Experience on System Integration and Simulation
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Transcript of Experience on System Integration and Simulation
EXPERIENCE ON SYSTEM INTEGRATION AND SIMULATION
Professor RUBENS MACIEL FILHO
•Laboratory of Optimization, Design and Advanced Process Control
•Department of Chemical Processes, School of Chemical• Engineering, State University of Campinas, Campinas - Brazil
e-mail [email protected]
VIRTUAL SUGAR CANE BIOREFINERY-CTBE - August 2009
Universidade Estadual de Campinas- UNICAMPSchool of Chemical Engineering
MOTIVATION• Process Simulation
– Evaluation of several possible routes –routes discrimination
–Investigation of different scenarios
- Process understanding
- Impact of operation variables on processperformance
Process Simulation (cont.)-Preliminary evaluation of costs, water andenergy consumption
-Studies of variable interaction and processdynamics
-Operator Training
-Dynamic simulation- process control strategiesmay be evaluated
Design of Equipments and plant conceptualdesign
PROCESS MODELLINGSteady State Model
Dynamic Model
Simplified versus Detailed Model
Physico-Chemisty based Models (Deterministic) versus Empiric and or Statistical Models
Hybrid Model
Single Unit Models
Large Scale Plant Model
Process Simulation
System –can be seen a set of subsystem depending upon of required investigation
Interaction among subsystems – made through mass and heat transfer parameters
Subsystem 1– an important component of the process, inside an equipment where the phenomena are intrinsically taking place- for instance catalyst particle, bagasse to be hydrolyzed and microorganism in biotechnological process. When considered explicitly a heterogeneous model is formulated.
Subsystem 2 - Equipment - peace of the plant where the changes (reactions, mixtures or separations) are occurring. In this category it may be place reactors, separation columns, fermentors, etc.
Subsystem 3 – large scale plant or a set of equipments in which there exist interest to study
Subsystem 1 and 2 – normally require software development if detailed representation are desired.
Subsystem 3 – simulators, including the commercial ones (Hysis , Aspen, Gproms etc)
System Integration
There exist an incentive for high operational performance operation
Process optimization begins with better process control
Large Plant Optimization and controlRTO: Integrate economic objectives and control
Stability, controllability and safety
System Integration
Large Plant Optimization and Control
RTO (Real Time Operation): Integrate economic objectives and control
Stability, controllability and safety- may be expressed as plant restriction
Refinery process ⇒large scale units, high products output, monitoring difficulties,data reconciliation
Two main strategies are to be implemented:
One layer approach
two layers approach
Hybrid approach may be necessary
Optimization Strategies
Economical optimization problem is solved together with the control problem
very sensitive to model mismatch
dimension of the optimization problem can be very large
( on line applications can be restrictive) use of simplified model may not be suitable
One layer approach
One layer approach
non-measuredinputs
measured inputs Process
non-measuredoutputs
measuredoutputs
Estimationblock
controller/optimizer
hierarchical control structure where there is an optimization layer that calculates set-
points to the advanced controller
the optimization layer is composed of an objective function and a process steady-
state model
Two layers approach
Two layer approach
non-measuredinputs
measured inputs measured outputs
non-measuredoutputs
Process
Estimationblock
Controller
Optimizer
setpoints
Advanced Controllers
• CONTROLADORES LINEARES• NON LINEAR CONTROLLERS• PREDICTIVE CONTROLLERS• ROBUST CONTROLLERS• ADAPTIVE CONTROLLERS• HYBRID CONTROLLERS (NEURALNETWORK AND FUZZY COUPLED WITH MODELBASED CONTROLLER)
Simulation – Applications
Subsystem 1
STRUCTURED MATHEMATICAL MODEL
FOR ETHANOL PRODUCTION
Possible to handle with substrate to drive the fermentation
STRUCTURED MATHEMATICAL MODEL
Representative Metabolic Route (F. Lei et al. Journal of Biotechnology 88 (2001) 205-221)
Mass balance equations and reaction rate of the model
( ) ( )eglufeedeglu SSDXRR
tS
cos71cos −++−=
∂
∂
( ) ( )pyruvatepyruvate SDXRRR
tS
−−−=∂
∂321978.0
( ) adeacetaldehyedeacetaldehyiglu
egluea
heglu
egluha
leglu
eglul Xs
KsKss
kXKs
skX
Kss
kR11
cos1
1cos
cos1
1cos
cos11 1 ++
++
++
=
aeglu
eglu XKs
skR
7cos
cos77 +
=
aegluipyruvate
pyruvate XsKKs
skR
11
cos2222 ++
=
apyruvate
pyruvate XKs
skR
34
4
33 +=
( ) ( )deacetaldehydeacetaldehy SDXRRR
tS
−−−=∂
∂6435.0
( ) ( )acetateacetate SDXRRRt
S−−−=
∂∂
854363.1
Acdhadeacetaldehy
deacetaldehy XXKs
skR
444 +
=
aethanolrdeacetaldehy
ethanolrdeacetaldehy XsKKs
skskR
66
666 ++
−=
aegluieacetate
acetateea
acetate
acetate XsKKs
skX
Kss
kR1
1
cos555
555 ++
++
=
aegluieacetate
acetate XsKKs
skR
11
cos5588 ++
=
( ) ( )ethanolethanol SDXR
tS
−=∂
∂6045.1
( ) ( )XDXRRtX
−+=∂∂
87 619.0732.0
( ) ( ) aa XRRRRRR
tX
8710987 619.0732.0619.0732.0 +−−−+=∂
∂
aeglu
egluca
egluieethanol
ethanole
eglu
eglu XKs
skX
sKKss
kKs
skR
9cos
cos9
cos999
9cos
cos99 1
1+
++
++
+=
aeethanol
ethanolea
eglu
eglu XKs
skX
Kss
kR10
1010cos
cos1010 +
++
=
( ) ( ) AcdhAcdh XRRRRt
X87119 619.0732.0 +−−=
∂∂
AcdhXkR 1111 =
X → biomass; Xa → active cell material; XAcdh → Acetaldehyde dehydrogenase; D → dilution rate;
Ki → rate constant; Ki → affinity constant;Kji → inhibition constant
• Mass balance equations → 8
• Kinetic parameter → 37
• Parameter adjust → Genetic Algorithm
CSTR simulations
TRS → Total Reductor Sugars
Batch simulations
SugarGlycose
Sacarose
FER
MEN
TATI
ON Etanol
Ácido acético
Ácido lático
Acetona Butanol Etanol
CH
EMIC
AL
SYN
THES
IS
Acetaldeído
Ácido acético
Anidrido acético
Acetato de etila
Acetato de vinila
Crotonaldeído
Paraldeído
Butanol
Acetato de butila
Piridina
Nicotinamida
Glicol
Butadieno
Glioxalato
Produtos químicos produzidos por fermentação
Some Chemical Products via fermentation
Etileno
Etanol
Acetaldeído
Ácido acético
Propano
Propileno
Ácido acrílico
Glicerol
Ácido lático
Butadieno
Butanodiol
Ácido succínico
BIOMASS H
YD
RO
LYSI
SSugar
Glicose
Sacarose
Xilose
Arabinose
FER
MEN
TATI
ON
Produção de novos produtos químicos a partir de biomassa
Other Products to be obtained from biomass
Fermentation process – piuvirate is formed in glycolysys
GLICOSE
Glicose 6-fosfato
Frutose 6-fosfato
Frutose 1,6-bifosfato
Gliceraldeído 3-fosfato 1,3-Difosfoglicerato 3-fosfoglicerato
2-fosfoglicerato Fosfoenolpiruvato PIRUVATO
ADP ATP
ADP ATP
ATP ADP
ATP ADP
NAD+ NADH +Pi +H+
10'1 146
2 2 Piruvato 2 2cos−
++
−=∆
++→+
kJmolG
HNADHNADeGli
10'1
2
61
2 2 2 2−=∆
+→+
kJmolG
OHATPPATP i
Processo de glicólise
Condições anaeróbiasCondições anaeróbias
2 Lactato
GLICOSE
2 Piruvato
2 Etanol + 2CO2
2 Acetil CoA
4 CO2 + H2 O
Rota (EMP)10 reações sucessivas
O2
CO2
Condições aeróbias
Ciclo do ácido TCA
O2
2 Ácido Acrílico + 2H2O
Rota glicolítica
Metabolic pathways for the synthesis of acrylic acid (Straathof et al., 2005)
STRUCTURED MODEL WITH IMOBILIZED CELLS
Structured Models based on the work of Lei et al. (2001) e Stremel (2001).
