Controlling and Optimizing Integrated Fractionation Processes in a Petroleum Refinery

Copvright © IL\C C0l1t1'01 of Di stil latiol1 Colullllls and Chemica l Reactors. Bournemouth . L' K I ~lx6

CONTROLLING AND OPTIMIZING INTEGRATED FRACTIONATION PROCESSES IN A PETROLEUM REFINERY M. Sourander* and S. Gros** *PrucrS,1 AlI/ullla/ion, Neste Technolugy , Porvoo, Finland ** Prucess Teclllwlug)', Nes/e Technology, Porvuu , Finland

Abstract In this paper comp ut er contro l of integrated fraction atio n processes in a petroleum refinery is di sc ussed . The impl ementat ion of on -l ine optimization of product yie lds versus energy usage i s s hown. The combination of on -l i ne eco nom i c closed l oop descissio n making with constraint controls, re prese nt ing system l imitatio ns in f r actionators and heat exhangers are also illustrated. A compariso n of the sit uation before and af t er t he computer contro l impl eme ntatio n is s hown in terms of obta ined co ntr o l qu a lit y as well as yie ld improvements. Keywords .

Computer control On - line operation Optimisation Decoup ling Models Oil refining

UNIT CONFIGURATION OF THE NAANTAL I REF INERY

Valmet's Damatic systems , each of whi ch is also co nn ected to t he process control computer . The pr ocess comput er system wa s delivered by Nokia oy from Finland and cons ist s of Digital Equipment Corporation's VAX 11 / 750 computer usin g VMS 4 . 2 operat ing system, No kia PMS process management system so ftw are, instrumentation interface and M/ M communi ca ti on hardwa re and software.

A brief int roduct i on of the Naa nt al i refinery is needed before a detailed discussio n about optimization strategies can be presented . The refinery has two a t mospheric crude units . The long residue from these are further fractionated in a vacuum unit.The vacuum short residue is termally cracked in a visbreaker unit to obtai n heavy fuel oil. The lig ht and heavy vacuum gasoi l is the main feed fo r the catal ytic cracker unit besides the heavy gasoil ( HGO ) from th e crude units. The light and medium atmospher i c gasoi l s ( LGo and GO ) a l ong wit h the correspondin g fractions from t he catalytic crac ker are hydrotreated to form th e diesel and light gasoil pools. Kerosine i s treated a nd used as jet fuel or used f or blending. The he avy napht ha i s hydr otreate d and reformed to obtain motor gasoline while the light nap hth a is s tabili zed in debu tanisers and further treated in naphtha splitter, deisopentanizer and depentanizer co lumns. It may be empha sized t hat the product cooling, pumparound a nd feed pre heat systems of crude unit 1, crude unit 2 and the vacuum unit form a comple x integrated heat exchanger network.

During the mod e rni za tion all co ntro l function s of the refinery were recon str ucted to s upp ort the r efinery pr ofi t maximization objective . Th is was done by the Neste Technology Group in c lose co-operat ion with th e operat ing personnel of t he Naantali Refinery. As a general rule it was esta bli s hed that all basic contro l s, th at are needed to run the process unit s without excessive manual operations have to be implemented in the digit a l ins trument at ion systems . Because th e instrumentation systems are all redundant this impro ves the security. On the other hand, control functions that require l arge amount s of su pport i ng ca l c ul ations , optimization or specia l co ntr ol algoritms were reali zed i n the computer . To plan, implement and commission the the comp ut e r controls 18 ma nyears of work were required. In th is paper no distinction is made between controls in the instrume ntation system and in the computer because of their functi ona l simila rit y. Similar control projects, as the one discussed in thi s paper, have been anal yzed erlier by (Nas i, 1983; Rinne, 1982; Sourander, 1984 ) .

In the computer control system implemented at the Naanta l i Refinery 14 f ractio nation units were involved. In this paper, however, the attentio n is focused mainly on the two cr ude unit s , the vacuum unit and the light napht ha spl itt e r system .

