PhD presentation

Transcript PhD presentation

Plantwide Control for Economically Optimal Operation of Chemical Plants

- Applications to GTL plants and CO 2 capturing processes Mehdi Panahi PhD defense presentation December 1 st , 2011 Trondheim M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 1

Outline

 Ch.2 Introduction  Ch.4 Economically optimal operation of CO 2 capturing process; selection of controlled variables  Ch.5 Economically optimal operation of CO 2 capturing process; design control layers  Ch.6 Modeling and optimization of natural gas to liquids (GTL) process  Ch.7 Self-optimizing method for selection of controlled variables for GTL process  Ch.8 Conclusions and future works

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 2

Outline

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 3

Skogestad plantwide control procedure

I Top Down • Step 1: Identify degrees of freedom (MVs) • Step 2: Define operational objectives (optimal operation) – Cost function J (to be minimized) – Operational constraints • Step 3: Select primary controlled variables

CV1s (Self-optimizing)

• Step 4: Where set the production rate? (Inventory control) II Bottom Up • Step 5: Regulatory / stabilizing control (PID layer) – What more to control ( CV2s ; local CVs)?

– Pairing of inputs and outputs • Step 6: Supervisory control (MPC layer) • Step 7: Real-time optimization

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 4

Optimal Operation Mode I:

maximize efficiency

Mode II:

maximize throughput

Self-optimizing control

is when we can achieve acceptable loss with constant setpoint values for the controlled variables without the need to reoptimize the plant when disturbances occur

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 5

Selection of CVs: Self-optimizing control procedure

Step 3-1: Define an objective function and constraints Step 3-2: Degrees of freedom (DOFs) Step 3-3: Disturbances Step 3-4: Optimization (nominally and with disturbances) Step 3-5: Identification of controlled variables (CVs) for unconstrained DOFs Step 3-6: Evaluation of loss

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 6

Maximum gain rule for selection the best CVs

Let G denote the steady-state gain matrix from inputs u (unconstrained degrees of freedom) to outputs z (candidate controlled variables). Scale the outputs using S 1

L max = 1 2   1 1 1   -1/2 uu span(z )= max z -z i d,e i i,opt = max e d i,opt (d)+ max e e i

For scalar case, which usually happens in many cases, the maximum expected loss is:

L max = J uu 2 1 S G 1 2

Maximum gain rule is useful for prescreening the sets of best controlled variables M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 7

Exact local method for selection the best CVs

max. Loss= 1 2 σ(M) 2 1/2 uu y -1 n y -1 F=G J J -G uu ud d y F is optimal sensitivity of the measurements with respect to disturbances; F= Δy opt.

Δd

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 8

Applications of plantwide procedure to two important processes

1. Post-combustion CO 2 (Chapters 3, 4 and 5) capturing processes 2. Converting of natural gas to liquid hydrocarbons (Chapters 6 and 7) M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 9

Importance of optimal operation for CO 2 capturing process

Dependency of equivalent energy in CO 2 capture plant verses recycle amine flowrate An amine absorption/stripping CO 2 capturing process * *Figure from: Toshiba (2008). Toshiba to Build Pilot Plant to Test CO2 Capture Technology. http://www.japanfs.org/en/pages/028843.html

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 10

Gas commercialization options and situation of GTL processes A simple flowsheet of a GTL process M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 11

Outline

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 12

Economically optimal operation of CO 2 capturing Step 1.

min. (energy cost + cost of released CO 2 to the air)

Step 2.

Objective function

10 steady-state degrees of freedom Cooling Water in Water Make up V-9 V-8 Cooling Water out Cooler Absorber To Stack n=15 Pump 2 V-7 V-10 Surge Tank Amine Makeup Stripper Condenser V-3 CO2 V-4 Cooling Water in n=20 V-2 Cooling Water out Step 3.

3 main disturbances Rich/Lean Exchanger Step 4.

Optimization 4 equality constraints and 2 inequality Flue Gas from Power Plant n=1 Pump 1 n=1 Reboiler Steam V-5 Condensate V-6 2 unconstrained degrees of freedom;10-4-4=2 V-1 Steps 5&6.