Model of Lei et al. (2001) -a structured biochemical modelthat describes the aerobic growth of Saccharomycescerevisiae in a medium limited to glucose and / or ethanol.
Model of Stremel (2001) -alternative structured model torepresent the dynamic simulation of a tubular bioreactorwith immobilized cells of Saccharomyces cerevisiae foralcoholic fermentation.
Para desenvolvimento deste modelo foi considerado:
Continuous isothermal process
heterogeneous model ;
biomass composition: CH1,82O0,576N0,146;
spherical particles ;
heterofermentative processproduction associated with cell growth;
axial dispersion .
Solution by orthogonal collocation
Metabolic route
Model Reactions
aa
aa XKS
SkXKS
SkR1
11
11 ++
+=
a
i
X
KLKS
SkR
+
+=
2
222
1
1
aXKP
PkR3
33 +=
aXKL
LkR4
44 +=
ai
XSKKL
LkR
++
=55
55 11
aia
XAAKKL
LKS
SkR
+
+
++
=1
1
66666
aa
aa XKAA
AAkXKS
SkR
+
+
+
=7
77
77
Reaction Rates
Mass Balances for the solid phase Glicose
Piruvato
Lactato
Ácido Acrílico
Células
Células ativas
Enzima lactato desidrogenase
( ) XeRRrSr
rrR
Dt
S AAAKAS −+−
∂∂
∂∂
=∂
∂21
222
1
( ) XeRRrPr
rrRD
tP AAAKAP −−+
∂∂
∂∂
=∂∂
312
22 978,01
( ) XeRRRrLr
rrRD
tL AAAKAL −−−+
∂∂
∂∂
=∂∂
5432
22 023,11
( ) ( ) XeRRr
AArrrR
Dt
AA AAAKAAA −−+
∂∂
∂∂
=∂
∂74
222 8,01
( ) XkeX
XXRRtX
dAAAK
sat−
−+=
∂∂ − `
52 1821,0732,0
( ) ( ) aa XRRRRRR
tX
527652 821,0732,0821,0732,0 +−−−+=∂
∂
( ) LADHLADH XRRR
tX
526 821,0732,0 +−=∂
∂
Mass Balance for the Fluid Phase Glicose
Piruvato
Lactato
Ácido Acrílico
( )[ ]XeRRzSu
zSD
dtS AAAK
az−+
−−
∂∂
−
∂
∂=
∂212
2 1 ηε
ε
( )[ ]XeRRzPu
zPD
dtP AAAK
az−−
−+
∂∂
−
∂
∂=
∂312
2978,01 η
εε
( )[ ]XeRRRzLu
zLD
dtL AAAK
az−−−
−+
∂∂
−
∂
∂=
∂5432
2023,11 η
εε
( )[ ]XeRRz
AAuzAAD
dtAA AAAK
az−−
−+
∂∂
−
∂
∂=
∂742
28,01 η
εε
SIMULATION RESULTS
0 2 4 6 8 10 12 14 16 18 20 22 24 26 28 30 32
60
75
90
105
120
135
150
Tempo (h)
Conc
entra
ção d
e Glic
ose (
kgm
-3)
0
5
10
15
20
25
30
Concentração de Ácido Acrílico (kgm-3)
0 2 4 6 8 10 12 14 16 18 20 22 24 26 28 30 320,00
0,75
1,50
2,25
3,00
3,75
4,50
5,25
6,00
Tempo (h)
Conc
entra
ção d
e Piru
vato
(kgm
-3)
0,0
0,5
1,0
1,5
2,0
2,5
3,0
3,5
4,0 Concentração de Lactato (kgm-3)
0,0 0,2 0,4 0,6 0,8 1,0
60
75
90
105
120
135
150
Posição axial
Conc
entra
ção d
e Glic
ose (
kgm-3 )
0
5
10
15
20
25
30 Concentração de Ácido Acrílico (kgm-3)
Simulation – Applications
Subsystem 2
Multitubular Catalytic Reactor
Tube-side : catalytic fixed bed
Detailed modeling
where: A = Parallel flow in the baffles holes B = Flow near the baffle endC = Parallel flow in the space between bundle of tubes and shellD = Flow between baflles and shellE = Cross Flow in the window zones
Multitubular Fixed Bed Catalitic Reactor
Co-current Design Alternative Design
Temperature Profiles
Radial mean temperature profile along the reactor length for different reactor configurations
Heat Transfer Coefficient Profiles
Co-current Design
Alternative Design
RHYDROLY
REACTOR DESIGN FOR HYDROLYSE
Adsorption
Enzymes(Cellulase,
β-glucosidase)
R3
R2
R1
AdsorbtionCellulase on cellulose and lignin, β-Glucosidase on ligninR1Cellulose to Cellobiose (Catalized by cellulase adsorbed on cellulose)R2Cellulose to Glucose (Catalized by cellulase adsorbed on cellulose)R3Cellobiose to Glucose (Catalized by non-adsorbed β-Glucosidase)
REACTION SYSTEM
312 056.1 rr
dtdG
−=
32 053.1111.1 rrdtdG
+=
21 rrdtdC
−−=
Cellulose
Cellobiose
GlucoseFig. 1 Observed time course of glucose (G) andcellobiose (G2) profiles. Enzymatic hydrolysisof AHP-pretreated sugarcane bagasse at differentinitial solid loadings (% w/w).
0
5
10
15
20
25
0 12 24 36 48 60 72Time (h)
Glu
cose
[G
] -
Cel
lob
iose
[G
2]
(g/L
)
G-1% G-3% G-5% G2-1% G2-3% G2-5%
EXPERIMENTAL DATA ANDMASS BALANCES
REACTION SCHEMES
Three reaction Scheme(General)
Two reaction Scheme(No direct glucose formationfrom cellulose)
One reaction Scheme(Nor direct glucose formationfrom cellulose neithercellobiose accumulation)
MATHEMATICAL MODELING
Enzyme adsorption on cellulose and lignin• One site Langmuir isotherm• Two sites Langmuir Isotherm
Enzyme inhibition by cellobiose and cellulose• Competitive• Non-competitive
Recalcitrance• Substrate reactivity• Substrate susceptibility
Enzyme deativation (Thermal, mechanical)
• First order kinetic
Non-mechanistical, fit experimental data,most used in the literature
Both are used in the literature. There is no consensus
α(S/S0)n+cte (S:substrate)v=v0Exp(-Krec(1-(S/S0))) (v0:adsorbed enzyme)
Very important for design of continuous reactionsystems at industrial scale
EXPERIMENTAL PROCEDURE AND KINETIC PARAMETER ESTIMATION
Adsorption• Enzyme adsorption on pretreated substrate• Enzyme adsorption on hydrolyzed substrate• Enzyme adsorption on ligninHydrolysis• Hydrolysis of pretreated substrate• Hydrolysis of partially hydrolyzed susbtrate• Hydrolysis with backgrond sugars (Cellobiose, glucose)• Fed batch (enzyme and susbtrate) hydrolysisParameter estimation with global and local optimization techniques• Genetic algorithms + quasi Newton• Simulated annealing + quasi Newton• Particle swarm method + quasi NewtonModel validation
Enzyme Loading5 FPU –CBU/g cellulose
500 FPU –CBU/g cellulose
Substrate Loading1%(W/W) 8%(W/W)
CONTINUOUS REACTION SYSTEMS I
CSTR
•Continuous substrate and enzyme feeding
n-CSTR
Continuous substrate and enzyme feeding at the first tank
n-CSTR with distributed feeding
•Ad hoc distributed feeding strategy of substrate and/or enzyme
•Model-based distributed feeding strategy of substrate and/or enzyme
Goals•Subs conc.•Subs conv.•Enzy consump.•Power Consump.•Resid time
CONTINUOUS REACTION SYSTEMS II
PFR with or without side feeding
Bafled PFR with or without side feeding
•Continuous substrate and enzyme feeding
•Ad hoc side feeding strategy or model-based side feeding strategy of substrate and/or enzyme
Goals•Subs conc.•Subs conv.•Enzy consump.•Power Consump.•Resid time•Overcome viscositylimitations
λ λ
CONTINUOUS REACTION SYSTEMS III
+
Liquefactor
Goals•Subs conc.•Subs conv.•Enzy consump.•Power Consump.•Resid time•Overcome viscositylimitations
Reactors•Liquefactor + n-CSTR•Liquefactor + PFR•Liquefator + Bafled PFR
REACTOR MODELING
n-CSTR Microfluid model
n-CSTR Macrofluid model• Ideal residence time distribution
• Substrate conversion
)()1(
i
iiRii Sr
SSV −== −
ϕτ
itni
n
en
ttE τ
τ/
1
)!1()( −
−
−=
∫∞→
=
=−
t
t Batchh
hsh dttE
ss
X0 0
)(1
PFR
CFD based model•Virtual tracerExperiments•Virtual determination ofRTD•Application of macrofluid model
)( h
hR
SrdSdV
−=ϕ
RESULTS FOR n-CSTRMacrofluid Model
10
20
30
40
50
60
70
80
90
100
110
120
0,650 0,670 0,690 0,710 0,730 0,750Xc
tao[
h]
NR=1 NR=2NR=3 NR=5NR=20 PFR
Microfluid Model
10
20
30
40
50
60
70
80
90
100
110
120
0,650 0,670 0,690 0,710 0,730 0,750Xc
tao[
h]N=1 NR=2NR=3 NR=5NR=20 PFR
Initial bagasse concentrationST0=50 g/L; initial cellulose concentrationSC0=40g/L.
