Reac t or Des i g n
Reactor ct or Design si gn • Predic Predictio tion n of reacto reactorr perf performa ormance nce,, prod product uct yields yields etc. etc. – See earlier lecture
• Detail Detailed ed discus discussio sion n of of reactio reaction n kine kinetics tics,, cata catalysi lysis, s, deactivation, mass transfer, etc. – See reactors classes and textbooks textbooks
• Focu Focus s of this this lec lectu ture re is is on how how rea reall reac reacto tors rs are are designed and sized in industry • Specia Speciall case case of of biol biolog ogical ical reacto reactors rs is treate treated d in in next next lecture
Reactor ct or Design si gn • Predic Predictio tion n of reacto reactorr perf performa ormance nce,, prod product uct yields yields etc. etc. – See earlier lecture
• Detail Detailed ed discus discussio sion n of of reactio reaction n kine kinetics tics,, cata catalysi lysis, s, deactivation, mass transfer, etc. – See reactors classes and textbooks textbooks
• Focu Focus s of this this lec lectu ture re is is on how how rea reall reac reacto tors rs are are designed and sized in industry • Specia Speciall case case of of biol biolog ogical ical reacto reactors rs is treate treated d in in next next lecture
Reactor ct or Sizin izing g & Costi os ting ng • Estim stimat ate e re requir quire ed vol volu ume – From residence time for non-catalytic non-catalytic reactors – From catalyst space velocity for packed bed catalytic catalytic reactors • Space Space velocity velocity = lbs/h lbs/h per lb cata catalys lystt • Hence use catalyst catalyst averag average e bed density density to estimate estimate cataly catalyst st bed volume volume
– From hydraulics & residence time for for fluidized and slurry reactors – Make allowance for for head space, internals, etc.
• Deci Decide de pres pressu sure re vess vessel el size size and and sha shape pe – See pressure vessel design lecture
• Cost Cost reac reacto torr she shellll as as a pres pressu sure re vess vessel el • Add extra extra costs costs for mixers mixers,, inte interna rnals, ls, contro controls, ls, etc. etc.
Complications of Real Reactor Design
How do we handle multiple phases? How do we get good mixing & segregation?
How do we introduce catalyst?
What gives lowest cost?
How do we add or remove heat? How tight does RTD have to be?
Real Reactor Design • Very often, the design of real reactors is a lot more complicated than just estimating the reactor volume • Much of the cost comes from reactor internals – Mixers, agitators, baffles – Heat transfer (jackets, coils or external loops) – Catalyst handling
• The mixing and heat transfer performance of real reactors can be very difficult to model and understand, and can have significant effects on process yields and product purity
Reactor Design • Basics of Reactor Design • Mixing in Industrial Reactors • Heat Transfer in Industrial Reactors • Vapor-Liquid Reactors • Reactors for Liquid Catalysis • Reactors for Solid Catalysis
“ Ideal” Reactors
WMR or CSTR
Plug Flow Reactor
•
Perfect mixing
•
No axial mixing
•
Product and entire vessel contents are at uniform temperature, concentration
•
Sharp residence time distribution
•
Material flowing through the reactor experiences a profile of concentrations and temperatures
•
Material sees a distribution of residence times
Idealized reactor performance is seldom attained in practice, but is useful as a first approximation
Reactor Performance • Plug flow reactor: dV G
G = molar flow rate V = volume X = conversion R = reaction rate per unit volume
Balance across element of reactor: -G dX = R dV Integrated form depends on rate expression R(X)
• Well mixed reactor: G V
Balance across reactor: G (Xin – Xout) = R V R is evaluated at outlet conditions
Reaction Kinetics Complications • Reactions are seldom simple first or second order • Most catalytic reactions can be fitted with LangmuirHinshelwood expressions – Inhibition terms are often significant
