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A Plantwide Control Procedure Applied to the HDA Process
Antonio Araújo and Sigurd Skogestad
Department of Chemical EngineeringNorwegian University of Science and Technology (NTNU)Trondheim, Norway
November, 2006
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Outline
• General procedure plantwide control• HDA process• Active constraints• Self-optimizing variables• Maximum throughput mode• Regulatory control• Dynamic simulations
– comparison with Luyben
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General procedure plantwide control
y1s
y2s
Control of primary variables(MPC)
“Stabilizing” control:p, levels, T (PID)
Part I. “Top-down” steady-state approach - identify active constraints and primary controlled variables (y1)
– Self-optimizing control
Part II. Bottom-up identification of control structure – starting with regulatory (“stabilizing”) control layer.
– Identify secondary controlled variables (y2)
RTO. min J (economics). MV = y1s
u (valves)
Skogestad, S. (2004), “Control structure design for complete chemical plants”, Computers and Chemical Engineering, 28, 219-234.
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Part I. Top-down steady-state approach
Step 1. IDENTIFY DEGREES OF FREEDOMNeed later to choose a CV (y1) for each
Step 2. OPERATIONAL OBJECTIVES Optimal operation: Minimize cost J
J = cost feeds – value products – cost energy subject to satisfying constraints
Step 3. WHAT TO CONTROL? (primary CV’s c=y1)
What should we control (y1)?1. Active constraints2. “Self-optimizing” variables
These are “magic” variables which when kept at constant setpoints give indirect optimal operation by controlling some “magic” variables at– Maximum gain rule: Look for “sensitive” variables with a large scaled steady-state gain
Step 4. PRODUCTION RATE
y1s
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Part II. Bottom-up control structure design
Step 5. REGULATORY CONTROL LAYER (PID)
• Main objectives– “Stabilize” = Avoid “drift”– Control on fast time scale
• Identify secondary controlled variables (y2)
– flow, pressures, levels, selected temperatures– and pair with inputs (u2)
Step 6. SUPERVISORY CONTROL LAYER – Decentralization or MPC?
Step 7. OPTIMIZATION LAYER (RTO)– Can we do without it?
y2 = ?
u (valves)
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Two main modes of optimal operation for chemical plants
Depending on marked conditions:
Mode I: Given throughputWhen: Given feed or product rate
Optimal operation: Max. efficiency
Mode II: Maximum throughput (feed available). When: High product prices and available feed Optimal operation: max. flow in bottleneck
1. Desired: Same or similar control structure in both cases2. Operation/control: Traditionally: Focus on mode I But: Mode II is where the company may make extra money!
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Mixer FEHE Furnace PFR
Quench
Separator
Compressor
Cooler
StabilizerBenzeneColumn
TolueneColumn
H2 + CH4
Toluene
Toluene Benzene CH4
Diphenyl
Purge (CH4 + H2)
HDA process
Toluene + H2 = Benzenje + CH4
2 Benzene = Diphenyl + H2
References for HDA:McKetta (1977) ;
Douglas (1988) Wolff (1994)Luyben (2005)++....
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Mixer FEHE
Furnace
Reactor
Quencher
Separator
Compressor
Cooler
StabilizerBenzeneColumn
TolueneColumn
H2 + CH4
Toluene
Toluene Benzene CH4
Diphenyl
Purge (H2 + CH4)
1
2
3
64
7
5
1113
12 10 8
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Step 1 - Steady-state degrees of freedom
NEED TO FIND 13 CONTROLLED VARIABLES (y1)
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Step 2 - Definition of optimal operation
• The following profit is to be maximized:
-J = pbenDben + Σ(pv,iFv,i) – ptolFtol – pgasFgas – pfuelQfuel – pcwQcw – ppowerWpower - psteamQsteam
• Constraints during operation:– Production rate: Dben ≥ 265 lbmol/h.– Hydrogen excess in reactor inlet: Fhyd / (Fben + Ftol + Fdiph) ≥
5.– Reactor inlet pressure: Preactor,in ≤ 500 psia.– Reactor inlet temperature: Treactor,in ≥ 1150 °F.– Reactor outlet temperature: Treactor,out ≤ 1300 °F.– Quencher outlet temperature: Tquencher,out ≤ 1150 °F.– Product purity: xDben ≥ 0.9997.– Separator inlet temperature: 95 °F ≤ Tseparator ≤ 105 °F.– Compressor power: WS ≤ 545 hp– Furnace heat duty: Qfur ≤ 24 MBtu– Cooler heat duty: Qcool ≤ 33 MBtu– + Distillation heat duties (condensers and reboilers).