Fig. 2 Total mean hydraulicresidence time (tao=τ) as afunction of cellulose conver-sion (Xc) predicted by the macrofluid and microfluidmodel.
CFD APPLIED TO REACTOR DESIGN I
ANSYS CFX (of Ansys Inc., EUROPE)xy velocity field Modeling approaches
Pseudo-homogeneous suspension with apparent rheological properties
‘or’
Multiphase
•Eulerian-Eulerian approach
•Eulerian-Lagrangian approach
CFD APPLIED TO REACTOR DESIGN IIBaffled PFR
Mesh details andPipe geometry
CFD APPLIED TO REACTOR DESIGN IIBaffled PFR
Predicted solids volume fraction distribution (1)and solid velocity (2)
1.
1.
2.
2.
HYDROTREATING OF MIDDLE DISTILLATES IN A TRICKLE BED REACTOR
The hydrodesulfurization (HDS), hydrodenitrogenation(HDN), hydrodeoxygenation, hydrocraking and saturativehydrogenation of middle distillates has been studied in thiswork.
An adiabatic diesel hydrotreating trickle bed packedreactor was simulated numerically by a heterogeneousmodel in order to check up the behaviour of this specificreaction system. Alternative design is proposed
The model consists of mass and heat balance equations for the fluid phase as well as for the catalyst particles, and take into account variations in the physical properties as well as of the heat and mass transfer coefficients. Heterogeneous model is developed
GAS in
LIQUID in
QUENCH
GAS out
LIQUID out
Bed 2
Bed 1
SHHnHydrocarboH2SnHydrocarbo 222 +=→+=
OHHnHydrocarboHOHnHydrocarbo 22 +−→+−
332 NHHnHydrocarboH3NnHydrocarbo +≡→+−
1 - Sulfur – containing hydrocarbons:
423 CHHnHydrocarboHCHnHydrocarbo +−→+−
2 - Oxygenated hydrocarbons:
3 - Nitrogenated hydrocarbons:
4- Hydrogenated hydrocrackable hydrocarbons:
5 - Unsaturated hydrocarbons with double bonds:
2HnHydrocarbo2HnHydrocarbo =→+
REACTOR PREDICTIONS
0 2 4 6 8 10650660670680690700710720730740750760770780
Tem
pera
ture
(K)
Bed length (m)
Figure 1 – Temperature profile along the reactor length.
0 2 4 6 8 100,0
0,1
0,2
0,3
0,4
0,5
0,6
0,7
0,8
0,9
1,0
Conv
ersi
on
Bed length (m)
Figure 2 – Sulfur conversion profile along the reactor length.
Pressure : 96 atm
0 2 4 6 8 10
650
655
660
665
670
675
680
685
690
695Te
mpe
ratu
re (K
)
Bed length (m)
Figure 3 – Temperature profile along the reactor length.
0 2 4 6 8 100,0
0,1
0,2
0,3
0,4
0,5
0,6
0,7Co
nver
sion
Bed length (m)
Figure 4 – Sulfur conversion profile along the reactor length.
Pressure: 68 atm
Efficient Mathematical Procedure for Calculating
Dynamic Adsorption Process
System for Adsorption Process
Different numericalmethods
Different equilibriumrelationships
Different operationalparameters,
and adsorbentcharacteristics
Different modelling approach
Feed Conditions:single adsorbate
binary or multicomponentcontinuos or pulse
Arrangement of the columns:fixed
in sequencysimulated moving bed
Equilibrium isothermsAdsorbent type and
characteristicsMass transfer model
Column parameters:dimensionsbed porosity
CONCENTRATION BREAKTHROUGH CURVESCONCENTRATION-DISTANCE PROFILES
MONOCOMPONENT ANDMULTICOMPONENT
ADSORBENT LOADING BREATHROUGH CURVESADSORBENT LOADING PROFILES
ELUTION CURVES (CHROMATOGRAPHY)
TYPES OF RESULTS
In the developed software:
•1 different numerical methods •1 different isotherms•1
were carried out in order to be possible to take decisions in relation to:
1 the evaluation of an operating adsorber 1 the possibility to apply this separation process for
recovering a given component from a mixture
Simulation of packed bed adsorption columnsusing the pore diffusion model, in which two masstransfer processes were considered: the external mass transfer from the bulk
liquid phase to the particle surface
internal pore diffusion within the adsorbent particle itself
Model and Solution
In the model formulation the following assumptions were made
• Diffusion coefficients independent of the mixture composition
• Spherical particles with equal sizes• Constant temperature and porosity• Not including axial dispersion
• Solution Procedure: orthogonal collocation method coupled with the DASSL routine
0 2000 4000 6000 8000 10000
0,0
0,20,40,6
0,81,0
1,2DELTA 200
c/c 0
t(s)
10 Elem 20 Elem 40 Elem 80 Elem Exper.
0 2000 4000 6000 8000 10000
0,0
0,2
0,4
0,6
0,8
1,0
1,2 DELTA 12.5
C/C
0
t (s)
10 Elem 20 Elem 40 Elem Experim.
0 2000 4000 6000 8000 10000
0,0
0,2
0,4
0,6
0,8
1,0
1,2 DELTA 25
C/C
0
t (s)
10 Elem 20 Elem 40 Elem Experim.
0 2000 4000 6000 8000 10000
0,0
0,20,40,6
0,81,0
1,2 DELTA 100
c/c 0
t(s)
10 Elem 20 Elem 40 Elem 80 Elem Exper.
Alternative Process Modeling
Fuzzy Logic
Artificial Neural Networks
Neuro Fuzzy
Hybrid Modeling
STATE UNIVERSITY OF CAMPINAS BRAZILDepartment of Chemical Engineering
SOFT SENSOR FOR MONITORING AND CONTROL OF AN INDUSTRIAL POLYMERIZATION PROCESS
OBJECTIVE:
To develop a Soft Sensor for polymer viscosity of an industrial PET Process.
PET Plant- the liquid phase (105.000 ton/year)
Figure 3- Schematic of virtual sensor.
RESULTS AND DISCUSSIONS
Input variable Name1 PE temperature T-12 SE temperature T-23 Temperature of the LP second stage T-44 Vacuum of the LP first stage P-15 Vacuum of the LP second stage P-26 HP temperature T-57 HP Vacuum P-38 Additive flow rate (catalyst). F-1
Output variable1 Measured viscosity by viscometer V-1
The variables, related to intrinsic viscosity, used for the neural net training are given in Table 1.Table 1- Variables for neural net training
0,980
0,990
1,000
1,010
1,020
0 4 8 13 17 21 25 29 33 38
Time (h)
Visc
osity
Viscosimeter Soft-Sensor
Figure 4 Viscosimeter versus Soft-Sensor (real time measurements - normalized values)
0,950
0,975
1,000
1,025
1,050
0 4 8 13 17 21 25
Time (h)
Visc
osity
Polymer viscosity Set-point
Figure 5. Process controlled using viscosity values estimated by Soft-Sensor (normalized values)
SETCIM INTEGRATION
(Industrial Test)
1340
1360
1380
1400
1420
1440
1460
1 6 11 16 21 26 31 36 41 46 51
Viscosímetro
Soft-
Sens
or
Viscosímetro Soft-Sensor
“Industrial Test”
R2 = 0.9086
1340
1360
1380
1400
1420
1440
1460
1340 1360 1380 1400 1420 1440 1460
Viscosímetro
Soft-
SEns
or
Soft-Sensor Linear (Soft-Sensor)
DATA DISPERSION (“Industrial test-several months
running ”)
H. POLIMERATION SCREEN OPERATION
HIGH POLIMERATION SCREEN OPERATION
Viscosimeter versus Soft-Sensor (Real Time Optimization)
Process Control by Soft-Sensor
Column Temperature- First Esterification Reactor
Extractive alcoholic fermentation process
•Usual existing processes: 3 or 4 tanks in series •Alternatives processes are under tests as flocculation and extractive
Flash
FilterFermentor
Vapour
Permeate
Feed
Purge
Return
Tf Pf
Ff
D pH TbT
Flash
FilterFermentor
Vapour
Permeate
Feed
Purge
Return
Tf Pf
Ff
D pH TbT
EXTRACTIVE FERMENTATION PLANT
Extractive Process
• This process was build up and validated for bioethanol production inbench scale by Atala (2004);
Development of Real-time State Estimators for Extractive Process - Introduction
- On-line monitoring by SS:- Allow real time monitoring of key variables of processes;
- Off-line monitoring:- Leads to time delay between sampling and results;- Requires advanced analytical instruments (including
near infrared spectrophotometers) → difficult to calibratedue to presence of CO2 in the media.
