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Ind. Eng. Chem. Res. 2000, 39, 718-723
New Control System for Crude Oil Processing Improvement of Control Quality and Plant Operability Stanislav Krejcˇ ı´,* Jirˇ ı´ Macha´ cˇ ek, and Frantisˇ ek Dusˇ ek Department of Process Control and Computer Techniques, University of Pardubice, 53210 Pardubice, Czech Republic
Pavel Hrubesˇ PARAMO a.s., 53006 Pardubice, Czech Republic
Most control loops in many installed control systems use only standard, linear PID controllers. Such situations offer the possibility to improve both control quality and plant operability without additional investments. The design of the new control system for crude oil processing in the PARAMO refinery a.s. in Pardubice, Czech Republic, was divided into three different parts: (1) A new control strategy based on the gain scheduling method for the pressure control and the feedforward method for the level control has been proposed for stabilization of the desalter tank. (2) A special ratio controller connected in cascade to the original flow controllers has been used for the minimization of the temperature difference between the heated streams (not controlled in the original system). (3) A combined feedforward-feedback cascade system has been designed for the feed temperature control. All proposed systems have been tested and set in operation. It has been proven that there is outstanding improvement over the original control system in stabilization of the output temperature, in simplification of the system operability, and in savings of heating medium. 1. Introduction In chemical industry we can relatively often meet situations when an installed control system has not utilized all possible functions for the control of complicated technologies. Such a system usually comprises perfectly worked out data acquisition and visualization subsystems for a great number of measured and manipulated variables, but most control loops use only standard, linear PID controllers. This state of affairs offers a wide area for improving both the control quality and the plant operability without additional investments. The challenging procedure has been realized in the PARAMO refinery a.s., Pardubice, Czech Republic. 2. Description of Technology
Figure 1. Simplified technology scheme.
The described technology can be divided into two main parts. Raw crude oil enters the first part of apparatuses where the stream is preheated and desalted. In the second part of the technology the desalted crude oil is heated to the temperature needed for the following atmospheric distillation. The simplified technology scheme can be seen in Figure 1. Raw crude oil is pumped from input tanks by pump A or B (chosen by an operator according to the flow demands) through a series of preheaters into a mixer. In the preheaters the crude oil is heated to the temperature 110-130 °C using the heat of some outlet streams from atmospheric distillation. Heated crude oil is then mixed with hot demineralized water, and the mixture enters the desalter tank. The mixture is then divided into two layerssthe refined crude oil and water with
dissolved salts under a relatively high pressure and using an electrostatic field. The high pressure (700900 kPa) must be kept in the desalter tank in order to prevent releasing of gases from crude oil. Water with dissolved salts is drained continually from the lower layer in the desalter tank. Both the pressure and the level in the desalter tank are stabilized by control loops with PID controllers. The pressure is kept on the reference value by a pneumatic throttled valve, which is linked up after a pump. The level between two layers in the desalter tank is controlled by a pneumatic control valve, which can change the water flow. Refined crude oil from the upper layer is pumped from the desalter tank through a series of three preheaters. The heating to 180-215 °C is performed by outlets removed from different distillation columns in the plant. After leaving the preheaters, crude oil is split into two parallel streams that enter two different heating coils inside the heating furnace. The both streams are heated to 330350 °C by combustion of gas or light oil. The outlets of
* To whom correspondence should be addressed. Tel: 0042040-6037504. Fax: 00420-40-6037068. E-mail: stanislav.krejci@ upce.cz.
10.1021/ie990241j CCC: $19.00 © 2000 American Chemical Society Published on Web 01/22/2000
Ind. Eng. Chem. Res., Vol. 39, No. 3, 2000 719
Figure 3. Gain-scheduling controller for the desalter tank.
Figure 2. Designed control system.