Exact Local method

: The candidate CV set that imposes the minimum worst case loss to the objective function

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 13

Exact local method for selection of the best CVs 39 candidate CVs

- 15 possible tray temperature in the absorber - 20 possible tray temperature in the stripper - CO 2 recovery in the absorber and CO 2 content at the bottom of the stripper - Recycle amine flowrate and reboiler duty

Applying a bidirectional branch and bound algorithm for finding the best CVs The best self-optimizing CV set in region I:

CO 2 temperature of tray no. 16 in the stripper recovery (95.26%) and

These CVs are not necessarily the best when new constraints meet M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 14

Optimal operational regions as function of feedrate Region I.

Nominal feedrate

Region II.

Feedrate >+20%: Max. Heat constraint

Region III.

Feedrate >+51%: Min. CO 2 recovery constraint

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 15

Proposed control structure with given flue gas flowrate (region I) M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 16

Region II: in presence of large flowrates of flue gas (+30%) Flowrateof flue gas (kmol/hr

)

Pumps duty (kW) Self-optimizing CVs in region I Cooler Duty (kW) CO 2 recovery % Temperature of tray no. 16 °C 106.9

Optimal nominal point +5% feedrate +10% feedrate +15% feedrate +19.38% feedrate, reboiler duty saturates +30% feedrate (reoptimized) 219.3

230.3

241.2

252.2

261.8

285.1

3.85

4.24

4.22

4.64

4.56

(+18.44%) 4.61

95.26

91.60

106.9

103.3

321.90

347.3

371.0

473.3

419.4 (+30.29%) 359.3

Saturation of reboiler duty (new operations region, region II); one unconstrained degree of freedom left Maximum gain rule for finding the best CV: 37 candidates

Temp. of tray no. 13 in the stripper: the largest scaled gain

Reboiler duty (kW) Objective function (USD/ton) 1161 1222 1279 1339 1393 (+20%) 1393 2.49

2.49

2.50

2.65

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 17

Proposed control structure with given flue gas flowrate (region I) M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 18

Proposed control structure with given flue gas flowrate (region II)

Reboiler duty at the maximum

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 19

Region III: reaching the minimum allowable CO 2 recovery Optimal nominal case in +30% feedrate +40% feedrate +50% feedrate +52.78% feedrate, reach to minimum CO 2 recovery

Flowrate of flue gas (kmol/hr)

285.1

Pumps Duty (kW) CO 2 recovery % Self-optimizing CV in region II

4.61

91.60

Temperature of tray 13 °C

109 307.02

328.95

335.1

4.58

4.55

4.54

86.46

81.31

80 109 109 109

Cooler Duty (kW) Reboiler Duty (kW) Objective function (USD/ton)

359.3

315.5

290.3

284.6

1393 1393 1393 1393 2.65

2.97

3.31

3.39

A controller needed to set the flue gas flowrate

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 20

Outline

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 21

Design of the control layers Regulatory layer

: Control of secondary (stabilizing) CVs (CV2s), PID loops • Absorber bottom level, • Stripper (distillation column) temperature, • Stripper bottom level, • Stripper top level, • Stripper pressure, • Recycle surge tank: inventories of water and amine, • Absorber liquid feed temperature.

Supervisory (economic) control layer:

Control of the primary (economic) CVs (CV1s), MPC • CO 2 recovery in the absorber, • Temperature at tray 16 in the stripper, • Condenser temperature.

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 22

RGA analysis for selection of pairings 1. Dynamic RGA Recycle CO 2 recovery

G dyn.

(s)=

Temp. no.16 in the amine

     6.85s+1.74

2 19.7s +11.4s+1 (-9.51s-1.02)e -2s

stripper Reboiler duty

0.45s+0.0754

    RGA dyn.