Software Sensor• Software sensor: an algorithm where several measurements are
processed together. The interaction of the signals from on-line instruments can be used for calculating or to estimate new quantities (e.g. state variables and model parameters) that cannot be measured in real-time.
• On-line measurements (input):- Temperatures;- Dilution rate;- pH;- Turbidity in the fermentor;- Pressure;- Feed flow rate in the flash vessel.• Off-line measurements (output):ethanol concentration in the fermentor and in the condensed stream from
the flash vessel.
Pf Ff Tf T D pH Tb
ANN-BASEDSOFT-SENSOR (1)
ANN-BASEDSOFT-SENSOR (2)
ESTIMATEDPferm
ESTIMATEDPflash
POTENTIAL INPUT VARIABLES
Pf Ff Tf T D pH Tb
ANN-BASEDSOFT-SENSOR (1)
ANN-BASEDSOFT-SENSOR (2)
ESTIMATEDPferm
ESTIMATEDPflash
POTENTIAL INPUT VARIABLES
ANN Structure Selection• Multilayer Perceptron (MLP) Neural Networks :- One of the most common ANN used in engineering;- understandable architecture and a simple mathematical form;• This NN consists of: input, output and one or more hidden
layers.• Numbers of neurons are N, M and K
fM(•)
θM
wM1
+
f2(•)
θ2
w21
+
f1(•)
θ1
w11
+
β1
W11
W12
W1M
x1
Input layer Hidden layer Output layer
xNw2N
w1N
wMN
F1(•)+
...
......
......
...
g1
fM(•)
θM
wM1
+
f2(•)
θ2
w21
+
f1(•)
θ1
w11
+
β1
W11
W12
W1M
x1
Input layer Hidden layer Output layer
xNw2N
w1N
wMN
F1(•)+
...
......
......
...
g1
θj
wj1
wj2
wjN
f(•)+...
yj
x1
x2
xN
θj
wj1
wj2
wjN
f(•)+...
yj
x1
x2
xN
(a) (b)
fM(•)
θM
wM1
+
f2(•)
θ2
w21
+
f1(•)
θ1
w11
+
β1
W11
W12
W1M
x1
Input layer Hidden layer Output layer
xNw2N
w1N
wMN
F1(•)+
...
......
......
...
g1
fM(•)
θM
wM1
+
f2(•)
θ2
w21
+
f1(•)
θ1
w11
+
β1
W11
W12
W1M
x1
Input layer Hidden layer Output layer
xNw2N
w1N
wMN
F1(•)+
...
......
......
...
g1
θj
wj1
wj2
wjN
f(•)+...
yj
x1
x2
xN
θj
wj1
wj2
wjN
f(•)+...
yj
x1
x2
xN
(a) (b)
Results and Discussion
• Even using on-line (input) datawith different levels of noise→The software sensor describedaccurately the ethanolconcentrations.
50
100
150
200
250
P f (m
mHg)
160
173
185
198
210
F f (L/
h)
32.5
33.3
34.0
34.8
35.5
T f (o C)
32.5
33.0
33.5
34.0
34.5
T (o C)
0.0
0.1
0.2
0.3
0.5
D (h-1 )
4.0
4.1
4.2
4.3
4.4
pH
19
22
25
28
31
200 250 300 350 400 450Time (h)
Tb (%
)
50
100
150
200
250
P f (m
mHg)
160
173
185
198
210
F f (L/
h)
32.5
33.3
34.0
34.8
35.5
T f (o C)
32.5
33.0
33.5
34.0
34.5
T (o C)
0.0
0.1
0.2
0.3
0.5
D (h-1 )
4.0
4.1
4.2
4.3
4.4
pH
19
22
25
28
31
200 250 300 350 400 450Time (h)
Tb (%
)
30
39
48
57
66
75Et
hano
l in th
e fe
rmen
tor (
g/L)
0.0
0.2
0.4
0.6
0.8
1.0
Dilut
ion fa
ctor (
h-1)
340
358
376
394
412
430
200 250 300 350 400 450Time (h)
Cond
ense
d et
hano
l (g/
L)
0.0
0.2
0.4
0.6
0.8
1.0
Dilut
ion fa
ctor (
h-1)
Dilution factor
(a)
(b)
30
39
48
57
66
75Et
hano
l in th
e fe
rmen
tor (
g/L)
0.0
0.2
0.4
0.6
0.8
1.0
Dilut
ion fa
ctor (
h-1)
340
358
376
394
412
430
200 250 300 350 400 450Time (h)
Cond
ense
d et
hano
l (g/
L)
0.0
0.2
0.4
0.6
0.8
1.0
Dilut
ion fa
ctor (
h-1)
Dilution factor
30
39
48
57
66
75Et
hano
l in th
e fe
rmen
tor (
g/L)
0.0
0.2
0.4
0.6
0.8
1.0
Dilut
ion fa
ctor (
h-1)
340
358
376
394
412
430
200 250 300 350 400 450Time (h)
Cond
ense
d et
hano
l (g/
L)
0.0
0.2
0.4
0.6
0.8
1.0
Dilut
ion fa
ctor (
h-1)
30
39
48
57
66
75Et
hano
l in th
e fe
rmen
tor (
g/L)
0.0
0.2
0.4
0.6
0.8
1.0
Dilut
ion fa
ctor (
h-1)
340
358
376
394
412
430
200 250 300 350 400 450Time (h)
Cond
ense
d et
hano
l (g/
L)
0.0
0.2
0.4
0.6
0.8
1.0
Dilut
ion fa
ctor (
h-1)
Dilution factor
(a)
(b)
SOFT SENSOR FOR CONCENTRATION
'
Penicillinprocess
RNN -+
Kalman filtertrainingweight
adjustment
State measurement
Kalman filter(NLSTC)
Error
SubstrateN
Air flow
The proposed non-linear Self-tuning controller scheme
0 20 40 60 80 100 1205
10
15
20
25
30
35
Process Kalman filter
Biom
ass
conc
entra
tion
(g/l)
Time (h)
Estimation of the biomass concentration
0 20 40 60 80 100 120-2000
0
2000
4000
6000
8000
10000
12000
14000
Process Kalman filterPe
nicillin
conc
entra
tion (
g/l)
Time (h)
Estimation of the Penicillin concentration with the multiple extended Kalman filter algorithm
Fractional Brownian motion as a model for an industrial Air-lift Reactor
fBm (Mandelbrot, 1968) BH(t+τ)-BH(t) é estatisticamente igual ao [BH(t+τr)-BH(t)]/rH
fGn: definido como derivado do fBm: fGn = BH(t+1)-BH(t)
Comparação entre o sinal de pressão e o ruído Gaussiano
fracionário (fGn)
0 500 1000 1500 2000 25003.18
3.2
3.22
3.24
3.26
3.28
3.3
3.32
Industrial Air-Lift Reactor Data Fractional Brownian Modelwith H = 0.7
0 500 1000 1500 2000 2500-4
-3
-2
-1
0
1
2
3
4
Synthesis of a fuzzy model for linking synthesis conditions with molecular
characteristics and performance properties of high density polyethylene
Cognitive Dynamic Modely(k)- prediction by linear equation – Takage Sugenoapproach:
y(k) = w0i + w1iu1(k-τu1) + w2iu1(k-τu1 -1) +...+ wp1iu1(k-τu1-p1)+w(p1+1)iu2(k-τu2) + w(p1+2)iu2(k-τu2 -1) +...+ w(p1+p2)1iu2(k-τu2-p2)+
w(p1+p2 +1)iy(k-1) + w(p1+p2 +2)iy(k-2) +...+w(p1+p2 +m)iy(k-m).