• Mass transfer, mixing & equilibrium often limit the overall rate • Catalyst deactivation is often significant • Simple first order model is usually adequate for predicting conversion, but not for predicting byproduct yields or understanding catalyst behavior
Mass Transfer • Mass transfer processes often reduce the overall rate of reaction to a slower rate than intrinsic kinetics • Mass transfer limitations can occur: – Between phases (V/L, L/L, L/S, V/S, etc.) – Inside catalyst pores
• Inter-phase transport is strongly influenced by interfacial area, i.e., particle, droplet or bubble size (hence agitation rate) • See reaction engineering textbooks for numerous examples with neat analytical solutions
First Order Approximation • Very often we can write: R = k eff CA
• CA is the concentration of one of the reagents (the limiting reagent) • k eff is effective first order rate constant – Includes mass transfer resistances – Includes concentrations of reagents that are present in excess and so roughly constant
• For an equilibrium reaction, expression is: R = k eff (CA – CA*)
CA* = equilibrium concentration
Reactor Heat Balance Reactor design must account for enthalpy difference between feed and products, which can come from: • Heat of reaction:
dH = G.(Xout – Xin).ΔHrxn
• Heat of reaction must be calculated at reaction temperature and pressure
• Sensible heat changes:
dH = m.C p.dT
• Latent heat due to phase changes:
dH = δm.ΔHL
• In industrial practice, all of these are usually estimated using process simulation software: dHreactor = H products - Hfeeds
Reactor Design • Basics of Reactor Design • Mixing in Industrial Reactors • Heat Transfer in Industrial Reactors • Vapor-Liquid Reactors • Reactors for Liquid Catalysis • Reactors for Solid Catalysis
Mixing in Industrial Reactors Tubular Reactors
• Tubular reactors are almost always designed to be in turbulent flow • A static mixer is usually placed immediately downstream of any feed point to ensure reactor contents are mixed quickly • Static mixer usually consists of baffles to induce turbulence
Source: Komax Inc. www.Komax.com
Mixing in Industrial Reactors Stirred Reactors • Agitator consists of impeller mounted on shaft driven by motor • Motor is usually mounted above the reactor • Reactor usually contains baffles or other internals to induce turbulence and prevent contents from swirling
© 2007 Chemineer Inc. Used with permission. www.Chemineer.com
Impeller Types
Straight Blade
Helical Ribbon
Pitched Blade
Anchor
Hydrofoil
Rushton Turbine
© 2007 Chemineer Inc. Used with permission. www.Chemineer.com
Propeller (Turbine)
Screw
Baffles •
Flow pattern
If the tank has no baffles then the liquid will swirl and develop a vortex: Liquid level
Side view
•
Usually four baffles are placed around the perimeter to break up swirl – Typically, baffles are 1/10 of diameter and located 1/20 of diameter from wall
Top view Flow pattern
Baffle
Impeller Reynolds Number • Can be used to determine extent of mixing and correlate power consumption and heat transfer to shell (jacket) • Defined as 2
Re =
Da N ρ
µ
Da = agitator blade diameter, m N = agitator speed, revs/s ρ = density, kg/m3 μ = viscosity Ns/m2
• Different definitions are used for agitators without blades
Power Consumption •
Power consumption P (in W or Nm/s) can be made into dimensionless power number, N p, which can be correlated against impeller Reynolds number
N p •
P =
ρ N 3 Da
5
For Re > 103, power number is roughly constant and mainly a function of impeller type N p
10
•
102
103
Re
See Perry’s Handbook or vendors for correlations
Non-Ideal Flow and Mixing • In some cases, simple correlations may not be adequate: – If dead zones cannot be tolerated for reasons of product purity, safety, etc. – If reactor internals are complex – If reaction selectivity is very sensitive to mixing