(0)=   0.77 0.23

0.23 0.77

  4.5

4 3.5

3 2.5

2 1.5

1 0.5

0 10 -3 Off-diagonal pairing alt.2

Diagonal pairing alt.1

10 -2 10 -1 10 0 Frequency [rad/min]

2. Steady-State RGA

G SS  10  2      0.5232 1.48

 8.47

5.17

  RGA = SS    0.27

 1.27

 0.27

  10 1

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’

10 2

”Break through” of CO 2 at the top of the absorber (UniSim simulation) 0,055 0,05 0,045 0,04 0,035 0,03 0,025 0,02 0,015 0 50 Liquid mole fraction of CO2 in trays of the Absorber 100 150 200 250 Time (min) 300 350 400 tray 1 tray 15 450

tray 15 tray 14 tray 13 tray 12 tray 11 tray 10 tray 9 tray 8 tray 7 tray 6 tray 5 tray 4 tray 3 tray 2 tray 1

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 24

Proposed control structure with given flue gas flowrate, Alternative 1 M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 25

Proposed control structure with given flue gas flowrate, Alternative 2 (reverse pairing) M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 26

Proposed control structure in region II, Alternative 3 M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 27

Modified alternative 2 Modified Alternative 2: Alternative 4 M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 28

Control of self-optimizing CVs using a multivariable controller M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 29

Performance of the proposed control structure, Alternative 1 M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 30

Performance of the proposed control structure, Alternative 3 M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 31

Performance of the proposed control structure, Alternative 4 M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 32

Performance of the proposed control structure, MPC M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 33

Comparison of different alternatives

• Alternative 1 is optimal in region I, but fails in region II • Alternative 2 handles regions I (optimal) and II (close to optimal), but more interactions in region I compare to Alternative 1. No need for switching • Alternative 3 is optimal in region II. Need for switching • Alternative 4 is modified Alternative 2 ,results in less interactions. No need for switching • MPC, similar performance to Alternatives 2 & 4

Alternative 4 is recommended for implementation in practice M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 34

Outline

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 35

CH 4

A simple flowsheet of GTL process

CO+H 2 +CH 4 CO 2 CO+H 2 (CH 2 ) n (CH 2 ) n

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 36

Pre-reformer reactions Converting higher hydrocarbons than methane, For Methanation Shift Reaction

 2

C H +nH O n m 2 m

) 2

H +nCO 2 CO+3H 2



2 CO+H O 2



CO +H 2 2

Auto-thermal reformer (ATR) reactions Oxidation of methane: Steam reforming of methane:

4  3  2

2 

 2

H O

2 

 3

2 Shift Reaction:



H O

2 

2 

2 Fischer-Tropsch (FT) reactions

nCO

 2

2  (-

2 -)



nH O

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 37

Fischer-Tropsch (FT) reactor

Simulation of a slurry bubble column reactor (SBCR) Reactions:

nCO

 2

2  (-

2 -)



nH O

2 Kinetics (the model developed by Iglesia et al):

r CH

4   8

P P H

   5

P CO

0.05

(

mol CH

4 g-atom surface metal. s )

r CO

 1.96 10  8

P H

0.6

2  5

P CO P CO

0.65

(

mol CO

g-atom surface metal. s ) FT products distribution (ASF model):

w n



(1  )

 1 41 reactions: 21 reactions for C n H 2n+2 and 20 reactions for C n H 2n FT products: C 1 , C 2 , C 3 -C 4 (LPG), C 5 -C 11 (Naphtha, Gasoline), C 12 -C 20 (Diesel), C 21+ (wax)

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 38

Detailed flowsheet of GTL process (UniSim) Natural Gas 8195 kmol/hr 3000 kPa, 40 ° C CH 4 :0.955

C 2 H 6 :0.030

C 3 H 8 :0.005

n-C 4 H 10 :0.004

N 2 :0.006

Steam Pre-Reformer Natural Gas (fuel) 5204 kmol/hr 455 ° C, 3000 kPa 2937 kmol/hr 3000 kPa 200 ° C 675 ° C Pure Oxygen (from ASU) 5236 kmol/hr Compressor CO 2 936 kmol/hr Water 9736 kmol/hr Heater 210 ° C 2000 kPa FT Reactor Vapor 210 ° C MP Steam Water MP Steam Fired Heater 1030 ° C Autothermal Reformer (ATR) MP Saturated Steam 48887 kmol/hr 252 ° C, 4000 kPa Syngas 35100 kmol/hr Extra Steam 27020 kmol/hr To fired heater (not shown) to produce superheat steam Water 38 ° C 4743 koml/hr 16663 kmol/hr 400 ° C, 4000 kPa ASU Turbines Superheated Steam 10 kPa Liquid Tail Gas 12673 kmol/hr CH 4 : 0.3587