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Disturbances
D1 Fresh toluene feed rate [lbmol/h] 300 285
D2 Fresh toluene feed rate [lbmol/h] 300 315
D3 Fresh gas feed rate methane mole fraction 0.03 0.08
D4 Hydrogen to aromatic ratio in reactor inlet 5.0 5.5
D5 Reactor inlet pressure [psi] 500 520
D6 Quencher outlet temperature [oF] 1150 1170
D7 Product purity in the benzene column distillate 0.9997 0.9960
Typical disturbances :• Feeds• Utilities• Constraints
Caused by: implementation error or change
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Step 3: What to control?
• 13 steady-state degrees of freedom• 70 Candidate controlled variables
– pressures, temperatures, compositions, flow rates, heat duties, etc..
• Number of different sets of controlled variables:
• Cannot evaluate all !
1370 70!4.75 10
13 57!13!
æ ö÷ç ÷= = ×ç ÷ç ÷çè ø
OPTIMAL OPERATION:1. Control active constraints!
Find from steady-state optimization (step 3.1)
2. Remaining unconstrained DOFs: Look for “self-optimizing” variables (step 3.2)
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Operation with given feedMode I
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Step 3.1 – Optimization distillation• Distillation train:
– Optimized separately using detailed models– Generally: Most valuable product at its constraint– Other compositions: Trade-off between recovery and energy– Results:
Stabilizer
xD,benzene 1 · 10-4
xB,methane 1 · 10-6
Benzene column
xD,benzene 0.9997
xB,benzene 1.3 · 10-3
Toluene column
xD,diphenyl 0.5 · 10-3
xB,toluene 0.4 · 10-3
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Step 3.1 – Optimization entire process
• Reactor-recycle part• With simplified distillation section (constant compositions)
Distillation compositions
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Step 3.1 – Optimization: Active Constraints
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Mixer FEHE
Furnace
Reactor
Quencher
Separator
Compressor
Cooler
StabilizerBenzeneColumn
TolueneColumn
H2 + CH4
Toluene
Toluene Benzene CH4
Diphenyl
Purge (H2 + CH4)
8
1
4
2
610
4
3
5
1. Max. Toluene feed rate 2. Min. H2/aromatics ratio3. Min. Separator temperature4. Min. quencher temperature5. Max. Reactor pressure6. Max. impurity product
+ 5 distillation purities
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Step 3.2: What more to control?
• So far: Control 6 active constraints + 5 compositions (“self-optimizing”)
• What should we do with the 2 remaining degrees of freedom?– Self-optimizing control: Control variables that
give small economic loss when kept constant
• But still many alternative sets
• Prescreening: Use “maximum gain rule” (local analysis) for prescreening– Maximize σ(S1·G2x2·Juu
-1/2).– Optimal variation and implementation error enters in S1
59 59!1711
2 57!2!
æ ö÷ç ÷= =ç ÷ç ÷çè ø
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σ(S1·G2x2·Juu-1/2) = 2.33·10-3 Average Loss (k$/year)
Mixer outlet inert (methane) mole fractionQuencher outlet toluene mole fraction
15.39
σ(S1·G2x2·Juu-1/2) = 2.27·10-3 Average Loss (k$/year)
Mixer outlet inert (methane) mole fractionToluene conversion at reactor outlet
26.55
σ(S1·G2x2·Juu-1/2) = 2.25·10-3 Average Loss (k$/year)
Mixer outlet inert (methane) mole fractionSeparator liquid benzene mole fraction
31.39
• Linear model• All measurements: σ(S1Gfull·Juu
-1/2) = 6.34·10-3
• Best set of two measurements involves two compositions:
c1c2
Step 3.2 – “Maximum gain rule”
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Step 3 - Final selection in mode I
c1 c2
Mixer FEHE
Furnace
Reactor
Quencher
Separator
Compressor
Cooler
StabilizerBenzeneColumn
TolueneColumn
H2 + CH4
Toluene
Toluene Benzene CH4
Diphenyl
Purge (H2 + CH4)
8
1
4
2
7
6
9
10
11
4
3
5
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Step 3: What to control in Mode II ?
