plant design_separation_tower design

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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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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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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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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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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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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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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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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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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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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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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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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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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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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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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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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Sizing of a deisobutanizer Go through this example thoroughly.


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plant design_separation_tower design

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