Possibly to keep an inert atmosphere in the purge piping (to prevent oxygen ingress) or to flush acid gases (corrosion issues) or ensure dry piping....
however Fuel Gas is more common when available (and you won't risk snuffing out you pilot)
The question you should ask is can the surface of the pipe be in the same pool fire as the vessel?
If the line is from the top of a tall tower or in a pipe rack, than it might not be necfessary to include that surface area since it will not really be affected by the fire;
however the bottoms...
Wouldn<t a 2 step process also provide you with a back-up; As the exhaust is usually going to Atm and you are trying to minimise SO2 or HCl venting, having the second step acts as a polishing unit and as a back-up in case the absorption in the 1st step suddenly decreases.
Good to check - but your concern may be misplaced.
A quick calculation (in hysis) shows that relieving to atmosphere you need roughly 30 psi pressure drop in the piping
For the piping this can't be more than 10% of the set pressure (according to API) or else you risk having chattering or other...
As stated by the first post, I would take Normal Operating to be 150 psi
Note also that differential pressure of the control valve in this case is the normal operating pressure at the inlet and the relief pressure plus over pressure of lower system at the outlet.
so, assuming 10% overpressure...
First, have you determined why "valves do not reseat"
as per psafety, you might be operating too close to your set-point. With a traditionnal valve, you will see the valve lift as early as 90% of set-point. You might want to consider going to a pilot which can let you operate at higher...
Yes, I agree that specifying a max Cv implies the sytem is operating at minumum flow, and am comfortable doing so as a no-flow scenario is not really credible.
The only valves between the cooler and the pump suction are the pump maintenance valve which are 18 inches and will not be...
I can not thank you enough for your answers;
Our conclusion is roughly what you both described; By limiting the capacity of the recirc valve (specifying a max Cv) and imposing a minimum pressure drop into the system, we are staying away from any continuous high presure scenario.
We are going...
sorry, it is a pump,
actually, it is 2 pumps in series with a cooler in the middle.
The first pump (small) normally discharges at 5 bar and has a shut-off of 10.
the second pump (big) goes to 90 and has a shut-off of 100.
The recycle of the second pump returns upstream of the cooler (designed...
Thank you all,
I am glad to see that this mirrors the debate we are having here.
I was trying to "simplify" and generalise to what I thought would be the worst case scenario, best described as a "pump test loop" (thank you BigInch).
The REAL system consists of the following:
It is an lean...
I see your point, and agree the safest oute is to have a large PSV for pump flow but is it really necessary:
the acknowledging the possibility of overpressure brings up other issues.
ISSUE 1: If the relief valve is set to open on low pressure piping design pressure of 15 bar, once it opens...
On a centrifugal compressor.
assume normally operating at a suction of 10 bar (constant, very large volume tank) and a discharge of 100 bar. Flowrate of 100m3/h.
Recirculation line for full flow, going from high pressure to low pressure rated piping (note, NOT back to source tank).
There is...
Yes, HAZOP done and flag raised however the feasability of overpressure came up when trying to calculate a relief rate.
Considering it is a liquid, therefore incompressible, if there were to be an overpressure, the relief valve will open and burp some liquid... since the source of the fluid is...
Theoretical question:
On a centrifugal compressor.
assume normally operating at a suction of 10 bar (constant, very large volume tank) and a discharge of 100 bar. Flowrate of 100m3/h.
If I have a blocked outlet and a recirculation line which is grossly oversized (worst case scenario), what...
I have a fairly large (100m3) expansion tank in a closed loop hot oil system.
the cases looked at include:
Gas blow through due to blanketing failure,
Gas blow through due to user failure - tube rupture
Thermal expansion of the liquid (ambient to working pressure)
they are all relatively small...
We are using a Hysis model. Initially based on a Lee-Kessler EOS, and then added extra checks using advanced equations-of-state (as suggested by Aspen). Even the guys at Aspen told me that the phase indicated by the software in the label is incorrect but they reassured me that that the...
Yes, that is what we did: breaking the 33km pipeline into 30m segments, we found that in the winter scenario we just barely dip into the liquid (or dense liquid) part of the graph (somewhere around the 25th km)
As mentionned by Dcasto, the density should be close enough between the 2 fluids, we...
Thank you both for your help, very informative.
To clarify, "at what point does the fluid stop being in dense phase and is considered a liquid"
On the P-H graph, is it :
a) Anthing to the left of the critical point and above the bell (557 kj/kg)?
b) Anthing colder than the critical point (304K)...
Sorry, you are right, Dense Phase.
To paraphrase, even if the pressure remains above the critical point, but the temperature is low enough, the fluid goes from dense phase to a liquid and eventually to a solid.
The melting line is shown.
Which brings up the question of WHEN does it stop being...
