Distillate hydrotreaters are large energy consumers. As more stringent sulfur specifications are introduced, refiners must increase energy usage for hydrotreating purposes. As much as 10% of a refinerys total energy consumption is attributed to hydrotreating products. Furthermore, these units can be quite energy inefficient. This case study examines the economics of retrofitting distillate hydrotreaters options that can improve energy efficiency: 1) increasing the surface area of feed preheat exchangers, and 2) installing a hot separator. The study also addresses the economics of using alternative heat exchange techniques, such as twisted tubes, plate exchangers and printed circuit exchangers.
Distillate hydrotreaters are one of the main energy consumers within a refinery. New environmental rules mandate even lower sulfur levels for transportation fuels. Thus, refiners have few options but to increase hydrotreating severity to remove sulfur-containing compounds from refined products and product streams. Unfortunately, many existing hydrotreating operations waste energy.
Room for improvement. The energy efficiency of a process unit, or of an entire refinery, can be benchmarked and compared against a selected reference value. Various methods are used: some are statistical; some are based on historical performance; and others apply a more solid engineering basis.
Some options apply a best technology (BT) benchmark, whereby the energy efficiency of an existing unit is compared with the best design target.1 The ratio of the actual and the targeted consumption yields a BT index for the process unit.
When benchmarked against such targets, many hydrotreaters show BT indices of 500%. In other words, these units consume five times more energy than that of the best design unit. The BT unit used for comparison would be one with the same throughput, feed quality and severity of operation as the actual unit.
Challenged energy performance.
The poor energy efficiency of an average hydrotreater is primarily due to insufficient heat integration, resulting in the loss of high-grade waste heat to water and air coolers. The effectiveness of heat integration is normally assessed by using pinch analysis techniques and its targeting part.2
Fig. 1 shows a conventionally designed diesel hydrotreater. In this particular case, the BT index was estimated at 550%. Part of this inefficiency is immediately attributed to the units electric power consumption, which is generated at suboptimal total efficiency (24%). If an assumption is made that all motive power is generated at the BT efficiency (80%), so that the effect of using inefficiently generated power is annulled, the remaining BT index would be 450%.
This unit is heat-integrated to some extent: the reactor effluent preheats the gasoil (GO) feed in exchanger E-1, before preheating the stripper feed in exchanger E-3. The unit is hot fed at 160°C. If the unit were cold fed, there would normally be another feed/effluent exchanger downstream of E-3 and before the air cooler.
In many existing hydrotreaters, the feed preheat temperature is considered low, bearing in mind that it is preheated using very hot reactor effluent. In the unit as shown in Fig. 1, the feed is preheated to 313°C, although the reactor effluent is available at 385°C. Surely, the approach between the two temperatures could have been engineered to be lower, say to 40°C, instead of the actual 72°C.
Obviously, if the sizes of exchangers E-1 and E-3 are increased, more heat could be transferred between the hot stream (effluent) and the cold stream (feed). This would increase the feed-preheat temperature and reduce the furnace duty, at a constant furnace outlet temperature.
The other major inefficiency is the absence of a hot separator. The effluent is condensed in the air cooler upstream of the high-pressure and low-temperature (HPLT) gas/liquid separator. This heatstill available at a high temperature of 172°C, and quite valuableis lost to the atmosphere. It would be more energy efficient to partly condense the effluent at 172°C in a hot separator, send the hot liquid directly to the stripper, only cool the remaining gas against air and send the smaller stream to the existing cold separator.
The problem of improving the energy efficiency in a grassroots distillate hydrotreater was addressed elsewhere.3 It is proposed that a hot separator, combined with enhanced heat transfer between the feed and effluent, can lower unit energy consumption to practically a BT value. At such conditions, the feed would be at a temperature high enough to be sent directly to the reactor. The furnace would not be used in normal operation, as all remaining heat would be supplied by the reaction exotherm. The proposed enhanced heat transfer is to be accomplished by using plate-type heat exchangers. Such a solution is feasible, and a number of units have been designed, built and operated in this way.