both heating coils are put together, forming the feed to the distillation column. The flows in both heating coils are controlled by simple PID controllers, whose set points are set by an operator in order to hold outlet temperatures of both streams on the same values. The feed temperature is stabilized by a PID controller which makes adjustment in the pressure of heating gas or oil. The whole technology is controlled by the control system Honeywell TDC 3000. PID controllers are digitalized with the sampling period T ) 1 s but are taken as continuous. The transfer function of the controllers can be expressed as
(
C(s) ) K 1 +
1 + Tds Tis
)
(1)
where K is a proportional gain, Ti is an integral constant, and Td is a derivative constant. The control system enables one to use some functions for setting other controller properties (e.g., limited speed of the valve position change) and to create new control programs. Both of these possibilities were used in the designed control strategy. 3. Design of an Improved Control System The main target of the improved control system was better control quality in the sense of stabilization of the whole system operation and simplification of the operator’s work. The design of the improved control system (Figure 2) was divided into three different parts: (a) stabilization of the desalter tank, (b) minimization of the temperature difference between both heating coils, and (c) stabilization of the feed temperature. Because these parts follow in sequence, every improvement in previous parts positively influences subsequent parts, as well. 3.1. Desalter Tank. The desalter tank is a multiinput multioutput system with inner interactions, where the basic controlled variable is the pressure inside of the tank. The second controlled variable is the level between two layers in the tank. The main disturbances are changes of the refined oil flow from the tank, which depend on the requested production. This system is highly nonlinear; its gains, time delays, and time constants change in wide ranges according to an actual operation state. On top of that, the dynamic nonlinearities depend on the change in direction. In the original state, decentralized one-dimensional control loops with PID controllers were used.
Figure 4. Relay autotuning method.
3.1.1. Pressure Control. The pressure inside of the desalter tank is influenced by values of oil and water flows and by pressure of the supplied oil depending on the used pump type (A/B). Both flows are controlled by PID controllers. The pressure controller with fixed parameters was used in the original design. Because operation conditions are often changed, the pressure controller had to be adjusted to ensure the stability of the control loop in all cases. A satisfactory control quality then could not be ensured for a great deal of operating situations. The newly designed control strategy based on the gain scheduling method ensures a good quality of pressure control for all operation conditions with the same PID controller. The gain scheduling was chosen as an adaptive method because of its advantages in comparison with the other adaptive techniques.1 It is a simple method, which can follow changes in the operating points conditions. The operating range is divided into several smaller ranges where the process can be well approximated by linear models. A controller for the full operating range can then be obtained by determining one model for each operating range and changing the controller parameters with operating conditions. The crude oil flow and the pump type (A/B) were chosen as the scheduling variables for the PID controller tuning in our design. The controller gain depends on the oil flow and the pump type (A/B). The controller time constants depend on the type of the used pump only (Figure 3). The PID controller was tuned for selected values of the flow and the procedure was repeated for both types of pumps. The system dynamics was tested by the relay autotune method (see ref 2). A nonlinear feedback of the relay type with the output amplitude (M was introduced in order to generate a limit cycle oscillation (Figure 4). The ultimate period Tu and amplitude of the oscillation A were determined when steady-state oscillation was obtained (Figure 5). The ultimate gain ku is then given by
ku ) 4M/πA
(2)
PID controller parameters were then computed according to the Ziegler-Nichols rule
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Ind. Eng. Chem. Res., Vol. 39, No. 3, 2000 Table 1. Controller Parameters pump
flow [tons h-1]
K
Ti [min]
Td [min]
A
65 80 90 100 70 80 90 100 110
0.10 0.30 0.70 0.80 0.07 0.08 0.11 0.15 0.30
0.33 0.33 0.33 0.33 0.7 0.7 0.7 0.7 0.7
0.084 0.084 0.084 0.084 0.18 0.18 0.18 0.18 0.18
B
Figure 5. Signals from the autotuning test.
K ) 0.6ku
Ti ) 0.5Tu
Td ) 0.125Tu
(3)
Values of the controller parameters as a function of the oil flow and the pump type are given in Table 1. The newly designed controller has limited speed of a valve position change to 60%/min. In such cases all derivative parts will be suppressed. The direction nonlinearity was suppressed by reducing of the controller gain to 60% with exception of the case when the controlled pressure deviates more than 20 kPa compared with the set point. Responses of the new and the original control systems are shown in Figure 6 and numerical results in Table 2. When the crude oil flow is low, the process output oscillates around the set point. It is caused by the friction in the valve (its position changes in a narrow range). It is not possible to improve such an undesirable behavior by the controller detuning.3 3.1.2. Level Control. The level control of the boundary between crude oil and salted water layers in the desalter tank influences the pressure inside the tank and decreases the quality of the pressure control. Every change of the manipulated variable immediately changed the pressure inside the tank. For this reason the level controller was originally set up in such a way that the manipulated variable changed only slowly and the resulting control quality was then bad. The new strategy was designed to improve the level control quality and at the same time the pressure control did not get worse. The acceptable solution of this problem was found on the simulation model of the desalter tank using a feedforward compensator. The approximate transfer functions were identified from step or pulse responses. The feedforward compensator with the transfer function GFF(s) has to fulfill the following condition:
∆u2(s) G21(s) + ∆u2(s) GFF(s) G11(s) ) 0
Figure 6. Pressure control with a new and original controller.