Co ntr ol objectives Instr umenta ti on The economic performance of a given petroleum r efi ner y is governed by interactions bewteen yield and produ ct quality infractio nation un i t s, by interaction between feed quality, process conditio ns and product qualit y in co nve r sio n units and by blending characteristics in sales

During the modernization period 1984 - 1985 the Naa ntali refinery anal og i nstrume ntation was complete l y replaced by digital instrumentation. The manufacturer is Valmet oy from Finland. The digital inst rumentat i on consists of four of 61

\1. L. Sourandcr and S. (;ros product finishing. This means d1fflculties in the control mechanisms but also possibilities for additional benefits from interunit manipulations. Hence also a refinery wide, closed loop opt1mi zation system, developed within ~este OY, was installed in the Naanta!i computer , but is not described in this paper. The control objectives for the refinery section whi ch is described here are as follows: 1.

Maximize recovery ( or yield ) of every distillate product aga1nst its most critical quality property

2.

Allocate energy optimally to emphasis separation vs. energy interaction

3.

Operate within the equipment limitations and drive against the constraint which represents the biggest profit

4.

Maximize control responsIveness by use of proper control structure, implicit decoupling and dynamic element

5.

Retain control robustness to emphasis the predicted process condition variations

CONTROL STRATEGY Product quality control In fractionation of petroleum, different column products represent distinctive cuts of the feedstock which are measured by ( standard laboratory testing) distillation curves. The split between two adjacent cuts is determined by the cutpoint between them. Th e cutpoints can be calculated by using co l umn direct measurements and they are used for every petroleum fraction as a primary quality property to control the quality of the lighter product heavy end. The quality control consists of yield and balancing control, cutpoint control and analyzed property control as cascade type logic as shown in the Fig. 2. The sidedraw product yields ( in percentage of column feed ) are stabIlized using dynamically compensated yield contro l lers and feed forward controllers to decouple the propagation of dIsturbances further than to the adjacent product below. Cutpoints are controlled to targets by manipulating the yield controller setpoints in orde r to maintain co nstant product fraction width. The cutpoints in their turn are adjusted to keep constant analyzed product qualities ( Fig. 2 ) . These controls are significantly slower than the cutpoint controls because of analyser and process delay times. The two crude units are identical in contro l structure. The overhead product of the atmospheric crude units is further distilled to a number of fractions In downstream units and has no distillation range requirement. This fact, the overhead system in two stages and the overhead noncondensible gas being charged to the vapor recovery compressor gives the oportunity to minimize the overhead light naphtha for the benefit of heavy naphtha without the ordinary drawbacks of light overhead product. Therefore the overhead temperature minimizing constraint control is used. Note that ligh t naphtha is fed to the debutanizers, light naphtha splitter, deisopentanizer and depentanizer in this order. The reboilers of these use the pumparound heat

of both crude columns and thus propagate disturbances in the overhead system and the subsequent fractionation columns back to each crude column mId section .Th e lowest side draw product in any co lu mn. heavy gasoi! or cutter stock IS maximized against a number o f constraInts using special loqic in the constra int controller . This 1S because of the sensitive contamination of the product by entrained residue when over flash approaches ni l . Maximization is economically justified because of the substa ntial value difference between sidedraw and residue. Pumparound control Since vac . overhead is returned to crude unit 1 pumparound as an additional feed, also dIsturbances in each crude unit sidedraws and in vacuum unit are propagated back to the crude unit 1 midsection. Similar disturbances ent er from other un1ts to crude unit nr 2. The above underlines the Importance of stabillzation and optimization of the pumparound section in the column . The special features of pumparound loops make the conditions of stabillzation difficult. The stabili zation of the pumparound duty, for example , does not stabilize the internal condit i ons of the column. Th erefore the calculated internal flow rates at some tr ays above the pumparound return are contro ll ed to targets by manipulating a pumparound flow controller setpoint and a bypass f l ow simultaneously. The total pumparound duty is stabilized by adjusting the heat returned t o crude preheat when changes in other pumparound heat exchangers ( light naphtha fractionators rebailer ) or elsewhere occur. In order to assure operation inside feasible range of the operatIng window , internal flow rate min l max constraint control features are internal flow control l ers. The added to constraints are, besides the physical equipment limits, heat exchanger maximum duties, valve positions, temperature limits etc, also the column flooding and downcomer back-up limits . Pressure control Pressure minimizing constraint cont r oller is employed to ride th e constraints affecting pressure ( and internal flow rates ) . During constrained operation , pressure setpoints are relaxed dynamIcally in the fastest way, whi le the internal flow rate controller reacts slowly to the constraints that affect both the vapor rate and the pressure. The longer term eco nomic impacts that need interunit calculations i s then performed in t he energy op ti mizer, whi ch calculates the optimum internal flow r a t es and column feed temperature simultaneously. Observations show that in our cases the fl ow rates have been maximized within other limitations, while sacrificing the pressure. Our steady sta te simulations, however, show that this somew hat unique observation is no more necessarily valid if a vacuum unIt operating at minimum floating pressure is used instead. Then the optim um operation would also be much more se nsitive to the relative product values and the value of fuel gas. This underlines the need to use e nergy optimizer, capable of taking into accou nt the refinery str ucture and economics . Energv versus YIelds In a crude distillation column the separatio n understood as a function of the must be distillation curves . This is illustrated in Fig . 3 . The figure shows the crude oil true