Retrofits prove more difficult. However, in retrofit situations, which are primarily considered in the present analysis, it is observed that many hydrotreaters are too small and are below the critical size for plate-type exchangers to be retrofitted economically. The problem is further aggravated if the unit is hot fed, which is otherwise a desirable feature of energy efficient designs.
With all this in mind, there are two options, not mutually exclusive, for improving the energy efficiency of an existing, medium-sized hydrotreater:
1. Add an exchanger area to E-1 and E-3 so that more reactor-effluent heat can be recovered by the GO feed
2. Install a hot separator and modify the exchanger network, as required.
As may be intuitively expected, the first option is likely to be less capital intensive, but offering lower energy savings. The presented study addresses the economics for both options, and provides conclusions that may direct selection.
Option 1: Increase area of E-1 and E-3.
Both exchangers E-1 and E-3 have large temperature approaches (72°C and 60°C, respectively) and offer the opportunity to economically install additional surface area. Table 1 lists four projects to be evaluated.
Based purely on its simple payback, Project Ainstalling twisted tubesappears to be the most attractive. This project could on average save 5.2 GJ/h (from start-of-run to end-of-run). Twisted tubes are easy to install. They may cost a fraction more than a new shell, but they allow duty increase without increasing systems pressure drop, thus keeping the recycle compressors power unchanged. Although the shortest of the four options listed in Table 1, the payback for Project A is not very attractive at 4.5 yearsa return that may not be justifiable unless other benefits may result.
The cost to add the first new shell (Project B) is high due to extensive piping modifications. However, once the first shell is installed, the incremental cost of adding more exchanger area is reduced. Following this logic, Projects C and D have been consideredadding 2 and 3 new shells, respectively. At 4.6 years, Project C shows a slightly improved payback time than Project B. The payback on the incremental area is 4.3 years.
As expected, each additional new shell recovers less heat, as the temperature driving forces in the exchanger are reduced. Project D (a third new shell in each exchanger) only recovers an additional 1.3 GJ/h, while the savings are mostly outweighed by the increase in compressor duty. Project D, alone, has an incremental payback of 14 years.
The fact that these modified exchangers show a temperature cross, and have relatively large duties with (now) tighter temperature approaches, indicates a possible application of plate-type or similar exchanger designs.
Option 2: Install hot HP separator.
An alternative approach to saving energy is installing a second separator. The bulk of the liquid downstream of exchanger E-1 would be sent directly to the stripper, rather than being cooled to 40°C and then reheated in the reactor effluent/stripper feed exchanger E-3.
In the example hydrotreater (Fig. 1), this option is particularly attractive as the pinch occurs precisely in exchanger E-3. A hot separator eliminates the need for this exchanger, thus larger energy savings are possible than what is achievable by adding surface area alone (as in Option 1).
Fig. 1. Flow diagram of a conventional distillate
Since E-3 becomes redundant, it can be reused as an additional feed/effluent exchanger. There are no other changes in the exchanger network. The stripper bottom stream continues to preheat the stripper feed from the cold separator and the treat gas. If the unit were cold fed, there would be an option to use more of the stripper bottoms heat and perhaps the gas stream from the hot separator to preheat the cold feed. The proposed revised flowsheet is shown in Fig. 2.
Fig. 2. Flow diagram of the original and retrofit for distillate
Table 2 summarizes the economics of two available hot separator options. The first option includes installing a new separator, and reusing E-3 as a further reactor feed/effluent exchanger (this includes replacing the shells for an increased pressure rating). The second option considers the incremental benefit of adding more area to E-3 and adding a recycle-gas heater to recover additional energy.
The investment cost associated with installing a hot separator is higher than the cost of simply adding exchanger area to the feed-preheat train. However, the savings are larger and the return on investment is improved.