• In these cases, it is usually necessary to carry out a more sophisticated analysis of mixing – Use computational fluid dynamics to model the reactor – Use physical modeling (“cold flow”) experiments – Use tomography methods to look at performance of real reactor
Computational Fluid Dynamics •
Calculate mass, energy and momentum balances discretely across a 2- or 3-dimensional grid of points as a function of time
•
Can include effects of heat and mass transfer, bubbles, suspended solids
•
Boundary conditions on grid are set up to reflect reactor geometry
•
Results are usually plotted as color coded pictures of velocity, mass transfer coefficient, void fraction, shear, etc., that let the designer see where the weak points of the design may be and propose changes to the design geometry
•
Commercial software such as Fluent®, CFX or FloWizard is used (see www.Ansys.com) Source: Ansys Inc. www.Ansys.com
Reactor Tomography •
Various methods can be used for non-invasive examination of reactor in-situ – Cat Scanning, Ultrasound, Gamma Scanning – Usually carried out by specialist contractors, & not cheap
Cat Scanning of FCC regenerator to validate MTO reactor catalyst distribution
Source: UOP
Gamma scanning to validate axial catalyst density profile in FCC regenerator
Reactor Design • Basics of Reactor Design • Mixing in Industrial Reactors • Heat Transfer in Industrial Reactors • Vapor-Liquid Reactors • Reactors for Liquid Catalysis • Reactors for Solid Catalysis
Non-Isothermal Liquid Phase Reactors
• Low heat duties can be achieved with a jacketed vessel: – Q ≈ U A ΔT
• Intermediate duties require an internal coil • • • •
But note: coil impacts mixing, fouling and cleaning Q = U A Lmtd U can be estimated using correlations for shell side of S&T HX Coil volume must be added to volume calculated from residence time
• High duties require an external heat exchange circuit
Estimating Heat Transfer Coefficients in Stirred Tank Reactors •
Reactor side heat transfer coefficient depends strongly on rate of agitation, reactor internals & coil design – Very case specific – Detailed understanding requires CFD or physical modeling
•
First approximation for jacket for design purposes: Nu = α Reβ Pr 0.33
•
Ch 19 (section 19.18) has values for different impellers: – α is in range 0.36 to 1.4, – β is in range 0.5 to 0.75, typically 0.67 – Re is the impeller Reynolds number – Nu = hd/k , where d is reactor internal diameter
Example •
A well-mixed reactor for manufacturing a specialty chemical has diameter 2m and liquid depth 3m. The agitator is a paddle with diameter 0.2m and speed is 60 rpm. The reactor operates at 75 °C, and a cooling rate of 200 kW is required. How would you cool the reactor?
•
Start by assuming typical organic chemical properties – Pr ~ 0.9, k ~ 0.14 W/mK, ρ ~ 700 kg/m3, μ ~ 0.6 × 10-3 Ns/m2
•
60 rpm = 1 rps, so Re = (0.22)×700×1/0.6 × 10-3 = 46700
•
From Ch 19, Nu = 0.36 Re0.67 Pr 0.33 = 467, and h = k Nu/d = 0.14 × 467/2 = 33 W/m2K
•
Heat transfer coefficient on jacket side using cooling water ~ 800 W/m2K, so U ~ (1/800 + 1/33)-1 = 31 W/m2K
•
Jacket area is π.d.L = 3.14 × 2 × 3 = 18.85m2, So cooling duty = 31 × 18.9 × dT ~594dT
•
If cooling water is available at 45 °C, then maximum delta T would be 30 °C and maximum cooling rate would be 594 × 45 = 26.7 kW
•
Jacket is not adequate and we should increase stirrer speed or agitator length or consider a coil or external loop
Non-Isothermal Vapor Phase Reactors • Heat transfer coefficients are usually too low to use jackets or internal coils • External heating or cooling loops are most common • For very endothermic processes, reaction is carried out in a fired heater tube – Reactor design is same as fired heater design – Allow extra residence time in radiant zone if necessary – See later
Reactor Design • Basics of Reactor Design • Mixing in Industrial Reactors • Heat Transfer in Industrial Reactors • Vapor-Liquid Reactors • Reactors for Liquid Catalysis • Reactors for Solid Catalysis
How would you get a vapor to react with a liquid?