influence on the pressure inside the tank is accordant with the design condition. 3.2. Temperature Difference Control. Desalted crude oil is divided into two heating coils before entering the furnace (see Figure 1). The flows in both heating coils are different at the same openings of the valves with regard to various hydrodynamic conditions, and the same is valid for temperatures. The temperature difference was not directly controlled in the original system but was kept at a minimum value by setting of the reference signals in both flow control loops. The total flow was kept at a constant value at the same time. The small temperature difference is required also for stabilization of the feed temperature and for reducing of the energetic wastes. Isobaric-adiabatic mixing of two streams with different temperatures is a typical irreversible process. Every irreversible process is accompanied by an entropy increase. The more irreversible the process is, the bigger entropy is and the bigger are energetic wastes. The new control strategy consists of a special ratio controller connected as a cascade controller to the original flow controllers (Figure 2). The ratio controller uses the temperature difference as the controlled variable and the flow ratio x defined by eq 5 as the
(4)
where G21(s) is the cross-transfer function and G11(s) is the direct one for the pressure control. Changes of the control variable from the level controller ∆u2 are taken as measured disturbances for the pressure control (Figure 7). The feedforward control requires the knowledge of accurate models, and it was not fulfilled in this case. That is why only static feedforward was suggested and set up for average operating conditions. The setting of the PI level controller was found by simulation for the condition in which pressure changes were smaller than 100 kPa. The designed level controller works faster than the original one (Figure 8 and Table 3), and the
x)
Q1 Q1 + Q 2
(5)
manipulated variable. The set point for the first controller is calculated as a product of the total flow Q ) Q1 + Q2 and the flow ratio x (i.e., Qx) and that for the second controller as the product Q(1 - x). The set point for the ratio controller (controlling the temperature difference) is zero. Changes in valve openings unfavorably influence the pressure inside the desalter tank, and that is why only limited speed of the valve position change was used so that the pressure controller had enough time for a disturbance suppression. Stabilization of the control loop
Ind. Eng. Chem. Res., Vol. 39, No. 3, 2000 721 Table 2. Comparison of the Original and the New Control of the Pressure IAE [kPa min] pump
flow [tons
A
h-1]
changes in pressure set point [kPa]
95
800 f 700 700 f 800 800 f 700 700 f 800 800 f 700 700 f 800 800 f 700 700 f 800 800 f 700 700 f 800 800 f 700 700 f 800
80 65 B
70 90 110
Figure 7. Feedforward compensator.
Figure 8. Level control with a new and original controller.
was also improved by better filtration in SMART sensors for both streams flow measurements. Dynamics of the controlled systems was approximated by a first-order model with time delay in the form
G(s) )
Ks -sτd e τs + 1
(6)
where Ks is the static process gain, τ the time constant, and τd the time delay. Parameters of these models were estimated from flow and temperature responses to step changes in valve positions. All controller parameters were then calculated according to the modification of the Ziegler-Nichols method for 0% overshoot:4
K)
0.95τ τdKsk
Ti ) 2.4τd
Td ) 0.42τd
(7)
where Ksk is gain Ks in engineering units recounted for the controller range 0-100. The new controllers keep the temperature difference in the range (1.5 °C for all disturbances (Figure 9), and its performance gives practically negligible changes of the total flow (see Table 4). Load disturbances have no
original
new
728.8 933.5
93.9 82.9 92.3 104.7 92.2 84.5
94.0 99.2
ISE [kPa2 min] original
new
48587 62235
5615 3773 5025 4525 5078 3809
6732 5620
76.8 82.0 90.3 130.1 230.5 201.3
4660 5920 80.1 89.2 67.3 105.1
6779 7131 13471 11650
4610 5389 5026 5548
effect on the control performance because it was proved by the F test (F ) 1.159, Fcrit(R)0.01,n1)n2)479)). 3.3. Feed Temperature Control. A fluctuation of the resulting crude oil temperature in the heating furnace output is produced by several effects: (a) changes in the total flow of the heated crude oil, (b) changes in an input temperature of the heated crude oil, (c) temperature difference between two streams before connecting, and (d) quality of a combusting medium. In the original state the temperature control was performed by one PID pressure controller, which could not suppress all mentioned disturbances. The new control strategy uses a combined feedforward-feedback cascade system where the first controller controls the output temperature (feed to the distillation column) and its output is connected with the second original controller as a set point (Figure 2). The feedforward part detects major disturbances (the load and the inlet temperature) and makes adjustment in the manipulative variable. The feedback part takes care of any errors that come through the process because of inaccuracies in the feedforward controller or other unmeasured disturbances. The feedforward controller contains steady-state gains only. The steady-state gains were determined from a mathematical model of the heating furnace. The same temperature Tp everywhere inside the furnace (see Figure 10) is assumed in the model, i.e., a perfectly mixed system. Heating coils for crude oil are considered to be plug-flow systems. The mathematical model of the furnace consists of heat balances for both crude oil flows:
QiciTi + o dh Ki(Tp - Ti) ) Qici(Ti + dTi)
i ) 1, 2 (8)
where Qi is the mass flow of an individual flow, ci the heat capacity of an individual flow, Ti the input temperature of an individual flow, o the pipe circumference, Tp the temperature inside the furnace, Ki the overall heat-transfer coefficient of an individual pipe, and dh the element of the coil length. After arrangement of heat balances, output temperatures of both crude oil streams can be solved in the form
T1b ) Tp - (Tp - T1a)e-K1F1/Q1c1 T2b ) Tp - (Tp - T2a)e-K2F2/Q2c2 (9) where F1 and F2 are the surfaces of the coils.