together with cognitive information
Implementations
• Du PONT Polymerization Process
• Rhodia Nylon-6,6 Process
• High Non Linear Process – large scale plantDeterministic model – difficult to assembly
Copolymer molar fraction
-200 0 200 400 600 800 1000 1200 1400 1600 18000,40
0,45
0,50
0,55
0,60
0,65
0,70
0,75
Y ap
tempo (h)
PLANTA MODELO
Teste para a fração molar do copolímero
0 200 400 600 800 1000
33000
34000
35000
36000
37000M
pw (k
g/km
ol)
tempo (h)
PLANTA MODELO
Validação para o peso molecular do copolímero
Polymer Molecular Weight
Nylon-66 Molecular weight
33000 34000 35000 36000 37000 38000
33000
34000
35000
36000
37000
38000
Mpw
(kg/
kmol
) - m
odel
o
Mpw (kg/kmol) - planta
par de dados da planta e do modelo
Phenol Hydrogentation Reactor
Módulo
ReactantsCoolant
Condição 1 2 3 4 5 6Ordem das entradas 23 17 7 23 17 7
Ordem do estado interno 1 1 1 2 2 2Regras 7 7 5 7 7 5
Fator erro indexado (J) 1.19e-3 1.2 e-3 1.22 e-3 1.17 e-4 1.19 e-4 1.21 e-3
100 200 300 400 500 600 7000,80
0,85
0,90
0,95
1,00
1,05
J = 1.2E-3
Tem
pera
tura
adi
men
siona
l dos
reag
ente
s Modelo determinístico Modelo Fuzzy
Tempo100 200 300 400 500 600 700
0,80
0,85
0,90
0,95
1,00
1,05
J = 1,21E-3
Tem
pera
tura
adi
men
siona
l dos
reag
ente
s
Tempo
Modelo determinístico Modelo Fuzzy
Ordem 7 para a entrada e1 para estado interno
Ordem 17 para a entrada e 1 para estado interno
Modelo Cognitivo Modelo Cognitivo
100 200 300 400 500 600 7000,80
0,85
0,90
0,95
1,00
1,05
J = 1,19E-3Te
mpe
ratu
ra a
dim
ensio
nal d
os re
agen
tes Modelo determinístico
Modelo Fuzzy
TempoOrdem 23 para a entrada e 1 para estado interno
Modelo Cognitivo
Density
MecanicalProperties
ThermicProperties
TensileProperties
Melt indexWeight molecularMolecular
Weightdistribution
Reologic properties
Correlation
Fuzzy model
Crystallinity
Properties Correlations
Output variablescontrol in deterministic
model
Density
MI
Performanceproperties
Thermicproperties
MechanicalProperties
RheologicProperties
TensileProperties
Weightmolecular
ProductFuzzyModel
Fuzzy model
FuzzyModel
Plant
Properties Product modelling from operationals dates throght Fuzzy Logic
Properties Product modelling from operationals dates throght Fuzzy Logic
PFR - trimer
Product
PFR
CSTR
Fuzzy Model - type A
Process
ConversionRate
productionMnMw
DensityPdMISE
StifnessImpact Strength
HardnessMelt StrengthStress CrackResistance
Tensile StrengthTmTcTg
crystallization percentmelt swell
softening Point
FuzzyModel -type C
Fuzzy Model - type B
Performance Properties
H2
CATCO-CAT
MonomerCo-monomer
Solvent
T PFRT CSTRP system
Feed Lateral
Results – Fuzzy model type A
Type A. Such model considers the linking of the property of flow stress exponent (SE) versus the variables of the synthesis process. The SE of a polymer is a measure of melt viscosity and is a direct measure of molecular weight distribution. The Stress Exponent, determined by measuring the flow (expressed as weight, in grams) through a melt index approaches (ASTM D 1238).
Optimization to achieve products with required properties
Optimization Based Polymer Resin Development
UFBA
Introduction
Polymerization process model
TemperatureConcentrations
Flow Rate
Input Conditions
0.10
0.20
0.30
0.40
0.50
0.60
0.00 0.20 0.40 0.60 0.80 1.00
Reactor Length (dim.)
SE (d
im.)
TemperatureConcentration
Conversion
Polymer Properties
Output Conditions
Optimization model
Improve Quality
Design of new products
Goal: Determine optimal operating policies in order to produce pre-specified polymer resins
PFR1
PFR2
CAT CC
Product
CSTR
CAT CC
H2
EthyleneHydrogenSolvent
EthyleneHydrogenSolvent
EthyleneHydrogenSolvent
Tubular ConfigurationStirred Configuration
Braskem Ethylene continuous polymerization in solution with Ziegler-Natta catalyst-
Industrial Plant
Mathematical Model
Product
CSTR
PFR2
CAT CC
MonomerH2Solvent
WoutWR
FZ 1
W0
B2
W1
Br+1
Wr
Br
Wr-1
BR
WR-1
.... ....CSTR1 CSTRr CSTRR
PFRJ+1
FZ r FZ R
PFR1
PFR2Product
CSTR
CAT CCH2
MonomerH2Solvent
W1 Wj WJ
WoutWR
....
Fj FJ
WP
B2
W1
Br+1
Wr
Br
Wr-1
BR
WR-1
.... ....
PFR1
CSTR1 CSTRr CSTRR
PFR J+1
PFRj PFRJ
Stirred Configuration
Tubular Configuration
Polymer Specification
• Melt Index (MI):
• Stress Exponent (SE):
• Density (DS):
( )βα wMWMI ⋅=
( )PDSE
⋅⋅+=
βγα exp1
SEMIDS ⋅+⋅+= γβα )log(
Desired polymer properties
end-point constraints of the optimization
• Specification at the end of reaction (z=zf)Embiruçu et al. (2000)
Objective Function
• Different operating policies can yield the same resin
Maximize Profit
( ) €/h SSCCCCCATCATHHMMPE WbWbWbWbWbWa ⋅+⋅+⋅+⋅+⋅−⋅=Φ
where
a: polyethylene sales price (€/kg)
b: reactant costs (€/kg)
W: mass flow rates (kg/h)
• Objective Function
Decision Variables
CSTR
PFR2Ws
CAT CC
MH2,0
TinPin
Wt
• Monomer Input Concentration (M)• Hydrogen Input Concentration (H2,0)
• Catalyst Input Concentration (CAT)
• Inlet Temperature (Tin)
• Inlet Pressure (Pin)
• Total Solution Rate (Wt)
• Side Feed (Ws)
Stirred Configuration
• Lateral Hydrogen Concentration (H2,j)• Lateral Hydrogen injection point (j)
PFR1
PFR2
CSTR
CAT CCH2,j
MH2,0
Wt
TinPin
Tubular Configuration
• Discontinuities ⇒ new stage
• Examples– Injection of mass along a tubular reactor– Reactor switch
Multi-stage Systems
z(0) z(1) z(2)
Event Event Event
f(1) f(2) ( )knf
( )1nkz − ( ) (f)n zz k =
kkkkkk ,...,nkz 1 ]z,[zz , 0),,,,,( k1-k =∈=puyxxf
0 ),,,,( =zkkkk puyxg
0)( 000 =− xx z
DAE system
0),,,,,,,,,( )()()()()()()()()( =zJ jjjjkkkkkj puyxxuyxx
Stage Transition
0.00
0.10
0.20
0.30
0.40
0.50
0.60
0.00 0.20 0.40 0.60 0.80 1.00
Reactor Length (dim.)
MI (
dim
.)Reactor Profile
PFR
PFR
CSTR
CAT CC H2
Tubular configuration
CSTR
PFR
CAT CC
Stirred configuration
PFR
0.20
0.25
0.30
0.35
0.40
0.45
0.50
0.00 0.20 0.40 0.60 0.80 1.00
Reactor Length (dim.)
MI (
dim
.)
Stage nº: 1 2 3 4 Stage nº: 1 2
Multi-stage Process
CSTR
CAT CC
Tubular configuration
CSTR
PFR2
CAT CC
Stirred configuration
Dynamic Optimization Techniques for multi-stage systems
DAE (axial coordinate) Steady-state
Analogy: axial coordinate ⇔ time
PFR1
PFR2
)(1 zf
H2
)(1 zf )(1 zf )(2 zf )(2 zf
3g3g3g
3g
)(4 zf )(4 zf )(4 zf )(4 zf
)(1 zf )(2 zf 3g )(4 zfz
)(zkfkg
: differential equation: algebraic equation: stage numberk: axial coordinatez
Results – Stirred Configuration
0.05
0.10
0.15
0.20
0.240 0.260 0.280 0.300 0.320SE (dim.)
