Refining Report Canadian refiner finds simple route to reformulated gasoline production

North Atlantic Refining Ltd. (NARL) added this 7,500 b/d hydrogenation unit at its 105,000 b/sd refinery in Come By Chance, Newf., to treat a portion of the plant's reformate output. Since start-up of the unit in late 1995, it has achieved 100% benzene hydrogenation, thus allowing NARL to produce RFG containing less than 1.0 vol % benzene.

When the U.S. Environmental Protection Agency instituted reformulated gasoline (RFG) requirements as part of the Clean Air Act Amendments of 1990, NARL had to find a way to meet the specifications, even though the refinery is not within U.S. territory. The refiner chose to add a small hydrogenation unit to treat the portion of the reformate stream containing benzene precursors.

Since start-up of the unit in late 1995, it has achieved 100% benzene hydrogenation, thus allowing NARL to easily produce RFG containing less than 1.0 vol % benzene.

The problem

NARL's mid-Atlantic refinery location has tanker docks that can accommodate VLCCs, which reduces crude importation costs. But the refinery's location and product exports demand a processing scheme capable of meeting Canadian, European, and U.S. product specifications.

Fig. 1 [43853 bytes] shows that scheme, including the new hydrogenation unit.

The principal gasoline specifications changed as a result of the RFG regulations were oxygen content, vapor pressure, and benzene content.

A hydrocracking refinery produces no olefins, so NARL must import an oxygenate such as methyl tertiary butyl ether (MTBE) or ethanol. The refinery can meet the reduced Rvp specification by decreasing the amount of butane blended into gasoline. Fortunately, blending oxygenates permits a reduction in reformer severity and a consequent reduction in butane production.

Meeting the benzene specification (<1.0 vol %) is more difficult in a hydro cracking refinery than in one employing catalytic cracking. the reason is that hydrocracking refineries produce neither catalytic naphtha, which normally contains less than 1.0% benzene, nor alkylate, which has zero benzene content.

The benzene contents of NARL's principal gasoline streams before the RFG project are shown in Table 1 [11031 bytes].

In order to meet U.S. RFG specifications, some benzene had to be removed from the pool.

The solution

NARL evaluated a number of possible solutions to the problem of meeting RFG specifications, including extraction, isomerization, and hydrogenation.

Extraction was clearly not a viable option. The low benzene content of the streams requiring treatment would lead to high capital cost per unit extracted, and the benzene yield would be inadequate to justify marketing.

An earlier study proposed isomerization of a C6 fraction. That route certainly would reduce benzene to low concentrations, but the economic justification was a function of the octane gain. Because NARL would be importing oxygenates and reducing reformer severity, the value of the incremental octane was vanishingly low.

The high value of reformer hydrogen in a hydrocracking refinery further reduced any incentive to lower reformer charge rate or severity.In addition, a large part of the investment required for an isomerization unit would be spent on facilities to reduce contaminant levels in the light gasoline to the very low levels demanded by sensitive isomerization catalysts.

IONA Ltd., a consulting firm based in Sarasota, Fla., performed an analysis of NARL's potential crude slates and operations. The analysis demonstrated that removal of benzene from only the reformate stream would be adequate to comply with the benzene specification. Catalytic hydrogenation of the removed stream over a non-noble metal catalyst was investigated.

United Catalysts Inc. (UCI), Louisville, Ky., had previously supplied IONA with catalysts for the hydrogenation of solvents to exacting specifications. IONA asked UCI to propose a catalyst and reaction conditions for a hydrogenation unit that would meet NARL's requirements. Because the reformate stream had already been severely hydrotreated and was essentially free from potential catalyst poisons, as was the hydrogen source, a highly active catalyst could be used without fear of poisoning.

The processing scheme illustrated by Fig. 2 [24978 bytes] comprises the following elements:

• The straight-run naphtha splitter and hydrocracked naphtha splitter both would be operated to maximize rejection of benzene precursors into the light products while maintaining the benzene content of those streams at about 1.0 vol %.

• Full-range reformate would be fractionated into a light reformate with zero toluene content and a heavy reformate with a benzene content of 0.1-3.0%. The permissible benzene content would vary with the refinery crude slate and gasoline blending requirements.

• Light reformate would be hydrogenated to convert the benzene to cyclohexane. A residual benzene content of 0.5 % was targeted in the combined light and heavy reformate stream.

• The product would be stripped to remove residual hydrogen and C1s in solution.

Procurement

The imminence of the imposition of RFG specifications dictated that long delivery items, such as reactors and large fractionators, had to be excluded from the project.