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Ind. Eng. Chem. Res., Vol. 39, No. 3, 2000
Table 3. Comparison of the Original and the New Control of the Level ISE [%2 min]
IAE [% min] changes in level set point [%] 70 f 80 80 f 70
settling time [min]
max. changes in pressure [kPa]
original
new
original
new
original
new
original
new
137.1
23.3 16.2
692.1
141.6 94.2
43
10 6.5
22
52 82
Table 5. Comparison of Steady-State Gains parameter
computed
measured
Zs4 - Zs1 [°C] Zs3 [°C h tons-1] Zs6 [°C h tons-1]
517.52 -3.09 -3.00
512.9 -3.20 -3.13
where deviations of temperatures ∆T1b and ∆T2b can be solved from steady states:
∆T1b )
∂T1b ∂T1b ∆T1a + ∆Q1 ∂T1a ∂Q1 ∂T2b ∂T2b ∆T2b ) ∆T2a + ∆Q2 (14) ∂T2a ∂Q2
where Figure 9. Temperature difference control - flow disturbances 5 tons h-1. Table 4. Results of the New Temperature Difference Control changes in flow [tons h-1]
IAE [°C min]
ISE [°C2 min]
standard deviation [°C]
NO (6:50-7:30) 95-90 (7:30-8:10)
18.4 20.7
12.9 14.5
0.561 0.604
∆Q1 ) x∆Q + Q∆x
Partial derivative terms consist of parameters that are not at disposal for such a complex system (overall heat-transfer coefficients and heat capacities). That is why the terms have to be rearranged using only known and measurable parameters:
Zsl )
Zs4 )
Figure 10. Heating furnace sketch.
The resulting temperature of the crude oil after mixing of both streams is given by the heat balance:
Q1c1T1b + Q2c2T2b ) (Q1 + Q2)csTs Q1c1 Q2c2 Ts ) T1b + T (10) (Q1 + Q2)cs (Q1 + Q2)cs 2b where cs is heat capacity of the resulting stream and Ts is the temperature of the resulting stream. The input flow Q is divided in two streams:
Q1 ) xQ
Q2 ) (1 - x)Q
(11)
Ts ) xT1b + (1 - x)T2b
(12)
Deviation of Ts from a steady state is given by
∆Ts ) (T1b - T2b)∆x + x∆T1b + (1 - x)∆T2b
∂T1b Tp - T1b Tp - T1b ) ln ∂x x Tp - T1a ∂T1b Tp - T1b Zs2 ) ) ∂T1a Tp - T1a ∂T1b Tp - T1b Tp - T1b Zs3 ) ) ln (16) ∂Q1 Q1 Tp - T1a ∂T2b Tp - T2b Tp - T2b )ln ∂x 1-x Tp - T2a ∂T2b Tp - T2b Zs5 ) ) ∂T2a Tp - T2a ∂T2b Tp - T2b Tp - T2b Zs6 ) ) ln (17) ∂Q2 Q2 Tp - T2a
Substituting eqs 14-17 into eq 13 and assuming T1a ) T2a:
∆Ts ) [(T1b - T2b) + QxZs3 - Q(1 - x)Zs6]∆x + [(xZs2 + (1 - x)Zs5]∆T1a + [x2Zs3 + (1 - x)2Zs6]∆Q (18) Operating conditions of the furnace are given by the parameters:
Q1 ) 41.5 tons h-1
Assuming c1 ) c2 ) cs, eq 10 can be rewritten as
(13)
∆Q2 ) (1 - x)∆Q - Q∆x (15)
Tp ) 655 °C
Q2 ) 43.5 tons h-1 T1a ) T2a ) 192 °C T1b ) 352 °C
T2b ) 355 °C
The total crude oil flow into the furnace was Q ) 85 tons h-1, and the resulting ratio x ) Q1/(Q1 + Q2) ) 0.488. Steady-state gains can be solved for these pa-
Ind. Eng. Chem. Res., Vol. 39, No. 3, 2000 723 Table 6. Comparison of the Original and the New Control of the Temperature ISE [°C2 min]
IAE [°C min] changes in flow [tons 95-90 75-70
h-1]
standard deviation [°C]
overshoot [°C]
original
new
original
new
original
new
original
new
696 101
92 54
4330 255
137 57
5.097 1.839
1.140 0.866
10.0 4.8
3.1 2.2
rameters:
Zs1 ) -263.26 [°C]
Zs2 ) 0.654 [1] Zs3 ) -3.09 [°C h tons-1]
Zs4 ) 254.26 [°C]
Zs5 ) 0.648 [1] Zs6 ) -3.00 [°C h tons-1]
Computed steady-state gains are compared with measured values in Table 5 and proved a good agreement. The basic notion of feedforward control is to detect disturbances as they enter the process and make adjustments in manipulative variables so that the output variable is held constant. The relationship among the output oil temperature Ts and significant disturbances (flow ratio x, flow Q, input temperature T1a, and heating medium-pressure Pg) is given by
