Implementing and Sustaining Process Optimization Improvements on a Deisopentanizer Tower

Bobby Loe HOVENSA LLC

John Pults KBC Advanced Technologies

Presented at the AIChE Spring Meeting April 25, 2001 Houston, Texas

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ABSTRACT

Optimization of process units in a refinery is a difficult task. Product pricing and unit constraints often change weekly, but changing unit operating philosophy and addressing hardware constraints can take months to accomplish. Even after the steps to improving optimal performance have been identified and implemented, if the pressure to improve is removed, operation tends to return to the older, more comfortable routine, or constraints in other parts of the refinery prevent operation in the most profitable mode. This paper provides a case history of how operational improvements for a single deisopentanizer fractionation tower in a large refinery were identified, implemented and sustained. HOVENSA and KBC worked in partnership to assess the current operation, identify the optimum operating point using a process model, and implement the improved strategy in the field. Both real and perceived constraints appeared to block the way to sustaining the improved profitability, and removal of these obstacles required continual change and communication throughout the refinery. The improvements on this single tower have generated over $500,000 as compared with historical operation over the first six months. Introduction The HOVENSA LLC refinery in St. Croix is a joint venture between Hess Oil and PDVSA, and can process over 400,000 barrels of crude per day. The refinery has significant capacity in catalytic reforming, visbreaking, gas oil and diesel hydrotreating, isomerization, FCC and alkylation. Primary products include gasoline, diesel and kerosene, aromatics and fuel oil. Most units are pushed to capacity constraints. HOVENSA is essentially a merchant refiner, and usually enjoys open market volumes for both regular and premium gasoline. The Deisopentanizer (DIP) tower, shown in Figure 1, processes light straight-run naphtha from two sources. The LPG fractionation unit debutanizer bottoms consists of primarily iso and normal pentane (iC5 and nC5), with small fractions of butane and C6+ components. The overhead from the naphtha fractionator tower contains mostly C5 and C6 paraffin compounds, with some benzene and C6 naphthenes and a small amount of butane. The combined feed to the DIP is typically in the range of 9,000 to 15,000 barrels per day. The DIP overhead product is normally rich in isopentane, and is routed to gasoline blending along with other high octane, low RVP gasoline components. The DIP bottoms, which is rich in nC5 and C6 paraffins, is routed to the light naphtha isomerization unit, along with light raffinate from the aromatics extraction unit. The DIP tower was originally designed as a naphtha splitter, to separate light naphtha from catalytic reformer feed. The tower was converted to a deisopentanizer to remove iC5 from the isomerization unit feed, which increased this unit’s capacity for nC5 and C6

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paraffins and thereby increased the refinery’s ability to produce premium gasoline. During the revamp, additional condensing capacity was added for the DIP. However, the number of trays remained at 50 as in the original design, which is fewer than the 70+ trays that a typical deisopentanizer would have. As a consequence, the DIP at HOVENSA often has a difficult time making a good split between isopentane and normal pentane components. In December of 1999, HOVENSA and KBC agreed to a program to work together to identify and implement profit opportunities in the refinery. This program consisted of a six-month period in which KBC provided a high level of on-site support to work with HOVENSA staff and rapidly implement identified improvements. The contract included a shared risk element that based a significant portion of KBC’s fees for the project on implemented benefits, calculated on a monthly basis using prevailing economics. This provided both parties with incentives to quickly find and implement changes to improve profitability, and to maintain these improvements over time. Deisopentanizer operation was initially identified as a potential area for optimization. Four steps were used to realize the opportunity of improving DIP operation. First, an assessment was made of the current performance of the tower. Second, the potential for optimization of the DIP was analyzed using a rigorous process model and refinery economics. Third, implementation of the idea was performed in the field. Finally, the improvements were sustained by tracking progress and identifying obstacles to continued top performance. Assessment of the Deisopentanizer Operation An understanding of the current operating philosophy for the DIP was first developed, based on discussions with HOVENSA operations and technical staff. Historically, the tower had been operated with a target of 10% nC5 in the overhead, and 20-30% iC5 in the bottom product. These targets were based on some modeling work done in 1986 at the time of the revamp. The tower was reported to be limited at times by reboiler or condenser duty. One of the two steam reboilers had been out of service for some time, and 20-30% of the condenser fin fan motors were not operating and in need of repair. The DIP equipment had not been a maintenance priority, in part because no economic penalty had been calculated for having a reboiler or condenser out of service. There was also a concern that the tower could flood if both reboilers were placed in service. The DIP process control was accomplished with a DCS system equipped with an advanced process control (APC) algorithm. The controller was set to target 10% nC5 in the overhead and 10% iC5 in the product, and would increase reboiler steam and reflux until reaching a maximum for these flows. The tower pressure was also controlled within a specified range by the APC, and this could indirectly limit the reboiler duty as well, if the tower pressure increased beyond its set maximum. Inferential estimates for product iC5 and nC5 qualities were calculated based on tower temperatures and pressures, and a bias for these values was continually updated based on daily lab data.