CO: 0.1454

H 2 : 0.2142

N 2 : 0.1173

CO 2 : 0.0838

H2O: 0.0024

Other: 0.0782

Light Ends Purge to fired heater (as fuel) 392 kmol/hr 3-Phase Separator To Upgrading Unit 721 kmol/hr, 30 ° C (141.4 m 3 /hr) vol. mole fraction LPG: 0.0468

Naphtha/Gasoline: 0.3839

Diesel: 0.3216

Wax: 0.2242

Other: 0.0235

Water 8589 kmol/hr M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 39

Different methods for calculation of α

1) 2) α α 1 2

r CH

r CO

 0.55

0.4

0.6

 (1   ) 2 16  

16 1 

1.6

 1 2

2 ( 1   2 ) )( 1 1    1 30   1  

1  2000 1 28  

25 ( 1 1   (1   ) 2  2000 1 30  1    1 28   3  2 1 352   1   1 350 )]   ...20

 19 )(1   2 ) ( 1 1   1 352   1   1 350 )]

1.1

1 0.9

0.8

1.5

1.4

1.3

1.2

0.7

0.02

0.03

0.04

0.05

Left hand side value in 6.12

0.06

0.07

Real roots α as a function of the selectivity (1.2 ≤ H 2 /CO ≤ 2.15)   (0.2332

y CO y CO



y H

2  

 533)] 3) Constant α = 0.9 (α 3 )

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 40

FT reactor performance (single pass) at H 2 /CO=2 feed for α 1 , α 2 and α 3

parameter CO conversion, % H 2 conversion, % CH 4 formation (kg/kgcat.hr) Other hydrocarbons formation (kg/kgcat.hr) α 1 83.56

90.98

0.0106

α 2 86.52

91.82

0.011

α 3 86.97

91.97

0.011

0.1877

0.1924

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 41

FT Products distribution when α 1 is used M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 42

FT Products distribution when α 2 is used M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 43

FT Products distribution when α 3 is used M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 44

Dependency of different α calculation methods vs. feed H 2 /CO M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 45

Optimization formulation Objective function

Variable income (P): sales revenue – variable costs

Steady-state degrees of freedom

H 2 O/C (fresh + recycled hydrocarbons to pre-reformer) O 2 /C (hydrocarbons into ATR) 3. Fired heater duty 4. CO 2 recovery percentage 5. Purge ratio 6. Recycle ratio to FT reactor

Operational constraints

Molar ratio H 2 O/C ≥ 0.3

ATR exit temperature ≤ 1030ºC, active at the max.

3. Inlet temperature to ATR ≤ 675ºC, active at the max.

The purge ratio is optimally around 2%, it is bounded at a higher value, active at the min.

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 46

Optimality of objective function ( α 2 model) with respect to decision variables and active constraints, wax price= 0.63 USD/kg M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 47

Outline

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 48

Optimal Operation of GTL process

Mode I:

Natural gas is given -

Mode II:

Natural gas is also a degree of freedom (maximum throughput)

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 49

Process flowsheet of GTL process with data for optimal nominal point (mode I) Natural Gas 8195 kmol/hr CH 4: 0.955

C 2 H 6 :0.030

C 3 H 8 :0.005

n-C 4 H 10 :0.004

N 2 :0.006

Steam Pre-Reformer Natural Gas (fuel) 4896 kmol/hr 455 °C, 3000 kPa 455 °C 3000kPa 2500kPa Tail Gas 10846 kmol/hr Compressor I 0.47 MW Fired Heater Oxygen(from ASU) 200 °C 675 °C 5190 kmol/hr Water 1030 °C Autothermal Reformer (ATR) HP Saturatied Steam Syngas 34627 kmol/hr CO 2 906 kmol/hr Compressor II 0.55 MW Recycle Tail gas to FT 7762 kmol/hr 2700 kPa 3000 kPa 2700 kPa V-1 ∆P=f(Q) Water 906 kmol/hr Heater 210 °C Vapor 210 °C 2500 kPa FT Reactor MP Steam Water Steam Drum Liquid Light Ends Purge to fired heater (as fuel) 325 kmol/hr 3-Phase Separator Liquid fuels to upgrading unit: 751 kmol/hr (144 m 3 /hr) vol. mole fractions: LPG: 0.0525

Naphtha/Gasoline: 0.3759

Diesel: 0.3156 Wax: 0.2257

Water 8560 kmol/hr Extra Steam To fired heater (not shown) to produce superheat steam Superheated Steam ASU Turbines M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 50

Mode I: Natural gas flowrate is given Step 1: Define the objective function and constraints Variables income Inequality Constraints

Three inequality constraints + capacity constraints on the variable units; fired heater (duty +40% compared to nominal), CO plant (+20% oxygen flowrate).