Integrated Fractionation Processes in a Petroleum Refinery boiling point distillation curve and the corresponding product TBP-curves. If the separation was complete the product distillation curves should coincide with the crude oil curve for the distillation product in question. A low separation gives distillation "tails" containing high boiling point fraction ( which for the product for instance LGO will result in a low cloud point ) . A higher separation gives a smaller "tail" (a nd a better cloud point) which in turn enables a better yield on account of the product yield below. PARAMETERS AFFECTING THE SEPARATION It is important to know which parameters are possible to manipulate in order to affect the separation efficiency. These parameters are the variables which are manipulated when the objective function is maximized. In the following a crude di s tIllation system is analyzed. The column pressure affect s the separation, due to the increased relative volatilities at decreased pressures. In our case the optimal operation is minimi zation of pressure against constraints as for instance the maximum control valve positIon, flooding etc. The stripping stream decreases the hydrocarbon partial pressure and thus has the same effect as the above discussed column pressure. The optimum stripping steam ratios in crude columns and side strippers have been investigated through steady state s imulations. The hydrocarbon to steam ratios are maintained through ratio controllers. The internal flowrates are important parameters when trying to affect the separation. Increased internal flows leads to a better separation. The crude columns have one pump around and one top reflux with both warm and cold reflux, see Fig. 2. I f the pumparound duty is decreased the top reflux duty will automatically increase. This will lead to increased internal flows in the upper part of the column and subsequently also to better separation in the upper part. The separation below the pump around is not affected. This shows that if higher yields of lighter ( and in our case more expensive ) products are desired the column should be cooled mainly by top reflux. A high feed temperature also leads to higher internal flows, and better separation. OPTIMIZATION OF ENERGY RECOVERY AND YIELDS The objective function to be optimized is the sum of product income and net energy cost. High feed temperature results in an increased energy consumption. Cooling in the upper part of the column using cold reflux means that energy is transFerred to water instead of preheating crude, having the same effect. To be able to calculate how much extra energy is needed, rigorous simulations of the heat exchanger network have been carried out (s ee Fig. 4 ) . Also rigorous distillation simulations have been used to adjust the functions describing the yields change with changing pumparounds, top reflux and feed temperature. In crude and vacuum units ( with no reboiler ) all the primary energy for fractionation is entered Into the column as feed enthalpy and its source