Simulation of the operations of the two naphtha splitters showed that the existing configuration was adequate for processing heavy crudes, but not for light crudes, which have greater naphtha yields. IONA therefore prepared process designs and equipment specifications so that NARL would be prepared to immediately execute a revamp, should that become desirable.

No changes to the reformer were necessary, other than a configuration to permit uncooled reformate to be routed to the reformate splitter and to permit heavy reformate to be returned to the reformate cooler.

An idle deisopentanizer was pressed into service as the reformate splitter. The vessel was oversized, but advantage was taken of the 60 trays to reduce utility consumption, because this piece of equipment has the highest operating costs in the process. Surplus 650-psi steam capacity was utilized as a source of reboiler duty.

The hydrogenation reactor was designed for a fresh feed capacity of 7,500 b/d. Taking into account the high quantity of recycle to the reactor, its total processing capacity is about 16,000 b/d.

The heat exchangers were purchased, but were short-delivery items, as was the high-pressure separator. The reactor initially appeared to constitute the critical path. Fortunately, an adequately designed surplus reactor was located on site.

NARL had to completely redesign the inlet distributor system of the reactor. Phase conditions in the reactor vary from close to 100% liquid at the outlet at start-of-run to a high percentage of vapor at the inlet at end-of-run. An inlet distributor to accommodate this wide range of conditions required careful design.

Thermocouples were installed at various bed elevations and diameters to assist in recording the deactivation rate.

An idle on-site fractionator was available for use as the stripper. While the splitter was the oversized, this vessel was a poor fit for use as a stripper. Fortunately, the duty-stripping dissolved hydrogen-is not exacting and the process conditions could be compromised with minimum economic penalty. Nevertheless, it was necessary to retray the vessel.

High-liquid-capacity trays were used in the stripping section. The trays in the rectifying section were removed to accommodate a loading which was about 20% of the former service.

The pressure rating of the available vessel was less than would normally be desired for this service. A simulation of the operation in Newfoundland's climate with an oversized overhead condenser, however, predicted a low C4-C5 loss under anticipated operating conditions. And the economic penalty of butane loss to fuel was small because of the lowered Rvp specification of RFG.

Hydrogen supply

Hydrogen was available directly from the site hydrogen plant and was charged on pressure control to satisfy process requirements. A small vent of tail gas was taken from the high-pressure separator to the feed supply header of the hydrogen plant.

Hydrogen consumption equaled the stoichiometric demand (680 scf/bbl, as designed) plus the small solution loss (about 40 scf/bbl). The design required no additional compressors.

Catalyst

Aromatic hydrogenation can be done with either platinum or nickel-based catalysts. One might choose platinum-based catalyst if sulfur is preset. Because the hydrogen and light reformate supplies are essentially free of sulfur and HCl, poisoning was not an issue, and a nickel catalyst was chosen.

Nickel can effect complete saturation at a much lower operating temperature. It also costs less and is expected to last 5 years, depending on trace sulfur adsorption. Additionally, once the nickel catalyst is spent, it can be put into a sulfur guard unit, where it will pick up an extra 10+% sulfur. After the catalyst is fully spent, it can be sold as a nickel source for stainless steel production.

NARL chose to use UCI's C46-7-03 1/16-in. CDS catalyst, consisting of 52% nickel on silica. UCI prereduced and stabilized the product. Thus, the catalyst was safe to handle and had no unusual loading restraints, yet it required no activation by NARL.

The catalyst has been used in a wide variety of hydrogenation units treating light-to-heavy solvents, and in chemical applications. The design for the NARL benzene saturation unit was based on C6 solvent hydrogenation units that can reduce benzene to nil at low temperatures.

Special features

The process unit appears at first sight to be a conventional hydrotreating unit, but there is a complication. The reaction is highly exothermic and the net reactor heat release is greater than the amount of heat required to bring the charge to reactor inlet conditions. Tight control of the reaction was therefore essential.

Reactor effluent was recycled from the high-pressure separator to control the temperature increase across the reactor at about 150° F. at start-of-run, and at l00° F. at end-of-run. Refinery upsets, however, do occur.

Responsive instrumentation was provided such that, in the event that the reactor outlet temperature exceeded a limit, the hydrogen supply and liquid feed would be shut off instantaneously. In addition, a control valve from the high-pressure separator to the flare was provided to depressure the reactor circuit.

The existing fractionators were not in a convenient location. In fact, the reformate splitter was about 1,000 ft from the product stripper. Pipe sizes and schedules, and control valve locations, had to be chosen with care. The reactor system was located midway between the existing fractionators (Fig. 3).