The second option shown in Table 2adding more area to E-3offers an attractive incremental payback of 2.4 years. Once the decision is made to install a hot separator, it may be more cost-effective to increase the size of the feed/effluent exchangers at the
Due to the higher separation temperature and increase in hydrogen solubility, installing a hot separator incurs two important process-related consequences:
1. Reduced hydrogen content in the recycle gas. A 5% reduction in hydrogen (H2) concentration is expected for the example hydrotreater. This will shorten the catalyst life and reduce the cycle length from about 3 years to 2.2 years. Assuming a catalyst volume of 120 tons, at 16 /kg, the extra catalyst replacement cost would be around 230,000/yr. Alternatively, the catalyst life could be restored to three years by increasing the purge and H2 makeup. About 1,500 Nm3 of additional H2 will be needed for each 100 m3 of feed. Which of the two alternatives will be selected depends on the hydrogen cost. Shorter cycle length would also incur a production loss of 12 days/yr and some added maintenance costs.
2. Increased H2 loss to fuel gas. This additional loss is estimated at 175 kg/h. In this particular case, this additional loss is estimated at 175 kg/h, losing around 275,000/yr as a difference between the cost of H2 (900/ton) and its values as fuel (6/GJ). This loss is, however, much reduced if the purge gas is recycled to the H2 manufacturing plant.
The additional processing cost reduces the benefits of installing a hot separator from 885,000/yr to a value between 375,000/yr and perhaps 550,000/yr, depending on the purge gas routing. This net benefit can be lower or slightly higher than the 490,000/yr achievable by revamping the preheat train only. The extra processing cost renders the hot separation unattractive in this particular retrofit. Similar results have been reported elsewhere.4 The conclusion, however, may change if the effects of unit debottlenecking and/or throughput increase become substantial. Those benefits would be greater with the hot separator than with just revamping the feed preheat train, and may again swing the project economics in hot separators favor.
Minor processing issues.
Other, minor process issues to be addressed are:
Wash-water and stripper operation. In the present study, it was possible to maintain the wash-water consumption and stripper conditions at present values, so that there would be little or no change to the downstream operation. This needs to be verified for each particular case.
Additional pressure drop from sending the feed through E-1, although any ∆P increase would be, to a large extent, offset by the lower flowrate through the cooler, upstream of the cold separator.
A process study may also address moving exchanger E-3 upstream of the feed pump, to avoid the need to increase the pressure ratingincluding the effect of high temperature on pump operation and cavitation.
Control outlet temperature.
As the furnace duty is reduced by the listed revamp projects, a question may arise concerning control of the reactor outlet temperature (which is affected by the feed temperature and reaction exotherm).
Improved heat integration greatly reduces the duty of the feed furnace, in some cases to zero. It can be argued that an important temperature controlling mechanism is removed. With the feed heater in operation, the reactor outlet temperature can be controlled simply by turning down the heater and reducing the feed temperature.
However, with the furnace on minimum firing or on standby, other controlling mechanisms can be used. These include a quench-gas flow to the reactor bed, bypassing the feed/effluent exchanger, using a feed cooler or installing an additional heat consumer (e.g., a steam generator) in the reactor effluent loop.
To enhance the heat transfer in a large shell-and-tube heat exchanger with tight temperature approach and a significant temperature cross, large additional area must be installed. The materials of construction and pressures involved add to the costs for such revamps. To lower this cost, alternative exchanger technologies can be considered. Such options include:
Plate/frame. The plate-and-frame type of exchangers offer high heat-transfer coefficients, and are fairly straightforward to mechanically clean. However, the pressure requirements of hydrotreating units may make some of these exchangers unsuitable in this application. Plate exchangers have been proven in large grassroots designs. For the particular unit considered here, the duty was too small to justify replacement.
