plant design_separation_tower design
DESCRIPTION
N9 Sep Tower DesignTRANSCRIPT
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SKF4153: Plant Design 1
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SKF4153: Plant Design 1
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Be able to determine the tower operating conditions (T,P) and the type of condenser
Be able to determine the equilibrium number of stages and reflux required
Be able to select an appropriate contacting method (plates or packing)
Be able to determine the number of actual plates or packing height required, as well as the locations of feed and product
Be able to determine the tower diameter
Be able to determine other factors that may influence tower operation
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SKF4153: Plant Design 1
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Proximity of critical conditions should be avoided
Typical operating P is 1 to 415 psia (29 bar)
For vacuum operation P>5 mmHg
Normally total condenser is used (except for low boiling components and where vapor distillate is desired)
Preliminary material balance to estimate the distillate and bottom product compositions
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Assume cooling water available at 90oF PD : Dist P PB: Bottom P PB=PD+10psia
End
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PD
V
TB
L D
R=L/D
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Applicable when cw (Ti=900F and To=120
0F) can be used.
P at exit of condenser PD (or in the reflux drum) is selected such that to condense stream V to liquid with the cw.
So, this P is bubble P at 1200F.
If PD is < than 215 psia (~15 bar), use total condenser
However, if PD is < 30 psia (~2 bar), use total condenser but reset PD to 30 psia to avoid vacuum operation
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SKF4153: Plant Design 1
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If PD is > than 215 psia, calculate the dew pressure at 1200F. If this pressure is 365 psia, use partial condenser with a refrigerant that give minimum approach T of 5-100F (to replace cw) such that the distillate dew P does not exceed 415 psia (~29bar).
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SKF4153: Plant Design 1
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Bottom pressure (Higher than top pressure) PB= PD + PCond + PColumn= PD + P = PD + (0 to 2 psia) + (5 to 10 psia) = PD + (5 to 12 psia) TB is at bubble point and calculated based on the PB and
bottoms composition.
If the TB is above limiting T (due to decomposition, close to Tcetc), then calculate
PB based on the limiting T, then calculate PD=PB-P and recalculate TD. Check whether this new TD requires different type of coolant and condenser.
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SKF4153: Plant Design 1
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If top stream contain both condensable and non-condensable components, the condenser is designed to produce both vapor distillate and liquid distillate
The PD is calculated at 120oF or lower (if
refrigeration) for the required recovery (composition) of condensable components in liquid distillate.
For vacuum operation, the vapor distillate is sent to vacuum pump.
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SKF4153: Plant Design 1
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If refrigerant is used, always consider placing water-cooled partial condenser ahead of it (to reduce coolant requirement)
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Favor high pressures and low temperatures
Cool the feed gas and the absorbent with cw or refrigerant.
Interstage coolers can be added if there is internal temperature rise.
Due to high compression cost, it might not be economical to increase the feed gas pressure.
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SKF4153: Plant Design 1
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Favor low pressures and high temperatures
Heat the feed liquid and the stripping agent.
Operate at near ambient pressure but not under vacuum.
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This method is valid under the following conditions. For single feed, distillate and bottom products i.e.
ordinary distillation To estimate reflux ratio number of equilibrium stages and feed location.
Quite accurate for ideal mixtures of narrow- boiling range
Not for non-ideal mixtures, azeotropes and mixtures of wide-boiling range (need to use rigorous model)
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SKF4153: Plant Design 1
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Step 1: Using FenskeEqn to determine minimum number of equilibrium stages (i.e. at total reflux, D=0, R=)
d and b are component flowrates at distillate and bottom respectively. HK (heavy key), LK (light key), is relative volatility
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SKF4153: Plant Design 1
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Step 2: Also using FenskeEqn to determine the distribution (d/b) of nonkey component between distillate and bottom streams (at total reflux)
This is a good estimate of d/b at finite reflux
condition.
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SKF4153: Plant Design 1
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Step 3: Using Underwood eqns to determine minimum reflux ratio (Rmin) that correspond to infinite number of equilibrium stages (N=).
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Step 4: Using Gilliland correlation to estimate the actual number of equilibrium stages (N) at a specified ratio of R/Rmin
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Step 5: Estimate the feed location by using Fenske Eqn. (or could use KirkbrideEqn)
A. Calculate NR,minfor rectifying section (between feed and distillate)
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B. Calculate NS,minfor stripping section (between feed
and bottom)
C. We then assume that,
NR,min/NS,min=NR/NS also N=NR+NS
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Example calculation using FUG is provided in Perrys Chemical Engineers Handbook.
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For column with one feed, one absorbent or stripping agent, and two product streams.