Prof
it (d
im.)
0.0
0.2
0.4
0.6
0.8
1.0
0.240 0.260 0.280 0.300 0.320SE (dim.)
Con
cent
ratio
n (d
im.) H 2,0
CATWsM
H2,0
Ws
0.55
0.60
0.65
0.70
0.75
0.80
0.24 0.26 0.28 0.30 0.32SE (dim.)
Rev
enue
, Cos
t (di
m.)
RevenueCost
0.40
0.50
0.60
0.70
0.80
0.24 0.26 0.28 0.30 0.32SE (dim.)
Q, W
PE(d
im.)
QWPEWPE
Results – Tubular Configuration
One H2 injection point at a pre-specified length (4 stages)
0.05
0.10
0.15
0.20
0.40 0.45 0.50 0.55 0.60 0.65 0.70 0.75SE (dim.)
Prof
it (d
im.)
0.0
0.2
0.4
0.6
0.8
1.0
0.40 0.45 0.50 0.55 0.60 0.65 0.70 0.75SE (dim.)
Con
cent
ratio
n (d
im.)
H 2,0CATH 2,jM
H2,0
H2,j
Benefits of the developed toolDevelopment of a potential tool able to improve
the polymer quality or to create new resins in a simple and quick manner.
– Better customer satisfaction.
• Robust approach
– Use of Dynamic Optimization algorithms for a stationary multi-stage process.
• Versatile tool, since other polymerization processes can be used as basis.
Large Scale Plant Simulation
MODELING A FCC UNIT
RESULTS
The distillation curve was determined from thetemperature and the percentage of distillateobtained experimentally through moleculardistillation and using ASTM D1160.
0 20 40 60 80 100
0
100
200
300
400
500
600
700
Temp
eratur
e (o C)
% Distillate accumulated (% w)
Through CENPES/PETROBRAS Through Molecular Distillation
Molecular Distillation of
the Alfa petroleum
obteined 10 %of distillate acumullated
SEPARATION SECTION OF THE FCCU
Product Industrial data (ton/day) Simullation Result (ton/day) Error (%)
Fuel Gas 360.0 360.4 0.11
LPG 1167.0 1191.4 2.09
Gasoline 3534.0 3436.2 2.77
LCO 667.0 677.0 1.50
Slurry 1107.0 1067.5 3.57
Products recovery: industrial data and simulation results.
Green Ethyl Acrylate
O
O-
CCH3
CH
NH3+
O
O-
CCH3
CH
OH
O
O-
CCH3
CH2L-Alanina Lactato
Propianato
O
O-
CCH2
CH
Ácido Acrílico
Glicose Lactose Sacarose VáriosC5 e C6
S U B S T R A T O S
1 2
4
3
5 6
Fermentação1) Fermentação de ácido Láctico (ex. Lactobacilli, Bacilli Streptokokki).2) Fermentação de ácido Propiónico.3) Redução Direta (ex. Clostridium propionicum).4) Desidratação5) Conversão Química6)Caminho Oxidativo (ex. Pseudomonas aeroginosa)
FEED
REC1
TOPO
COOL
WATER
RAF
EXT
REC3
ACRYLATE
REC2
ETHANOL
STRIPPER
COOLER
EXTRACT
DISTIL2
DISTIL1
REACTOR
ACID
WASTE
Conceptual Plant design for Green Ethyl Acrylate
Reactor Mathematical ModelEquações adimensionalizadasBalanço de Massa para o Ácido Acrílico
Balanço de Energia no Tubo
Balanço de Energia do Fluido Térmico
Queda de Pressão
Aad
rBuGu
uGB
zG ... 22
2
1 +
∂∂
+∂∂
=∂∂
All
ad
l rBu
uu
Bz
... 42
2
3 +
∂∂
+∂∂
=∂∂ θθθ
( )FNTad
BdzdQ θθ −= .5
7BdzdP
ad
ad =
Solução por Colocação Ortogonal
Reactor simulationConversion for several temperatures Tubular reactor 5,0 meters long
0,0 0,2 0,4 0,6 0,8 1,0
0,0
0,1
0,2
0,3
0,4
0,5
0,6
0,7
Conv
ersão
Coordenada Axial
Conversão @ 75 C Conversão @ 80 C Conversão @ 85 C
Conceptual Plant design for Green Ethyl Acrylate
FEED
REC1
TOPO
COOL
WATER
RAF
EXT
REC3
ACRYLATE
REC2
ETHANOL
STRIPPER
COOLER
EXTRACT
DISTIL2
DISTIL1
REACTOR
ACID
WASTE
Vazão(kmol/h) FEED
REC1 TOP COOL
WATE
RAF EXT ACRYL REC2 WAST REC3
Ácido Acríl 20,82 20,82 0,00 0,00 0,00 0,00 0,00 0,00 0,0000 0,0000 0,0000
Etanol 20,82 0,00 20,82 20,82 0,00 0,00 20,82 0,00 0,0000 0,3790 20,4409
Água 29,18 0,00 29,18 29,18 20,00 4,59 44,58 0,00 4,5999 36,3804
8,1995
Acril de Etila 29,18 0,18 29,00 29,00 0,00 22,64 6,36 19,74 2,9003 0,00 6,3595
Total (kmol/h) 100,0 21,00 79,00 79,00 20,00 27,23 71,70 19,74 7,5000 36,76 35,00
Total (Kg/h) 5.907 1.518 4.388 4.388 360 2349 2.399 1976 373,53 672,91 1726,11
Temp. (ºC) 78,0 140,5 79,0 25,00 25,0 29,9 29,3 99,4 82,42 97,10 77,71
Pressão (atm) 1 0 1 0 1 0 1 0 1 0 1 0 1 0 1 0 1 0 1 0 1 0
Green Acrylic Acid – Unicamp/CTC/Braskem
Cana –fonte de açúcar
Seleção de microrganismos Otimização do meio de cultura
Seleção das rotas metabólicas
Fermentação
Ácido LácticoSeparação/purificação Cinética
Processodesidratação reduçãoÁcido Acrílico
Ácido Propiônico
Otimização dos processos
Cinéticas Modelagem Controle dos processos
Fontes petroquímicas
Acetaldeído (CH3CHO)
Lactonitrila (CH3CHOHCN
Mistura racêmica DL - ácido láctico
Adição de HCN e catalisador
Hidrólise por H2SO4
Fontes renováveis
Carboidratos fermentescíveis
Caldo fermentado
Ácido láctico L(+) ou D(-) opticamente puro
Pré-tratamento (hidrólise)
Fermentação microbiana
Recuperação e purificação
Síntese química Fermentação microbiana
SSF
PROCESSOS DE PRODUÇÃO DE ÁCIDO LÁCTICO[1]
PRODUTOS OBTIDOS A PARTIR DO ÁCIDO LÁCTICO [2]
Ácido Láctico
Açúcar
Ferm
enta
ção
descarboxilação
desidratação
redução
condensação
+ CO + H2OCO2+ H2
O
H
O
OH
O
OH + H2O
Acetaldeído
Ácido Acrílico
Ácido Propiônico
H2+ 1/2 O2
O
O
+ CO2 + 2H2O
2,3 pentanodiona
Proposed Process Intensification and Green Technology for Ethyl Acetate Production
45 ton/month
The proposed system design: coupled reactor/column configuration.
Reactor is the column reboiler
Configuration of the overall ethyl acetate process by Reactive Distillation
Sustainable Global Process (SGP)
Process environmentally non aggressive, including total water recovery and the lowest energy consumption.
