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HDA case study

HDA case study. S. Skogestad, May 2006 Thanks to Antonio Araújo. CH 3. →. +. H 2. +. CH 4. +. Heat. +. H 2. →. 2. ←. Process Description. Benzene production from thermal-dealkalination of toluene (high-temperature, non-catalytic process). Main reaction: Side reaction

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HDA case study

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  1. HDA case study • S. Skogestad, May 2006 • Thanks to Antonio Araújo

  2. CH3 → + H2 + CH4 + Heat + H2 → 2 ← Process Description • Benzene production from thermal-dealkalination of toluene (high-temperature, non-catalytic process). • Main reaction: • Side reaction • Excess of hydrogen is needed to repress the side reaction and coke formation. • References for HDA process: • McKetta (1977) – first reference on the process; • Douglas (1988) – design of the process; • Wolff (1994) – discuss the operability of the process. • No references on the optimization of the process for control structure design purposes. Benzene Toluene Diphenyl

  3. Purge (H2 + CH4) Compressor H2 + CH4 Toluene Quench Mixer FEHE Furnace PFR Separator Cooler CH4 Toluene Benzene Toluene Column Stabilizer Benzene Column Diphenyl Process Description

  4. Steady-state operational degrees of freedom 14

  5. Steady-state operational degrees of freedom 8 Purge (H2 + CH4) 7 Compressor 1 Furnace 3 H2 + CH4 Quencher Toluene Mixer FEHE Reactor 4 5 2 Cooler 6 13 11 9 Separator Benzene CH4 Toluene Toluene Column Benzene Column Stabilizer Diphenyl 14 12 10

  6. Cost Function and Constraints • The following profit is maximized (Douglas’s EP): -J = pbenDben – ptolFtol – pgasFgas – pfuelQfuel – pcwQcw – ppowerWpower - psteamQsteam + Σ(pv,iFv,i) • Constraints during operation: • Production rate: Dben≥ 265 lbmol/h. • Hydrogen excess in reactor inlet: FHyd / (Fben + Ftol + Fdiph) ≥ 5. • Bound on toluene feed rate: Ftol ≤ 300 lbmol/h. • Reactor pressure: Preactor ≤ 500 psia. • Reactor outlet temperature: Treactor ≤ 1300 °F. • Quencher outlet temperature: Tquencher = 1150 °F. • Product purity: xDben ≥ 0.9997. • Separator inlet temperature: 95 °F ≤ Tflash ≤ 105 °F. • + Distillation constraints • Manipulated variables are bounded.

  7. Disturbances

  8. Optimization Benzene price Disturbance

  9. Optimization • 14 steady-state degrees of freedom • 10 active constraints: • Pure toluene feed rate (UB) • By-pass valve around FEHE (LB) • Reactor inlet hydrogen-aromatics ratio (LB) • Flash inlet temperature (LB) • Methane mole fraction in stabilizer bottom (UB) • Benzene mole fraction in stabilizer distillate (UB) • Benzene mole fraction in benzene column bottom (UB) • Toluene mole fraction in benzene column distillate (UB) • Toulene mole fraction in toluene column bottom (UB) • Diphenyl mole fraction in toluene column distillate (UB) • 1 equality constraint: 11. Quencher outlet temperature • 3 remaining unconstrained degrees of freedom.

  10. Optimization – Active Constraints Purge (H2 + CH4) Compressor Equality Furnace 11 2 H2 + CH4 Quencher Toluene Mixer FEHE Reactor 3 1 Cooler 8 6 10 Separator 4 Benzene CH4 Toluene Toluene Column Benzene Column Stabilizer Diphenyl 9 7 5

  11. Candidate Controlled Variables • Candidate controlled variables: • Pressure differences; • Temperatures; • Compositions; • Heat duties; • Flow rates; • Combinations thereof. • 138 candidate controlled variables might be selected. • 14 degrees of freedom. • Number of different sets of controlled variables: • 10 active constraints + 1 equality constraint leaving 3 DOF: • What should we do with the remaining 3 DOF? • Self-optimizing control!!!