Available feed and good product pricesMaximum throughput
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Optimization in mode II: Maximum throughput• 14 steady-state degrees of freedom (one extra) • Reoptimize operation with feedrate Ftol as parameter:
– Find same active constraints as in Mode I.– At Ftol = 380 lbmol/h: Compressor power constraint active.– At Ftol = 390 lbmol/h: Furnace heat duty constraint active.– Further increase in Ftol infeasible: Furnace is BOTTLENECK!
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Step 3 - Controlled variable mode II• 8 active constraints (including WS and Qfur )
• + 5 distillation compositions• One unconstrained degree of freedom:
– To reduce the need for reconfiguration we control x-methane
– Average loss 68.74 k$/year
c1
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Step 4 – Throughput manipulator
• Mode I: Toluene feedrate (given)• Mode II: Optimal throughput manipulator is
furnace duty (bottleneck)– Minimizes back-off– But furnace duty is used to stabilize reactor– So use toluene feedrate also in mode II
c1
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Part II: Bottom-up designstarting with regulatory layer
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Step 5: Regulatory layer - Stabilization• Control reactor temperature and liquid levels in separator and
distillation columns (LV configuration).
LC01
LC11LC21LC31
LC32 LC22 LC12
TC01
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Regulatory layer - Avoiding drift I: Pressure control
LC01
LC11LC21LC31
LC32 LC22 LC12
PC01
PC11PC22PC33
TC01
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Regulatory layer - Avoiding drift II: Temperature control
LC01
LC11LC21LC31
LC32 LC22 LC12
PC01
PC11PC22PC33
TC02
TC03
TC22
TC11
#20
#3#5
TC33
TC01
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Regulatory layer - Avoiding drift III: Flow control
LC01
LC11LC21LC31
LC32 LC22 LC12
PC01
PC11PC22PC33
TC02
TC03
TC22
TC11
#20
#3#5
TC33
FC01
FC02
TC01
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Step 6: Supervisory layer – Mode I
LC01
LC11LC21LC31
LC32 LC22 LC12
TC01
PC01
PC11PC22PC33
TC02
TC03
TC22
TC11
#20
#3#5
TC33
FC01
FC02
RC01
CC01
CC02
CC21
CC22
CC32
CC31
CC12
CC11
Decentralized control (PID-loops) seems sufficient
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Step 6: Supervisory layer – Mode II
LC01
LC11LC21LC31
LC32 LC22 LC12
TC01
PC01
PC11PC22PC33
TC02
TC03
TC22
TC11
#20
#3#5
TC33
SETPOINT=Max.fuel-backoff
FC02
RC01
CC01
CC21
CC22
CC32
CC31
CC12
CC11
FixedDecentralized control (PID-loops) seems sufficient
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Dynamic simulations – Mode IDisturbance D1: +15 lbmol/h (+5%) increase in Ftol .
Ours Luyben’s
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Dynamic simulations – Mode IDisturbance D2: -15 lbmol/h (-5%) increase in Ftol .
Ours Luyben’s
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Dynamic simulations – Mode IDisturbance D3: +0.05 increase in xmet.
Ours Luyben’s
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Dynamic simulations – Mode IDisturbance D4: +20 psi increase in Prin.
Ours Luyben’s
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Conclusion
Procedure plantwide control:
I. Top-down analysis to identify degrees of freedom and primary controlled variables (look for self-optimizing variables)
II. Bottom-up analysis to determine secondary controlled variables and structure of control system (pairing).