All reactants are renewable
Steady State and Dynamic simulation – Reactive Distillation is more stable then configuration coupling reactor and column
Sugar and Ethanol Integrated Process
Separation units
The hydrous ethanol, with 96°GL, is an azeotropicmixture of ethanol and water, and therefore can not be more concentrated by pure distillation. The additional water removal is accomplishedin the so called dehydration process:•Azeotropic distillation
•Extraction distillation with monoethyleneglycol
•Molecular sieves
Ethanol dehydration technologies Technology Dehydrating Steam consumption Agent (kg steam/L ethanol) Azeotropic Cyclohexane 1.7 Extraction Monoethyleneglycol 0.7 Molecular sieves zeolite beads 0.6
Simulation of bioethanol production processes from sugarcane juice and
bagasse, using an Organosolv process with dilute acid hydrolysis
L O PCALaboratório de Otimização, Projeto e Controle Avançado
Goals
To perform the simulation of conventional bioethanol production process from sugarcane juice,
considering the introduction of technologies that improve its energetic efficiency
Integrate the production of bioethanol from sugarcane bagasse, using an Organosolv process
with dilute acid hydrolysis
Block flow diagram – conventional bioethanol production process
Block flow diagram –bioethanol production process from bagasse
Process simulation
• Mass balance based on literature and industry data was made on a spreadsheet
• Simulation using Hysys• Thermodynamic models:
– Until distillation: NRTL and SRK– Hydrolysis: NRTL and SRK– Extractive distillation: UNIQUAC and SRK– Azeotropic distillation: NRTL and SRK
Process simulation
• Production of 1000 m³/day of anhydrous bioethanol from sugarcane juice 500 tons of sugarcane per hour
• Integrated process: use of sugarcane bagasse as raw material. along with sugarcane juice
• Fermentation of the hydrolyzed liquor in a mixture with juice
Simulation components – hypothetical
• Since not all components present in bioethanol production are available at Hysys database, some hypothetic components were created to represent:– Conventional process components: sugarcane
bagasse (cellulose, hemicellulose and lignin), dirt, impurities (salts, organic acids), lime, phosphoric acid, yeast
– Hydrolysis components: pentose and HMF
Unit operations
• Splitters were used to represent the following equipments/operations: Sugarcane dry cleaning system Mills Screens and hydrocyclones Settler Filters and separators
• Multiple effect evaporators were represented by a system comprised by separator, valve and heat exchanger
Unit Operations
• Conversion reactors Fermentation: conversion data based on those provided
by the industry Hydrolysis: data based on experiments available in the
literature (pre-hydrolysis of sugarcane bagasse and hydrolysis of chemical grade cellulose)
• Centrifuges were simulated as solids separators
Simulation of anhydrous bioethanol production process from sugarcane juice
Sugarcane bagasse hydrolysis
Ethanol dehydration
Ethanol dehydration processes
• Two different processes were analyzed:– Extractive distillation: both conventional and
alternative configuration– Azeotropic distillation
• Solvents evaluated:– Extractive distillation: monoethyleneglycol (MEG) and
glycerin– Azeotropic distillation: cyclohexane and n-heptane
Extractive distillation with MEG –conventional configuration
Extractive distillation – alternative configuration
Azeotropic distillation
Comparison between extractive and azeotropic distillation
Parameter
Extractive Distillation Azeotropic Distillation
Conventional Alternative Ciclo-hexane
n-HeptaneMEG Glyc. MEG Glyc.
Vapor consumption (kg/L anydr ethanol) 0.43 0.47 0.41 0.56 8.0 6.1
Saturated steam pressure (bar) 6 10 / 65 6 65 2.5 2.5
Ethanol losses (%) 10-5 10-5 9x10-5 6x10-5 0.017 0.017Solvent losses (%) 0.01 0.01 0.49 0.02 0.001 0.008
Solvent in anhydrous ethanol (wt%) No contamination with solvent 0.017 0.04
Double effect distillation
Steam consumption on column reboilers – conventional and double effect
distillation
Parameter Distillation processConventional Double-effect
2.5bar steam consumption – column A 1.53 0.002.5bar steam consumption – column B 0.27 0.386bar steam consumption – extractive
column 0.35 0.35
6bar steam consumption – recovery column 0.07 0.07
Total steam consumption 2.21 0.80Steam consumption - [kg/L anhydrous ethanol]
Anhydrous bioethanol production –integrated process
Energy consumption on conventional and integrated production process
Parameter
Energy consumptionConv Int Conv Int
(kJ/kg anhydrous ethanol)
(kJ/kg sugarcane)
Heating operations 15529 22771 1052 1803Cooling operations 9940 16951 672 1342Increase on heating (%) 46 71Increase on cooling (%) 71 100Conv: conventional bioethanol production process; Int: integrated process with hydrolysis of 70 % generated sugarcane bagasse
Process integration of the biorefinery
– Thermal integration considers that 50 % of sugarcane straw is used as a fuel in the boilers
– When the amount of energy produced is equal to that required by the biorefinery, the real amount of bagasse available for hydrolysis is determined
• It was found that 60 % of the bagasse generated in the mills may be used as raw material for hydrolysis and still make the biorefinery self sufficient on its energy production, in a conventional distillation system
Products and inputs of the biorefinery after process integration
Parameter Units Conventionalprocess
Integratedprocess
Sugarcane input t/h 493 493Bagasse input t/h 0 70.7Anhydrous bioethanolproduction
m3/day 1004 1178L/t cane 84.8 99.6
Increase in production % - 17.5Pentose liquor (9 wt%) t/h - 10.3Integrated process: sugarcane juice and 60 % of sugarcane bagasse as raw materials
Fermentation cooling
• Dias, M.O.S., Maciel Filho, R., Rossell, C.E.V., Efficient cooling of fermentation vats in ethanol production, Sugar Journal, V. 70, p. 11-17, 2007
• Fermentation is usually done at 34°C• Limits ethanol content of the wine• Increases energy consumption (centrifuges and
distillation columns)• Increases stillage volume• Promotes infection and yeast inhibition
Fermentation cooling
• Fermentation cooling is done using water from rivers– increasing environmental restrictions – or coolingtowers of low efficiency
• Carrying fermentation at lower temperatures (28°C)increases ethanol content of the wine, but demandsthe use of more efficient cooling equipment
• Options considered:• More efficient cooling tower• Water accumulator• Steam jet ejector• Absorption machine
Ethanol content of the wine – impact on ethanol losses in vinasse
Considering the production of 1000 m³/day of anhydrous
bioethanol and vinasse with 0.02 wt% ethanol
Ethanol content of the wine – impact on ethanol losses in vinasse
For lower ethanol content of the wine, more wine is
necessary to produce the same amount of anhydrous bioethanol (1000 m³/day),
thus increasing energy consumption on
fermentation and distillation as well as capital costs
Fermentation cooling – water accumulator
Cooling Tower
PHEVat
Water Accumulator
Cooling water
Cooling water in
Cooling water in
