a plantwide control procedure applied to the hda process
DESCRIPTION
A Plantwide Control Procedure Applied to the HDA Process. Antonio Araújo and Sigurd Skogestad Department of Chemical Engineering Norwegian University of Science and Technology (NTNU) Trondheim, Norway November, 2006. Outline. General procedure plantwide control HDA process - PowerPoint PPT PresentationTRANSCRIPT
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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
9
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!
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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
7911
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!
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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).