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HOVENSA and KBC then collected a set of operating and laboratory data from the past year, eliminating periods of known equipment failure or poor unit volume balance. The data showed an average of 19% nC5 in the DIP overhead, and 27% iC5 in the bottoms product, with a wide variation in the product quality, as shown in Figure 2. An average of 18% of normal butane was observed in the overhead product, indicating poor debutanization in the upstream fractionation towers. The process design for the tower revamp showed an available reboiler duty of 51,000 pounds per hour of steam, but the average for the data collected showed only 27,000 pounds per hour. The reboiler duty was also erratic. Tower Optimization Using the averaged process and lab data, a model for the DIP was developed with HYSYS.Process software. The feed rate and composition to the tower were fixed, as well as the reflux rate, tower pressures and overhead rate. The model results for reboiler duty, tower temperatures and compositions compared favorably with the unit data, as shown in Table I. The calibrated model was used to simulate DIP performance at the design reboiler duty, to determine if available condenser duty and tower tray capacity would be adequate for this operation. As shown in Table II, the predicted condenser duty for this operation was only slightly above the design value, and tower tray parameters indicated that flooding was unlikely. Also the separation of iC5 and nC5 improved dramatically vs. the historical operation as would be expected with the increased tower traffic. Economic Evaluation The profitability of the DIP column was determined based on the value of separating iC5 for direct blending to gasoline and nC5 to be used as feed for the C5/C6 isomerization unit, less the utility and downstream unit opportunity costs incurred to do so. Lighter feed components, such as n-butane, were assumed to always be fractionated into the DIP overhead, and components heavier than nC5 were assumed to always be found in the DIP bottoms stream. Thus only the disposition of iC5 and nC5 components were considered in the profitability calculation. The formula used for calculating the DIP fractionation benefit is shown below: DIP Upgrade Value = Overhead Value + Bottoms Value – Feed Value – Reboiler Steam Cost – Isom Operating Cost – Isom Capacity Penalty The DIP feed and overhead values are calculated as the gasoline blending value of the iC5 and nC5 in this stream, with corrections for road octane and RVP of these components vs. those of conventional regular gasoline. The DIP bottoms stream is normally processed at the isomerization unit, where 75% of the exiting C5’s are assumed to be iC5. After this equilibrium conversion, the value of the resulting iC5/nC5 stream

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is calculated at gasoline blending value as described for feed and overhead above. The reboiler steam cost is calculated assuming a 70% generation efficiency from refinery fuel gas, and the isomerization unit operating cost (for fuel, power and catalyst) was taken to be the same value per barrel as used in the refinery LP. During some periods, the refinery isomerization unit has more feed available than can be processed. If additional DIP bottoms is produced, less capacity is available to process light raffinate from the aromatics extraction unit. The Isom capacity penalty, or the opportunity cost for processing additional DIP tower bottoms at this unit, was therefore estimated by evaluating the octane upgrade of light raffinate, using KBC’s proprietary isomerization model. Case Studies Once the economic evaluation method was determined, numerous case studies were completed using HYSYS.Process to determine the optimum operating point for the DIP tower under different scenarios. It quickly became apparent that in nearly all economic and operating situations, maximizing the DIP reboiler duty up to the design value gave the highest profitability. For subsequent case studies, the simulation was completed with maximum reboiler duty, and tower pressure and nC5 content of the overhead product were also fixed. These constraints completely specified the tower operating conditions. The DIP profitability was first examined for scenarios where the isomerization unit has available capacity. The DIP feed rate and overhead nC5 content were varied, and profitability calculated, as shown in Figure 3. This analysis showed that the optimum target was around 5% of nC5 in the overhead, regardless of the tower feed rate. Profitability was then examined assuming the isomerization unit was at its maximum charge rate, and additional production of DIP bottoms would result in bypassing of light raffinate around the isom, direct to gasoline blending. The cost of losing the light raffinate octane upgrade can vary between $2 and $5 per barrel, and so HYSYS cases were completed for both of these scenarios as shown in Figure 4. In these cases, the optimum nC5 in DIP overhead target is dependent on the charge rate to the tower. At low charge rates, the available reboiler duty is sufficient to obtain good separation between iC5 and nC5 components, so that minimal iC5 is lost into the DIP bottoms when targeting 5% nC5 in the overhead. At higher charge rates, more iC5 is lost to the bottoms stream, and it is more profitable to increase the overhead nC5 target, reducing the DIP bottoms rate to the isom unit and allowing additional raffinate upgrading. Thus the optimal nC5 in overhead target varies between 5% and 20%, depending on DIP charge rate and the value of light raffinate upgrading. Based on a comparison of the optimal tower operation as determined above, and the historical performance, an incentive of around $1.5 million per year was identified to improve DIP fractionation.