Cooling water out
Fermentation cooling – steam jet ejector
Vat PHE
Evaporator
Ejector1,4 bar Steam
Make-up water
Purge
Cooling water out
Cooling water in
t=28ºC
t2
T1
T2q1
q2
q3
q4
m1
q5
Water
Control and Real Time Optimization
Non Linear Intelligent Controller ^
Processo Controlador
Modelo Neural Ajuste Controlador
Dados Passados Aprendizagem Redes Modelo Neural
Filtro Referência
J - 1 Y r
Y
Rotina de Otimização
)()()( kÛkUkê r −=
( ) ( ) ( )kYkYke wˆˆ −=
On-line Lerning
Dados Atuaisdo Processo
RNA 2Base - Padrão
ERRO 2
MenorErro
RNA 1Padrão
ERRO 1
RNA 3Nova
ERRO 3
Se erro 1 for menor
Peso
s Pa
drõe
s
Aju
ste
de P
esos
Aju
ste
de P
esos
Dados Atuaisdo Processo
RNA 2Base - Padrão
ERRO 2
MenorErro
RNA 1Padrão
ERRO 1
RNA 1Padrão
ERRO 1
RNA 3Nova
ERRO 3
RNA 3Nova
ERRO 3
Se erro 1 for menor
Peso
s Pa
drõe
s
Aju
ste
de P
esos
Aju
ste
de P
esos
Caso de Estudo 1
Processo Extrativo de Fermentação Alcoólica Contínua
Perturbação Estocástica
0 20 40 60 80 100 120 140 160 180 200
38
39
40
41
42
43
44
45
46
47C
once
ntra
ção
de P
rodu
to (k
g/m
3 )
Tempo (h)
Referência Malha Fechada Resultado com Perturbação Estocástica
CRAQUEAMENTO CATALÍTICO
Configuração da Rede
Rede composta por três camadas:
Camada de entrada: 8 variáveis para o
instante atual;
Camada de saída: 4 variáveis para o instante
futuro;
Camada oculta: 20 neurônios;
Erro inferior a 0,005;
Configuração final - 8 x 20 x 4;
RNA - FCC
Comparação entre o MPC - Neural e o PID
Controle preditivo do tipo baseado em modelo - MPC
Real Time Process Integration
Large Scale Cumene Oxidation Reactor
4 Air-Lifts Reactors in Sequence
TCR1
AIR1
FCR1
AIR1
FCR1
TCR1
AIR1
FCR1
AIR1
AIR1
TCR1
AIR1
FCR1
AIR1
TCR1
AIR1
FCR1
AIR1
Results from Industrial On Line Control
%HPOC R104 D
25,00
26,00
27,00
28,00
29,00
30,00
31,00
32,00
33,00
34,00
35,00
15/0
7/99
12:
00
16/0
7/99
00:
00
16/0
7/99
12:
00
17/0
7/99
00:
00
17/0
7/99
12:
00
18/0
7/99
00:
00
18/0
7/99
12:
00
19/0
7/99
00:
00
19/0
7/99
12:
00
20/0
7/99
00:
00
20/0
7/99
12:
00
21/0
7/99
00:
00
21/0
7/99
12:
00
22/0
7/99
00:
00
22/0
7/99
12:
00
23/0
7/99
00:
00
23/0
7/99
12:
00
24/0
7/99
00:
00
24/0
7/99
12:
00
25/0
7/99
00:
00
25/0
7/99
12:
00
26/0
7/99
00:
00
26/0
7/99
12:
00
27/0
7/99
00:
00
27/0
7/99
12:
00
28/0
7/99
00:
00
data
HPO
C (%
)
AI-114AC-114_SV
tempo
AC140_PVAC140_SP
∆SP
= 3
,5%
%HPOC R104 A
10,00
10,50
11,00
11,50
12,00
12,50
13,00
13,50
14,00
14,50
15,00
15/0
7/99
12:
00
16/0
7/99
00:
00
16/0
7/99
12:
00
17/0
7/99
00:
00
17/0
7/99
12:
00
18/0
7/99
00:
00
18/0
7/99
12:
00
19/0
7/99
00:
00
19/0
7/99
12:
00
20/0
7/99
00:
00
20/0
7/99
12:
00
21/0
7/99
00:
00
21/0
7/99
12:
00
22/0
7/99
00:
00
22/0
7/99
12:
00
23/0
7/99
00:
00
23/0
7/99
12:
00
24/0
7/99
00:
00
24/0
7/99
12:
00
25/0
7/99
00:
00
25/0
7/99
12:
00
26/0
7/99
00:
00
26/0
7/99
12:
00
27/0
7/99
00:
00
27/0
7/99
12:
00
28/0
7/99
00:
00
data
HPO
C (%
)
AI-111AC-111_SV
%HPOC R104 B
14,00
15,00
16,00
17,00
18,00
19,00
20,00
21,00
15/0
7/99
12:
00
16/0
7/99
00:
00
16/0
7/99
12:
00
17/0
7/99
00:
00
17/0
7/99
12:
00
18/0
7/99
00:
00
18/0
7/99
12:
00
19/0
7/99
00:
00
19/0
7/99
12:
00
20/0
7/99
00:
00
20/0
7/99
12:
00
21/0
7/99
00:
00
21/0
7/99
12:
00
22/0
7/99
00:
00
22/0
7/99
12:
00
23/0
7/99
00:
00
23/0
7/99
12:
00
24/0
7/99
00:
00
24/0
7/99
12:
00
25/0
7/99
00:
00
25/0
7/99
12:
00
26/0
7/99
00:
00
26/0
7/99
12:
00
27/0
7/99
00:
00
27/0
7/99
12:
00
28/0
7/99
00:
00
data
HPO
C (%
) AI-112AC-112_SV
%HPOC R104 C
20,00
21,00
22,00
23,00
24,00
25,00
26,00
27,00
28,00
29,00
30,00
15/0
7/99
12:
00
16/0
7/99
00:
00
16/0
7/99
12:
00
17/0
7/99
00:
00
17/0
7/99
12:
00
18/0
7/99
00:
00
18/0
7/99
12:
00
19/0
7/99
00:
00
19/0
7/99
12:
00
20/0
7/99
00:
00
20/0
7/99
12:
00
21/0
7/99
00:
00
21/0
7/99
12:
00
22/0
7/99
00:
00
22/0
7/99
12:
00
23/0
7/99
00:
00
23/0
7/99
12:
00
24/0
7/99
00:
00
24/0
7/99
12:
00
25/0
7/99
00:
00
25/0
7/99
12:
00
26/0
7/99
00:
00
26/0
7/99
12:
00
27/0
7/99
00:
00
27/0
7/99
12:
00
28/0
7/99
00:
00
data
HPO
C (%
)
AI-113AC-113_SV
%HPOC R104 D
25,00
26,00
27,00
28,00
29,00
30,00
31,00
32,00
33,00
34,00
35,00
15/0
7/99
12:
00
16/0
7/99
00:
00
16/0
7/99
12:
00
17/0
7/99
00:
00
17/0
7/99
12:
00
18/0
7/99
00:
00
18/0
7/99
12:
00
19/0
7/99
00:
00
19/0
7/99
12:
00
20/0
7/99
00:
00
20/0
7/99
12:
00
21/0
7/99
00:
00
21/0
7/99
12:
00
22/0
7/99
00:
00
22/0
7/99
12:
00
23/0
7/99
00:
00
23/0
7/99
12:
00
24/0
7/99
00:
00
24/0
7/99
12:
00
25/0
7/99
00:
00
25/0
7/99
12:
00
26/0
7/99
00:
00
26/0
7/99
12:
00
27/0
7/99
00:
00
27/0
7/99
12:
00
28/0
7/99
00:
00
data
HPO
C (%
)
AI-114AC-114_SV
Reator AReator B
Reator C Reator D
tempo
tempo tempo
tempo
Results from Industrial Reactor On Line Control
DMC + PCA
-4
-3
-2
-1
0
1
2
3
4
0 20 40 60 80 100n
Y
AC110AC120AC130AC140
-4
-3
-2
-1
0
1
2
3
4
0 20 40 60 80 100n
Y
AC110AC120AC130AC140
On-line Optimization
• Perfis inadequados de temperatura e concentração levam ao aumento da formação de subprodutos
• O otimizador em linha funciona em malha fechada com controladoresmultivariáveis, mantendo as concentrações e as temperaturas nos seus valores ótimos.
Função objetivo:
f(x) = (% impurezas) mínimoΣ
Restrições:
g(x) = [T1, T2, T3, T4, produção]
As concentrações das impurezas são calculadas pelo modelo
SQP Temperaturas máximas dos reatores
OTIMIZAÇÃO EM LINHA
Etapas envolvidas na otimização em linha
• Aquisição de dados do processo em tempo real;• Verificação da qualidade dos dados (filtro, intervalos de validade); • Teste de estado estacionário;• Execução do modelo e comparação entre a simulação e os dados reais do processo;• Ajuste de parâmetros do modelo para minimização dos erros;• Verificação da qualidade do ajuste;
SIMULAÇÃO EM LINHA
• Execução da subrotina de otimização;• Verificação da qualidade dos ótimos;• Decisões sobre os valores que serão enviados ao SDCD (trajetória de otimização);• Entrada no modo de espera (tempo entre otimizações);• Saída do modo de espera e re-início do algorítmo;
SP - valor do setpoint atualPV - valor atual da variável de processoOPT - valor ótimo calculadoDmax - máxima variação permitida do setpoint com relação
à variável de processoNSP - novo setpoint
OTIMIZAÇÃO EM LINHA
Trajetória de Otimização
Situações possíveis para NSP:
NSP = SP (nenhuma modificação é feita)
NSP = SP + / - Dmax (mudança incremental na direção do ótimo)
NSP = OPT (a solução ótima é admitida)
On-line Optimization with a Multivariable Controller
Unicamp software
22.00
22.20
22.40
22.60
22.80
23.00
23.20
23.40
16/0700:00
16/0706:00
16/0712:00
16/0718:00
17/0700:00
17/0706:00
17/0712:00
17/0718:00
18/0700:00
18/0706:00
18/0712:00
18/0718:00
19/0700:00
date
CH
P (%
)
CHP_PV reactor 3CHP_SP reactor 3
..... AC130_PVAC130_SP
On-line Optimization
600
650
700
750
800
850
900
07/0700:00
08/0700:00
09/0700:00
10/0700:00
11/0700:00
12/0700:00
13/0700:00
14/0700:00
15/0700:00
16/0700:00
17/0700:00
18/0700:00
19/0700:00
20/0700:00
21/0700:00
22/0700:00
date
DM
PC (k
g/h)
DMPC (kg/h)
(σ1 , x1)
(σ2 , x2)
without optimization
with online optimization
Impurities Reduction
Impureza 1 (kg/h)
σ1 / σ2 = 1.846
x1 / x2 = 1.012
Simulation Tools are able to help in process design, operation, optimization and control