Vapor-Liquid Reactors Goal
Types of V-L Reactor
Maintain low concentration of gas component in liquid
- Sparged stirred tank reactor
Contact gas and liquid over catalyst
- Trickle bed reactor
React a component out of the gas phase to high conversion
- Multi-stage V/L contactor (reactive absorption column)
- Sparged tubular reactor
Examples
- Liquid phase oxidations using air - Fermenters
-
-
Slurry phase reactor
Venturi scrubber
- Catalytic hydrogenation
- Chemisorption - Acid gas scrubbing
Sparged Reactors
•
Sparger is a pipe with holes for bubbles to flow out
•
For smaller bubbles, a porous pipe diffuser can be used instead
•
Balance between bubble break-up and coalescence is quickly established
•
If small bubble size must be maintained then additional shear is needed and an agitator is used as well
•
Designer must allow some disengaging space at top of reactor, or entrainment will be excessive
Sparger as Agitator •
If gas flow rate is large then gas flow can be used as primary means of agitation
•
Perry’s Handbook suggests the following air rates (ft3/ft2.min) for agitating an open tank full of water at 1 atm:
Degree of agitation
Liquid depth 9ft
Liquid depth 3ft
Moderate
0.65
1.3
Complete
1.3
2.6
Violent
3.1
6.2
Lift Reactors and Loop Reactors • If sparger is used to provide agitation then a baffle is often added to give better liquid circulation and ensure mixing of feeds • These reactors can be used for very large flowrates, where the liquid flow is driven by the vapor flow • Equipment design is governed by two phase flow hydraulics (see earlier lecture)
Baffle
Example: UOP/Paques Thiopaq Reactor
•
Biological desulfurization of gases with oxidative regeneration of bugs using air
•
Reactor at AMOC in Al Iskandriyah has six 2m diameter downcomers inside shell
Reaction in Vapor-Liquid Contacting Columns • Trayed or packed columns can be used to contact vapor and liquid for reaction – See separation columns lecture for design details
• Packing may be catalytically active, or could be conventional inert packing • Design is similar to design of absorption columns, but must allow for enhancement of absorption due to reaction
Vapor-Liquid Reaction Kinetics Liquid B CA,∞
•
Vapor
δ
A
CA,i
If liquid component B is present in excess then we can assume reaction is psuedo-first order in gas component A Rate of reaction = k 2 C A C B ≈ k 1C A
•
Start by assuming reaction in bulk is >> reaction in mass transfer film Mass transfer flux through film
= =
(
k L a C A,i
−
C A,∞
)
Rate of reaction in bulk = k 1 C A,
∞
Vapor-Liquid Reaction Kinetics Solving :
C A,∞
=
C A,i a k L
(k 1 + a k ) L
so rate of reaction (or flux) =
k 1 C A,i a k L
(k 1 + a k ) L
=
a k L C A,i
k 1
(k 1 + a k ) L
We can define two regimes: •
k 1 << ak L, rate ≈ k 1C A,i – Known as slow kinetic regime – Reaction rate occurs with concentration that would be predicted by phase equilibrium – Design is insensitive to increase in area a
•
k 1 >> ak L, rate ≈ a k LC A,i – Known as slow mass transfer regime – Reaction rate occurs at the rate that would be set by mass transfer with zero concentration in the bulk liquid – Design is sensitive to increase in area a
Vapor-Liquid Reaction Kinetics • For either of the slow regimes to occur we need reaction to mainly occur in the bulk liquid Reaction in film << Reaction in bulk a δ k 1 C A,i
if C A, D k 1 2
k L
∞
≈
<<
<<
a k L (C A,i
−
C A,∞ )
0, and δ = D / k L , where D is diffusivity
1
• We define the Hatta number, Ha as:
Ha
=
D k 1 k L
• If the Hatta number is ~1 or greater then we have the “fast” or “instantaneous” regimes and the analysis is more complicated: see reaction engineering textbooks