2 recovery unit (+20% feedrate), oxygen

Equality Constraints (Specs:9) Step 2: Identify degrees of freedom (DOFs) for optimization 15 degrees of freedom M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 51

Step 3: Identification of important disturbances

Natural gas flowrate, Natural gas composition, Natural gas price, FT reactions kinetic parameter, Change in active constraints value.

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 52

Step 4: Optimization MIXED method

: combines the advantage of global optimization of BOX and efficiency of SQP method 15 degrees of freedom, 9 equality constraints and 3 active constraints:

unconstrained degrees of freedom:

15 – 9 – 3 = 3, which may be viewed as: H 2 O/C, CO 2 recovery, tail gas recycle ratio to FT reactor

Step 5. Identification of candidate controlled variables

18 candidate measurements including the three unconstrained degrees of freedom      18!

3!15!

 816

Step 6. Selection of CVs Exact local method for selection of the best CVs M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 53

Individual measurements (mode I) worst-case loss for the best 5 individual measurement sets

no.

Sets Loss (USD/hr) 1 2 3 4 5 y 3 :CO 2 recovery y y y 3 3 3 :CO :CO :CO 2 2 2 recovery recovery recovery y 10 :CH 4 mole fraction in fresh syngas y 9 : CO mole fraction in fresh syngas y 2 : H 2 O/C y 2 : H 2 O/C y 6 : H 2 /CO in tail gas y 6 : H 2 /CO in tail gas y 12 : CO mole fraction in tail gas y 6 : H 2 /CO in tail gas y 5 : H 2 /CO in fresh syngas y 5 : H 2 /CO in fresh syngas y 5 : H 2 /CO in fresh syngas 1393 1457 1698 2594 2643

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 54

Control structure for mode I of operation with proposed CVs and possible pairings with MVs (red lines are by-pass streams) Natural Gas TC SP=455 °C Steam Pre-Reformer Natural Gas (fuel) CC SP (H 2 O/C) Fired Heater TC SP=200 °C TC SP=675 °C SP=30 bar PC Compressor I Autothermal Reformer (ATR) Self-optimizing CV2 SP (CO 2 recovery%)= 75.73

CC Oxygen(from ASU) TC SP=1030 °C Water SP=38 °C TC Syngas Self-optimizing CV1 SP (CO%)=25.67 CC HP Saturated Steam CO 2 SP=27bar PC V-1 ∆P=f(Q) Water RC SP (splitter ratio) Tail Gas Self-optimizing CV3 SP (CO%)=12.96 gas to FT SP=12.5bar

Vapor

Heater TC SP=210 °C FT Reactor Water Steam Drum Liquid 3-Phase Separator SP=3% purge MP Steam RC Light Ends Purge to fired heater (as fuel) Liquid fuels to upgrading unit TC SP=30 °C Water SP=455 °C TC Extra Steam To fired heater (not shown) to produce superheat steam Superheated Steam ASU Turbines M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 55

Mode II: Natural gas feedrate is also a degree of freedom

Point A: oxygen flowrate saturates 1 extra DOF, 1 new active constraint

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 56

Optimal values in: nominal point, saturation of oxygen flowrate and maximum throughput opt. nominal max. oxygen max. through put H 2 O/C 0.6010

0.5357

0.4084

O 2 /C 0.523

0.516

0.504

CO 2 recovery Recycle ratio to FT Purge of tail gas H 2 /CO fresh 75.73% 73.79% 3% 2.1

76.80% 76.04% 90% 97.13% 3% 3% 2.092

2.095

H 2 /CO into FT 2.03

1.91

1.80

CO conversion % per pass 85.74

67.08

51.25

overall 95.50

94.14

94.79

H 2 conversion % per pass 89.93

60.69

overall 96.92

74.705

95.88

96.39



0.87

0.86

0.87

Carbon efficiency Objective function (USD/hr) 74.59% 74.30% 74.31% 49293 59246 59634 FT reactor volume is the bottleneck M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 57

Individual measurements (mode II) worst-case loss for the best 5 individual measurement sets

no.