∆Ts ) 0 ) Zsx∆x + ZsQ∆Q + ZsT1a∆T1a + Zsp∆Pg
(19)
where values Zsx, ZsQ, and ZsT1a were determined from eq 18 and Zsp was determined from an experiment. Because the temperature difference control worked very well, the influence of flow ratio changes was neglected. The feedforward controller was designed as a static gain, which was computed in the same way as the level controller (chapter 3.1.2). Signals of both of the feedforward compensators, i.e., flow and temperature changes modified by calculated coefficients, are summed with the feedback part of the temperature controller. The resulting signal is the setpoint value of the gas-pressure controller. Parameters settings of the PID controllers were made on the basis of the measured step responses and eq 7. The new control system was proven in all cases: (a) stability for all types of changes and (b) a more favorable performance in comparison with the original control system. The effect of the load disturbances on the feed temperature was measured for all possible situations for long periods of time. A typical response is shown in Figure 11. Typical quantities used for the control quality characterization are overshoot, settling time, and the integrated function of the error, e.g., the absolute error IAE (integral absolute-error criterion) or the quadratic criterion ISE (integral square-error criterion).5 The comparison of the original and the new control systems is illustrated in Table 6. 3.4. Influence of the Temperature Difference on Heating Gas Consumption. The influence of the stream temperature difference (leaving the heating furnace) on the heating gas consumption was measured
Figure 11. Feed temperature difference control - flow disturbances 5 tons h-1.
as well. The objective of the measurement was to verify the reduction of the process irreversibility. The test was divided into two parts. During the first part, the set point of the temperature difference controller was 0 °C, and during the second part, 5 °C and the heating furnace performance was stabilized. The achieved heating gas savings was 63 kg h-1. 4. Conclusion The improved control system for crude oil preparation in the distillation part has been designed and set in operation. The changes were made only in the computer control system without changes in technological apparatuses. Outstanding improvement over the original system consists of stabilization of the output temperature, simplifying of the system operability, and savings of heating medium. Literature Cited (1) Åstro¨m, K. J.; Wittenmark, B. Adaptive Control, 2nd ed.; Addison-Wesley Publishing Company: Reading, MA, 1995. (2) Åstro¨m, K. J.; Ha¨gglund, T. Automatic Tuning of Simple Regulators with Specification on Phase and Amplitude Margins. Automatica 1984, 20, 645-651. (3) Åstro¨m, K. J.; Ha¨gglund, T. PID Controllers: Theory, Design, and Tuning, 2nd ed.; ISA: Research Triangle Park, NC, 1995. (4) Chien, K. L.; Hrones, J. A.; Reswick, J. B. On the automatic control of generalized passive systems. Trans. ASME 1952, 74, 175-185. (5) Ogata, K. Modern Control Engineering, 2nd ed.; Prentice Hall Inc.: Englewood Cliffs, NJ, 1990.
Received for review April 1, 1999 Revised manuscript received September 8, 1999 Accepted September 29, 1999 IE990241J