Start-up

The unit could not be commissioned, as a typical hydrotreater is, by circulating hydrogen, introducing charge stock, and bringing the reactor to the anticipated process conditions. Control of the exotherm requires a supply of benzene-free recycle for start-up.

The reformate splitter was initially operated as a depentanizer with the reactor system operating on a C5/C6 mixture containing less than 3.0% benzene. Because introduction of hydrogen, and even low-aromatic feed, can be quite exothermic, the unit had to be brought on stream carefully.

Exotherms were controlled by simply cutting off the hydrogen and waiting for the temperature to decrease. Once conditions lined out, the splitter operation was adjusted to produce the normal reactor charge. Observation of the increase in the reactor temperature change (delta T) permitted this change to be monitored closely without burdensome laboratory analyses.

Operations

Since start-up, all operating objectives have been attained. There is no detectable by-product production, and hydrogen consumption is exactly as predicted. Start-of-run temperature was 200° F. It was later reduced to 180° F., which is the lowest temperature at which the unit can operate. The effluent benzene from the reactor is nil (Table 2 [61258 bytes]).

Lower inlet temperature appears to be achievable without penalty. During start-up, satisfactory performance seems to be attained at temperatures as low as 120° F. For normal operations, a higher inlet temperature is required so that the effluent temperature is high enough to reboil the stripper. Absent the stripping requirement, the highly active catalyst can meet product specifications at significantly lower temperatures.

End of run normally occurs when trace sulfur trapped on the nickel reaches 2-4 wt %. As deactivation occurs and the reaction exotherm moves through the bed, the temperatures will be increased to maintain conversion. When the maximum temperature in the reactor reaches 450° F., end-of-run is reached.

After 363 days of operation, the catalyst life was 27.31 bbl/lb and the reactor inlet temperature had not been increased. Additionally, the exotherm at the top of the bed had not moved down at all.

Some process upsets have occurred. One failure of the splitter controls permitted a high concentration of toluene to be charged to the reactor. The increased reactor delta T resulting from increased aromatics conversion immediately set off an alarm. The operators diagnosed the cause of the problem and immediately restored normal conditions.

In addition, an upset in the hydrogen plant permitted the carbon monoxide content of the hydrogen to increase to about 100 ppm. Catalyst activity declined, but was restored after a few hours of normal operation.

Table 2 presents a summary of the stream compositions before and after addition of the hydrogenation unit.

Process control

The fractionation cut points between the light straight run gasoline, the light hydrocracked gasoline, and the reformer charge normally are dictated by the overall refining strategy. The primary consideration usually is optimizing reformer operation and hydrogen production.

While benzene removal could be controlled by variation of the reactor inlet hydrogenation temperature, this would not be satisfactory. The time constraint of the system and the hysteresis of the reactor temperature profile would combine to impede precise control.

It is good practice to choose a reactor inlet temperature which results in 97-100% benzene hydrogenation, and to operate steadily for a long time at that temperature.

Control of the benzene content of the gasoline pool is best obtained by adjusting the reformate splitter so that the benzene content of the heavy reformate results in the gasoline pool just meeting the RFG benzene specification. This operation virtually ensures zero toluene in the reactor charge and, therefore, minimum octane loss by toluene hydrogenation. Such operation also permits the utility consumption of the reformate splitter to be minimized.

IONA specified a pressure-compensated temperature controller to ensure close control of the reformate splitter. Although this was not included in the initial design, it likely will be incorporated later, when the unit controls are transferred to the site distributed control system.

The Authors

Brian Walsh is a process engineer with North Atlantic Refining Ltd. (NARL). His experience includes engineering, unit design, design coordination from P&ID development, process hazards analysis, and commissioning. His current duties include contact engineering for the catalytic reformer at NARL. He has a BS in chemistry from Memorial University of Newfoundland and a BS in chemical engineering from the University of New Brunswick.

G.W.G. (Gerry) McDonald is president of IONA Ltd., Sarasota, Fla. He specializes in providing technical and economic services to refiners and petrochemical manufacturers worldwide. His recent activities have concentrated on fluid catalytic cracking and downstream process improvement through economic analyses, revamps, and computer control.
Prior to forming IONA, he was manufacturing vice-president of both Commonwealth Oil Refining Co. and Amerada Hess Corp., and an FCCU and alkylation specialist with UOP. He has a BS (honors) from the University of Glasgow, and is a registered professional engineer in several states. He is a member of AIChE and holds several process patents.

J. Douglas Perkins is a technical sales representative with United Catalysts Inc., Louisville, Ky. He has been in the catalysis business for 26 years, specializing in hydrogenation. He holds a BS degree in chemistry from the University of Georgia, and an MBA (honors) from the University of Houston.

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