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Implementation of Changes In order to realize the benefit indicated by the optimization analysis above, several unit hardware, process control and operating philosophy changes were needed. First, it was clear that both reboilers would be required, so the spare bundle was leak-tested and returned to service by operations. Several fin fan motors were also quickly repaired to ensure that design condenser duty was available and overpressurization of the tower would not limit the reboiler duty that could be applied. Using the economic basis described above and current product prices, the Economics and Planning group began evaluating the optimum target for the DIP overhead nC5, and communicated this value to operations for input into the advanced process control system. The DMC controller on the DIP DCS system was reconfigured to operate at this nC5 target, while maximizing the reboiler duty as limited by the high limit on tower pressure. This system allowed the DIP operation to be maintained at an economic optimum, accounting for the isomerization unit capacity and the economics of the day. Communication of the new operating philosophy was also critical in improving DIP performance, and was accomplished by the refinery Technical Services and Operations staff. Operators and unit supervisors were trained on the importance of always maximizing the reboiler duty and setting the nC5 in overhead target. The iC5 content of the bottoms stream was still measured by lab and inferred analysis, but this was no longer a tower control variable. The reboiler duty and overhead nC5 were tracked on a daily basis, and performance for these Key Performance Indicators (KPI’s) were discussed at weekly operations and planning meetings. Performance Improvements The economic benefit of improving DIP fractionation was tracked on a monthly average basis, beginning in March, 2000 shortly after implementation of the new optimization strategy. Economic performance vs. the baseline operation is shown in Figure 5, using actual monthly averaged economics and unit operating and lab data. Monthly benefits of over $100,000 were achieved in several cases during the summer months, when octane values were at their highest level. As octane values dipped during spring and fall months, benefits from the improved DIP fractionation dropped off as well. A significant drop in benefits can be seen in Figure 5 for July. This was due to poor operation of the DIP tower, caused by a high butane content in the tower feed from the crude unit stabilizers. The C4’s caused the DIP tower pressure to increase up to its safe operating limit, and the reboiler duty was cut back to avoid overpressuring the tower. This resulted in a reduction in fractionation efficiency and profitability during part of July, and represented one of the challenges encountered in sustaining the improved DIP performance.