To estimate minimum absorbent (Lmin) or stripping agent (Vmin) flow rate and the number of equilibrium stages N.
Instead of relative volatility , this method uses absorption factor (Ae=L/KV) for absorption and stripping factor (Se=KV/L) for stripping.
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Min absorbent molar flow rate
Typical actual absorbent rate L
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To calculate number of equilibrium stages N, use
Where
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Min stripping agent molar flow rate
Typical actual stripping agent rate V
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To calculate number of equilibrium stages N, use
Where
See example 14.1.
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SKF4153: Plant Design 1
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Multistage, multicomponent, VL separation tower (Plate or packed)
Normally done by simulators
Assume equilibrium-stage model (other are mass-transfer models) Mole balance, enthalpy balance and VLE
calculation at each stage
Iterative solution with initial guesses (inside-out method or Newton method)
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From mass-transfer models calculation: We get the actual number of stages (trays) or packed
height.
From equilibrium stage calculation: We get number of equilibrium stages (Nequilibrium) .
We need an estimate of plate efficiency (Eo) to convert Nequilibriumto actual trays (Nactual), or
We need a height equivalent to a theoretical plate (HETP) to convert Nequilibriumto packed height.
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For tray towers,
For packed towers,
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Very approximate plate efficiency: 70% for distillation 50% for stripper 30% for absorber
One method to est. is by Lockett and Leggett version of empirical OConnell correlation
We need a product of average liquid-phase viscosity and average relative volatility
See Figure 14.3.
Another method by Chan and Fair (See Ref.)
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x
Relative volatility?
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Column Height = (Nactual- 1) x (Tray Spacing) + Height of sump below bottom tray + Disengagement height above top
tray.
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Values of HETP are usually derived from experimental data for a particular type and size of packing.
Packing vendors/manufacturers can provide HETP values.
Typical values: For modern random packing: 2 ft
For structured packing: 1 ft
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For modern random packing with low-viscosity liquids, HETP, ft=1.5(Dp, in)
For structured packings at low-to-moderate P and
low-viscosity liquids, HETP, ft=100/a, ft2/ft3 +0.333
For absorption with a viscous liquid, HETP, ft=5 to 6
Where Dp is the nominal diameter of random packings
a (ft2/ft3 ) is the specific surface area of structured packing
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SKF4153: Plant Design 1
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For vacuum service, HETP, ft=1.5(Dp, in.) + 0.50
For high P service with structured packings, HETP, ft>100/a, ft2/ft3 +0.333
For small diameter towers less than 2 ft in
diameter, HETP, ft=tower diameter in feet but not less than
1 ft.
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SKF4153: Plant Design 1
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HTU (Height of a transfer units) and NTU (Number of transfer units)
A more firm theoretical foundations
More accurate
See Transport Process and Unit Operation, Geankoplis
Also Seader and Henley (1998)
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Tower diameter is calculated to avoid flooding (i.e. liquid began to fill the tower and leave with vapor at top)
The diameter depends on,
Vapor and liquid flowrates
Vapor and liquid properties.
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SKF4153: Plant Design 1
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Entrainment flooding, At high vapor flow rate, more
droplets of liquids are carried by the vapor to the tray above and lead to flooding
More common
Downcomer flooding, due to liquid froth in the
downcomer backs up to the tray above.
Not common To avoid, downcomer area should
at least 10-20% of tower cross-sectional area
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SKF4153: Plant Design 1
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There will be liquid holdup in the tower
Under normal operation the amount of liquid holdup is unchanged
If the gas flow is increased, a point is reached where the holdup will increase significantly with increasing vapor flow rate where liquid will begin to fill the tower (flooding).
This is follow by rapid increase in pressure drop
Normally, for a given liquid flow rate, the tower diameter is calculated at 70% of flooding gas flow rate
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SKF4153: Plant Design 1
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Tower inside diameter,
Flooding velocity,
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0.1?
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Tower inside diameter,
For flooding velocity, use
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DT>10(nominal packing diameter) or DT --> 30(nominal packing diameter)
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dPtray tower >dPrandom packing >dPstructured packing
Tray dP can be calculated using simulators dP for tray and packed columns are discussed in
Perrys ChemEngrHbook (1997), Kister (1992), Seader and Henley(1998)
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Reduce downcomer liquid load Reduce tray liquid load Lower tray dP Shorter path length (might reduce tray
efficiency) More expensive Sensitive to maldistribution of liquid and
vapor.
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SKF4153: Plant Design 1
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Sizing of a deisobutanizer Go through this example thoroughly.
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SKF4153: Plant Design 1
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SKF4153: Plant Design 1
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SKF4153: Plant Design 1
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SKF4153: Plant Design 1
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