63

is the absorbed heat in crude heater fired with fuel gas or oil. The feed flashes and generates internal vapor flow, the volume of which expands while meeting internal reflux and becoming lighter at every tray. The vapor rate can decrease only by removing heat from the column at the pumparounds. The column geometry requires some heat to be removed and also the overhead condensers are sized to match this. Optimization of the pumparound duty Generally the pumparound duty can widely vary without exeeding column constraints. A decreased duty causes increased vapor flow in the upper part of the column, which in turn improves separation and the yields of the lighter products above the pumparound at the expense of the lower products. The pressure effect will cause an increased long residue yield and the preheat changes of the heater feed will cause an increased fuel gas usage. Interaction to vacuum pump around exchangers will increase their duty affecting favourably the heavy vacuum gasoil yield. Steady state modelling technique with feedback information from the unit to adjust the models for successively increased accuracy is used. The final feedback on long term is the product laboratory distillation curves which are used on daily basis. Optimization of column feed temperature In these columns the other possibility to affect energy usage is to adjust the column feed temperature. This optimization is even more complex because it is dependent on the factors described earlier, over flash recovery in crude units and in vacuum unit and vacuum unit heater operation credit and the optimum value of the pumparound duty. The optimum column feed temperature is calculated and implemented via feed temperature constraint controllers. Vacuum unit optimization The vacuum unit has total draw-off trays for products and the unit is operated against a fixed minImum overhead product temperature. Light and heavy vacuum gasoil are blended and the heater operates at its maximum metal temperature. Therefore, in our case, the energy optimization is straightforwardly maximization of heavy gasoil pumparound duty (at the expense of light gasoil pumparound duty) against operating constraints. Note that Hy.VGO is maximized against its colour and other constraints by manipulating the wash flow back to the column, whereas during normal operation the over flash is returned to the feed. Light ends fractionation optimization In light ends fractionation chemical compounds can be analyzed and defined as light and heavy key components and treat the column as pseudo binary distillation. For example, in debutanizer i-pentane and in deisopentanizer n-pentane is the heavy key component. In all these columns composition control hard target is the distillate product heavy key composition, which is controlled according to specification, by indirectly shifting the column material balance. The soft composition target at bottoms product is allowed to have more variations and is indirectly controlled to a long term optimum value by using energy optimization which determines optimum column internal flow rates.

6--1

M. L. Sourander and S. Cros

These are implemented through a min / max. constraint controller by adjusting the bottom reboiler duty. Since the reboiler energy is derived from the crude unit pumparound enthalpy this calls for a comparison of the relative usefulness of energy in light end columns, crude columns or crude unit preheat. CONTROL PERFORMANCE AND RESULTS The computer control system has been operating successfully for about a year. The availability of the control functions in the digital instrumentation system has been virtually 100 % due to the redundant equipment ( measurement and control valve faults not included ) . The computer control availability has been 99,6 per cent, the floating point accelerator card being the only observed faul t. The control performance and optimization effects have keenly been followed by the operating personnel and compared to the time prior to computer control. The economic improvements attributable to the computer control show that the payback time has been about one year. The result are shown in Table 1.

TABLE 2.

Computer Control Results

Product quality

before

after

2.5 N.A.

1.0 0.7 0.7 0.2 1.0

Standard deviation Hy Naphtha 90 % pt. Kerosene density Lt gasoil cloud pt. Gasoil viscosity Hy.gasoil colour

Product yield improvement per cent of crude oil charge Naphtha Kerosene Lt gasoil Gasoil Hy. gasoil long residue

1.5

0.4 1.6

base

+

+

0,8 1,0 0,5 2,1 1,0 0,2

REFERENCES Quality control The product qualities are clearly more stable than before the computer control. The yields of more valuable products have increased due to optimized pumparound duties and column feed temperatures, while keeping the same quality targets. This is due to less pronounced "tails" in the distillation curves of the products. Energy allocation In the product value situations of the past year, despite of the tremendous changes in the crude oil price, the product value differences have not changed directionally. Also the optimum strategy that the computer has implemented seems to be stable in nature. In all cases both crude units' pumparound duties have been at maximum limits and the column feed temperatures at maximum limits, both determined by column and feed heater design geometry factors.

Sourander, M. & al."Control and Optimization of olefinCracking Heaters", Hydrocarbon Processing Vol 63, No 6, 1984 Nasi, M. & al. "Experience with Computers Control of an Ethylene Plant", Hydrocarbon processing, Vol 62, No 6, 19B3. Rinne, R. & a1. "Experience with Distillation Unit Computer Control", Hydrocarbon Processing, Vol 61, No 4, March, 19B2.

Integra ted Fractionation Processes

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65

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