  12. Analysis of the linear model • Select 3 of 127 candidate measurements. • Scale variables properly and linearize. • Find max σ(G3x3) by a branch-and-bound algorithm. • Alternatively, find max σ(G3x3·Juu-1/2) by a branch-and-bound algorithm. • Calculate the loss by the exact local method for the most important disturbance: Namely, feed toluene rate. “input scaling”

  13. Analysis of the linear model • All measurements: • σ(Gfull) = 0.0445; σ(Gfull·Juu-1/2) = 0.1695 I I II II III III II II III III II II III III

  14. Self-optimizing variables II W Purge (H2 + CH4) Compressor Furnace 11 2 H2 + CH4 Quencher Toluene Mixer FEHE Reactor III 1 F 1 III Q Cooler L 8 6 10 I Separator 4 Benzene CH4 Toluene Toluene Column Benzene Column Stabilizer I xtol Optimal set Diphenyl 9 7 5

  15. Analysis of the linear model b. Distillation train excluded and separator pressure constant (controllability): σ(Gfull) = 0.0440; σ(Gfull·Juu-1/2) = 0.1663 I I II II III III Note: The two methods give the same order in this case

  16. Alternative self-optimizing variables W II Purge (H2 + CH4) Compressor Furnace 11 2 H2 + CH4 Quencher Toluene Mixer FEHE Reactor P 1 1 III Cooler 8 6 10 Separator 4 Benzene CH4 Toluene xtol Toluene Column Benzene Column Stabilizer I Diphenyl 9 7 5

  17. Conclusion steady-state analysis • Many similar alternatives in terms of loss • Need to consider dynamics (Input-output controllability analysis): • RHP-zeros • RHP-poles • Input saturation • Easy of implementation (decentralized control of final 3x3 supervisory control problem)! • Now: Consider “bottom-up” design of control system

  18. Bottom-up design of control system • Start with stabilizing control • Levels • Pressure • Temperatures • Normally start with fastest loops (often pressure) • but let is start with levels

  19. “Bottom-up”: Proposed Control StructureStabilizing Control: Control 7 liquid levels Purge (H2 + CH4) Compressor Furnace H2 + CH4 Quencher Toluene Mixer FEHE Reactor Cooler LC Separator LC LC LC Benzene CH4 Toluene Toluene Column Benzene Column Stabilizer LC LC LC Diphenyl LV-configuration assumed for columns

  20. Avoiding “Drift” I – 4 Pressure loops Pressure with purge Purge (H2 + CH4) Compressor Furnace H2 + CH4 Quencher Toluene Mixer FEHE Reactor Cooler PC LC Separator PC PC PC LC LC LC Benzene CH4 Toluene Toluene Column Benzene Column Stabilizer LC LC LC Diphenyl Column pressures are given

  21. Avoiding “Drift” II – 4 Temperature loops Purge (H2 + CH4) Compressor Furnace H2 + CH4 Quencher Toluene Ts Mixer FEHE Reactor TC ps Cooler PC LC Separator PC PC PC LC LC LC Benzene CH4 Toluene TC TC Toluene Column Benzene Column Stabilizer TC LC LC LC Diphenyl

  22. Now suggest pairings for supervisory control

  23. Control of 11 active constraints. Purge (H2 + CH4) Compressor SP CC Furnace SP SP TC FC H2 + CH4 Quencher Toluene Ts Mixer FEHE Reactor TC ps FC Cooler PC SP LC SP TC Separator PC PC PC LC LC LC Benzene CH4 Toluene SP SP SP CC TC CC CC TC Toluene Column Benzene Column Stabilizer SP methane SP SP CC TC CC CC LC LC LC 3 DOF left Diphenyl

  24. Control of 3 self-optimizing variables: Optimal set II Purge (H2 + CH4) Compressor SP CC Furnace SP SP TC FC H2 + CH4 Quencher Toluene Ts Mixer FEHE Reactor III TC ps FC Q Cooler PC SP LC SP TC Separator PC PC PC LC LC LC Benzene CH4 Toluene SP SP SP CC TC CC CC TC Toluene Column Benzene Column Stabilizer SP methane SP SP CC TC CC CC LC LC LC Supervisory control problem seems difficult I xtoluene Diphenyl

  25. Control of 3 self-optimizing variables:Alternative set: SIMPLE II Purge (H2 + CH4) Compressor SP CC Furnace SP SP TC FC H2 + CH4 Quencher Toluene Ts III Mixer FEHE Reactor TC ps FC Cooler PC SP LC SP TC Separator PC PC PC LC LC LC Benzene CH4 Toluene SP SP SP CC TC CC CC TC I Toluene Column Benzene Column Stabilizer xtol SP SP SP CC TC CC CC LC LC LC Could control another composition, e.g. quencher outlet diphenyl Diphenyl

  26. Conclusion HDA • Follow systematic procedure • May want to keep several candidate sets of “almost” self-optimizing variables • Final evaluation: Non-linear steady-state simulations + Dynamic simulations using Aspen (ongoing!)

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