Printed-circuit heat exchangers (PCHE). These exchangers use diffusion-bonded stainless steel plates, with channels for the fluid etched in them. The channel geometry leads to high heat-transfer coefficients, while the design is suitable for high pressures. PCHEs are compact, can be produced in small sizes and offer several advantages:
Fits in plot area of existing exchangers
Tighter temperature approach can be economically achieved
Lower overall pressure drop
Improved project economics. In the present study, results found are:
º Option 1adding exchanger area. Slightly higher benefits are obtained (670,000/yr vs. 537,000/yr, due to the lower pressure drop of PCHE), at 40% lower investment (2.1 million vs. 2.9 million), offering a 3 year payback (vs. 5.5 years)
º Option 2using PCHE in conjunction with hot separator. Again, slightly higher benefit (930,000/yr vs. 885,000/yr) are obtained, but at lower investment cost (2.8 million vs. 3.3 million), and faster payback (3 yr vs. 3.7 yr)
Potential disadvantages of using PCHE are:
Mechanical cleaning is not possible (although materials of construction and low liquid volume are suited for chemical cleaningsimilar to plate exchangers)
Small passages (slightly smaller than in plate exchangers) raising fouling issues
Insufficient experience. PCHE have many offshore applications, but are rarely used in downstream industry.
Of the two options available for revamping a hot-fed hydrotreater, 1) adding area to feed/effluent exchanger, and 2) installing a hot separator with some addition of heat exchange area, the hot separator option (with exchanger area addition), offers larger energy savings (13 GJ/h vs. 7.9 GJ/h of furnace process duty), and larger potential energy benefits (885,000/yr vs. 490,000/yr). Hot separator requires a higher investment cost (3.3 million vs. 2.3 million), but offers a more favorable return (3.7 yr vs. 4.6 yr).
With a hot separator, the unit energy efficiency, as measured by the BT index, would improve from 450% BT to around 240% BT. Installing a hot separator incurs additional processing cost due to reduced H2 concentration in the recycle gas. This affects the catalyst life and cycle length, and increases H2 loss to fuel gas. These factors substantially impact project economics; they can render the hot separation option economically unattractive in retrofit situations. In grassroots designs, however, it is expected that the hot-separator configuration would be chosen.
The analysis and the project economics are based on energy benefits only. In many cases, improved heat recovery de-bottlenecks the feed furnace and enables an increase in unit throughput. Both revamp options are open to this additional potential benefit. If the unit capacity can be increased, the refiner may find that the yield-related benefits outweigh the energy ones, and they more than compensate for the H2 purity loss in case of the hot-separator installation. HP
The authors thank Dharmesh Panchal and Joris Mertens for their valuable suggestions during the preparation of this article.
1Milosevic Z. and W. Cowart, Refinery energy efficiency and environmental goals, Petroleum Technology Quarterly, Summer 2002.
2 For an introduction to Pinch Analysis see: Hans-Joachim Leimkühler Managing CO2 emissions in the chemical industry, published by Wiley-VCH, 2010. A more detailed account is to be found in: Kemp I. C., Pinch analysis and process integration. A user guide on process integration for the Efficient Use of Energy, Elsevier, 2007.
3 Barnes, et al, HDS benefits from plate heat exchangers, Petroleum Technology Quarterly, Spring 2004.
4 Mertens, J., Rising to the CO2 challengePart 3. CO2 Emission reduction options in refineries, Hydrocarbon Engineer, March 2010.
|The authors |
||Dr. Zoran Milosevic is a senior staff consultant with KBC Process Technology Ltd., and an internationally renowned authority on energy optimization and profit improvement of oil refineries and petrochemical plants. He is best known through his work on profit improvement and energy conservation. He has over 40 published papers and articles on energy efficiency, refinery/petrochemicals profitability improvement, and energy economics. Dr. Milosevic has given numerous training courses in energy economics, refinery energy efficiency and Pinch technology. |
||Tim Shire is a senior consultant with KBC Process Technology Ltd, working in the Energy Optimization group. He manages energy and water optimization studies in refining, petrochemicals and gas processing industries in the Far East. Mr. Shire holds BS and MS degrees in chemical engineering from the University of Cambridge. |