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Sustaining the Improvement Without ongoing attention, refinery profit improvements from initiatives such as that on the Deisopentanizer tower tend to fade over time, for a multitude of reasons. Challenges to the new level of performance must be tackled as they arise, whether they result from hardware or control problems, misunderstandings, or operating changes in other parts of the refinery. Tracking the economic benefits of the initiative is a critical element of sustaining the change, since knowing the lost profits associated with a loss in performance helps to set work priorities within the refinery. Several problems occurred during the first six months of improved DIP operation which reduced the profit derived from this tower, and these issues were quickly addressed to sustain the improvement. During the summer months, a change in routing of a portion of the refinery condensate resulted in a hydraulic constraint on the amount of condensate that could be removed from the DIP reboilers into the condensate header. This caused a reduction in reboiler duty and a loss in fractionation. The economic calculations clearly showed that the benefit derived from the additional reboiler duty was much higher than the value of recovering the condensate. For this reason, condensate was safely spilled into the sewer until normal condensate header operation was restored, and DIP fractionation was maintained. A second challenge occurred when it was noticed that the DIP reflux drum temperature had increased above 140oF, which was higher than the recommended run-down limit to tankage. Initially reboiler duty was decreased to ensure safe operation. However, in meeting with the tank farm operators about the problem, it was found that the DIP overhead mixed with several other much larger streams before entering a gasoline blending tank. Calculations showed that the effect of the higher DIP rundown temperature on the tank temperature was minimal, and DIP reboiler duty was again increased. Another problem in maintaining reboiler duty was identified when the DIP tower pressure increased due to butanes in the feed, as mentioned above. It was found, however, that the maximum tower pressure was set well below the vessel design operating pressure, and so the safe operating limit of the tower could be increased after an appropriate Management of Change (MOC) review. HYSYS modeling showed that operation at the higher pressure did not significantly impact the tower profitability if the reboiler duty could be maintained near the maximum. The pressure limit was increased and performance of the tower again improved. A fourth reduction in DIP profitability was noticed when the DMC controller began cutting reboiler duty, for no apparent reason. Further investigation showed that the inferred NC5 content of the overhead stream had been deviating from the lab value for a few days, because of a computer glitch. This was quickly corrected and DIP operation returned to normal.

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Although all of these obstacles reduced profitability for a short period, timely identification and resolution of the constraints averted potentially long periods of underperformance. Conclusion The refinery profit improvement process passes through several phases. First, current performance must be assessed, evaluating upstream and downstream constraints and unit equipment limitations. Understanding of how the unit operation can be optimized can then be gained by use of an appropriate process model and refinery economics, and a new operating strategy is then developed to improve profitability. Implementation of this strategy requires good communication of the plan and its benefits throughout the plant. Also needed may be upgrading or repair of unit hardware to allow operation at desired conditions, safety reviews and process control changes. Tracking of key operating parameters as well as economic performance of the unit and distributing these results within the plant is essential to sustaining the improvement, as this process flags deterioration in profitability and encourages quick resolution. The profit improvement process requires input from technical, operations, safety/environmental and economics groups within the refinery. In the case of the HOVENSA Deisopentanizer tower, this overall process of identifying, implementing and sustaining benefits led to over $500,000 of additional profits over the first six months.

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Figure 1: HOVENSA Deisopentanizer Flow Scheme

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Figure 2: Historical Deisopentanizer Performance

Normal Pentane in DIP Overhead

051015202530354045

Dec-98Feb-99

Apr-99Jun-99

Aug-99

N-P

enta

ne (w

t%)

Target: <10%

Isopentane Recovery from DIP Bottoms

01020304050607080

Dec-98Feb-99

Apr-99Jun-99

Aug-99

Isop

enta

ne (w

t%)

Target: 20-30%

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Figure 3: Optimization Without Isom Capacity Constraint

Deisopentanizer Optimization

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-4000

-3000

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-1000

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1000

0.00 5.00 10.00 15.00 20.00 25.00

NC5 in Overhead

15 kBPD

12 kBPD

9 kBPD

Max Reboiler, 55 psig, no Penex limit

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Figure 4: Optimization With Isom Capacity Constraint

Deisopentanizer Optimization

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0.00 5.00 10.00 15.00 20.00 25.00N C 5 in Overhead

15 kBPD12 kBPD9 kBPD

Max Reboiler, 55 psig, $2/bbl Penex Charge

Deisopentanizer Optimization

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0

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0.00 5.00 10.00 15.00 20.00 25.00N C 5 in Overhead

15 kBPD

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9 kBPD

Max Reboiler, 55 psig, $5/bbl Penex Charge

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Figure 6: Deisopentanizer Economic Improvement

Deisopentanizer Fractionation Improvement

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60,000

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Table 1: HYSYS Simulation Results vs. Deisopentanizer Data

30% 31% IC5 in Bottoms (wt%)

16% 18% NC5 in Overhead (wt%)

183 179 Bottom Temperature (F)

153 143 Top Temperature (F)

26.7 23.6 Reboiler Duty (MMBtu/hr)

Model Data Result

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Table 2: HYSYS Results at Design Reboiler Duty

18% NA IC5 in Bottoms (wt%)

5% NA NC5 in Overhead (wt%)

37% 50% Max Downcomer Backup

74% 85% Max Jet Flood

48.1 45.0 Condenser Duty (MMBtu/hr)

48.8 48.8 Reboiler Duty (MMBtu/hr)

Model Design Result

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