Sets Loss (USD/hr) 1 2 3 4 5 y 3 :CO 2 recovery y 3 :CO 2 recovery y 3 :CO 2 recovery y 3 :CO 2 recovery y 3 :CO 2 recovery y 2 : H 2 O/C y 2 : H 2 O/C y 2 : H 2 O/C y 2 : H 2 O/C y 9 : CO mole fraction in fresh syngas y 7 : H 2 /CO into FT reactor y 6 : H 2 /CO in tail gas y 5 : H 2 /CO in fresh syngas y 17 : tail gas flowrate to syngas unit y 15 : CO mole fraction into FT reactor 3022 3316 3495 4179 4419

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 58

Control structure for mode II of operation with proposed CVs and possible pairings with MVs (red lines are by-pass streams) Steam Natural Gas Pre-Reformer TC SP=455 °C Natural Gas (fuel) Fired Heater SP=30 bar PC RC SP (splitter ratio) Compressor I Autothermal Reformer (ATR) Self-optimizing CV2 SP (CO 2 recovery%)= 76.04

Oxygen(from ASU) flowrate is at max.

TC SP=200 °C CC CO 2 SP=27bar PC Compressor II Recycle Tail gas to FT Tail Gas CC Self-optimizing CV3 SP (H2/CO)=1.8 SP=12.5bar

Vapor

FT Reactor TC SP=675 °C TC SP=1030 °C Water SP=38 °C TC Syngas V-1 ∆P=f(Q) Heater TC SP=210 °C Water Steam Drum Liquid 3-Phase Separator SP=3% purge MP Steam RC Light Purge to fired Ends heater (as fuel) Liquid fuels to upgrading unit TC SP=30 °C Water CC Self-optimizing CV1 SP (H 2 O/C=0.4084) HP Saturated Steam Water SP=455 °C TC Extra Steam To fired heater (not shown) to produce superheat steam Superheated Steam ASU Turbines M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 59

Concluding remarks of self-optimizing application for GTL process

• Self-optimizing method was applied for selection of the CVs for GTL • There are 3 unconstrained DOFs in both modes of operation • One common set in the list of the best individual measurements in two modes: CO 2 recovery H 2 /CO in fresh syngas H 2 O/C setpoint reduces from 0.6 to o.4

• Operation in Snowballing region should be avoided • Saturation point of oxygen plant capacity is recommended for operation in practice

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 60

Outline

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 61

Conclusions and future works

Systematic plantwide procedure of Skogestad was applied for a post-combustion CO 2 capturing process; a simple control configuration was achieved, which works close to optimum in the entire throughput range without the need for switching the control loops or re-optimization of the process A GTL process model suitable for optimal operation studies was modeled and optimized. This model describes properly dependencies of important parameters in this process Self-optimizing method was applied to select the right measurements for the GTL process in two modes of operation UniSim/Hysys linked with MATLAB showed to be a very good tool for optimal operation studies

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 62

Conclusions and future works

Implementation of the final control structure for CO 2 for implementation in practice capture plant is recommended Dynamic simulation of the GTL process should be done to validate the proposed control structures The application of plantwide control procedure is strongly recommended for other newer energy-intensive processes Developing a systematic method for arriving at a simple/single control structure, which works close to optimum in all operational regions can be a good topic for future work

Thank you for your attention!

M. Panahi ’Plantwide Control for Economically Optimal Operation of Chemical Plants’ 63

PhD presentation

Transcript PhD presentation

Plantwide Control for Economically Optimal Operation of Chemical Plants

Outline

Outline

Skogestad plantwide control procedure

Optimal Operation Mode I:

maximize efficiency

Mode II:

maximize throughput

Selection of CVs: Self-optimizing control procedure

Maximum gain rule for selection the best CVs

Exact local method for selection the best CVs

Applications of plantwide procedure to two important processes

Outline

Outline

Outline

Outline

Outline

Thank you for your attention!

Directory