overhead temperature or concentration of amines, ammonia and chloride alter partial pressures and therefore salt potential.
Salt deposits in towers and pumparounds can result in under-deposit corrosion and the generation of corrosion by-product deposits. These deposits may plug trays in the top and mid-sections of the tower, partially or completely block the heavy naphtha and middle distillate product draws, as well as severely limit the pumparound rates and heat removal from the tower. Performance of pumparound exchangers can be modelled using heat transfer simulation software. The build-up of salt deposits in trays can increase the ∆P across the top section of the tower. Uncontrolled fouling has resulted in tray flooding, evidenced by a rapid and significant increase in ∆P, which can have a negative impact on the fractionation efficiency between gasoline and middle distillate.1
The FCC unit and delayed coker main fractionator towers are particularly sensitive to salt fouling because of low temperature and relatively high concentration of ammonia (NH3) and hydrogen chloride (HCl) in the overhead system. An increase in ∆P at the main fractionator at both units can reduce wet gas compressor capacity and therefore limit the throughput. An increase in ∆P across the FCC unit’s main fractionator increases the pressure in the reactor and in the regenerator, which can limit air blower capacity and coke burn, as well as the production of propylene.1 Therefore, the control of chlorides into the distillation columns as well as frequent modelling of the salt potential and salt formation temperature are essential in adjusting distillation column cut points. Control of chlorides into the distillation columns can be accomplished through good desalting at the crude distillation unit and the injection of caustic downstream of the desalter. Optimisation of caustic and neutraliser injections through automation can reduce the chloride levels in the crude atmospheric columns and overhead systems while minimising the risk of downstream impacts such as sodium poisoning of the FCC catalyst.
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