CV Failure Philosophy 4

[Jefka]

Let's say there is a control valve with a system upstream designed for 200 psig (i.e. the upstream relief valve set pressure). Operations plans to actually run the upstream equipment at no more than 150 psig.

There's also a a relief valve downstream of the control valve (system downstream design = 50 psig). When sizing this relief valve for a CV fail open case, what does one assume for the CV upstream pressure in these calcs? Would you use:

a.) 150 psig
b.) 200 psig

Assume the client is indecisive and leaving it up to the designer.

 

[BenThayer]

from an operations viewpoint, any level set-point between high and low level alarms or any pressure set-point between alarms or essentially from vacuum relief and the PSV setting is typically considered to be fair game.

considering the wide range of conditions found during start-up, shut-down, responding to upsets, striving for efficiencies, etc. you essentially have to assume the operators will run it high, wide and handsome.

 

[psafety]

What I would use is (200*1.1)-(50*1.1) = 165 at 100% open Cv for the mass rate capacity for the 50-lb PSV.

 

[chemsac2]

This post has already had replies from many of us. I hope I am not that late in putting my views. Here is what I would do:

I would base my relief rate on maximum operating pressure in upstream vessel i.e. 150 psig in above case. My argument is two-fold:

1) API 521, 2007 edition, section 5.10.3 asks designer to consider operating pressure upstream and relieving pressure downstream for relief calculations for control valve failure

2) I don't buy the argument that operations people try to utilise margin between operating and design to increase throughput/profit margin etc. It is not job of designer to consider these cases. Operations people have to ensure that existing PSV is sufficient for their new operating case.

In example given above, should designer consider the possibility that in future operations department may realise that upstream vessel MAWP is 250 psig and may start operating close to it?

Codes take care of this with margins. Being strict on upstream side and ignorant on downstream side is not how I would design the system.

Regards,

Sachin

 

[pleckner]

chemsac2:

I couldn't disagree with you more. Operations is not responsible for ensuring the PSV is designed correctly for any condition. That is the responsibility of the Process Design Engineer or the unit Process Engineer. Operations cannot just change the operating parameters without checking with engineering if it is safe to do so.

Yes, to boost profits, operations will take whatever they can get, even if it means boosting pressure up to the MAWP of a vessel unless they are stopped by someone or procedure.

Psafety has got it 100% correct.

 

[chemsac2]

Pleckner,

I do not have operations background and hence may be little ignorant about operator behaviour. Consider this post as my way of getting first hand information on operator behaviour.

However, as a designer I would argue against using 200 psig upstream pressure on two counts: technical and administrative.

Technical arguments:

1) When I assume upstream system at design pressure, I assume pressure control system in upstream vessel is not working. Even if one does not consider favourable response of instrumentation system, it is highly unlikely that upstream vessel would reach design pressure due to something that happens downstream.

2) It depends on upstream vessel make-up stream. If make-up is much less than outflow, pressure in upstream vessel would actually diminish.

3) Consider make-up is significant. Now gas blow-by occurs with upstream vessel reaching design pressure. Due to increased pressure in upstream vessel, make-up would decrease (you should not consider negative response of make-up control system).

4) It depends on relative volumes of upstream and downstream vessels as well. In one particular case of ours, low make-up stream flowrate and large volume of low pressure system almost eliminated gas blow-by case.

5) Assume, I have sized PSV considering 150 psig in upstream. Now assume control valve fails during operation and upstream vessel somehow reaches design pressure. It still may not be an issue as following margins may help:
- additional flow is taken care of by rated flow through orifice
- Pressure in downstream vessel may exceed design pressure, but since MAWP is generally higher than design pressure no problem occurs
- Combined flow of PSV and downstream pressure control system (even without favourable response) at relieving pressure may avert the situation

6) Generally PSV for gas blow-by case is calculated using bypass valve full open. When gas blow-by occurs, bypass may not be open providing sufficient margin. Here I assume, administrative controls prohibit operators from keeping bypass open in their bid to increase throughput.

Administrative arguments:

During my limited interactions with site people I found them to be serious about PSV operation. If they think something can lead to PSV blowing, I have seen them avoiding it. Also, system of permits to be obtained from technical departments would also help avoid intentional maloperation.

Summary:

I would consider design pressure of upstream vessel (200 psig) for gas blow-by case only I see a chance of upstream vessel reaching design pressure on control valve failure. E.g. upstream vessel level is controlled by a level control valve in its feed.

I would explore such possibilities during HAZOP

Regards,

Sachin

 

[BenThayer]

think of operators at the plant as drivers on the road. some will drive the speed limit (i.e., operate at design pressure) and others will drive as fast as they think they can get away with and not get a ticket (i.e., operate closer to the MAWP or PSV set pressure).

for a relatively high pressure plant (3000 psig), we upgraded to pilot operated valves so we could run closer to the set pressure. and one operations leader told the PSV company to not bother delivering any PSV's that were on the low side of the tolerance.

a +/- 3% is +/- 90 psig and we wanted to run at 2930 psig. i have also heard tales of PSV's being sent back if they were not on +2-3% side.

 

[Jefka]

@BenThayer: That's a very good point.

Let's say you size the above PSV example for an upstream operational pressure of 150 psig.

Picture yourself walking into this plant and you see a pressure gauge on the equipment reading 155 psig. How comfortable are you standing there? Maybe the gauge reads 160 psig. Now how comfortable are you?

 

[maddocks]

In addition to the above questions, a large part of our work involves going into aging facilities, re-rating them, pushing the limits of flows, pressures, and temperatures. I don't want to leave a facility in the hands of an operator (or a future owner) that will arbitrarily jack the pressure up to 199 Psig with a 200 Psig limit. Having said that, I've heard of instructions to operations teams "push it till the valve simmers and then back off a bit".

 

[chemsac2]

After I posted my reply, I thought myself as an operator. I was expecting operator to be a rational being. But then I know how we use non-rational ways like judgement, rules-of-thumb due to schedule pressure during design stage.

In the end I thought it would be better to be safe than sorry. However, my worry is that assuming operator to be non-rational means thinking of lot many possibilities which I would not consider normally. Probably that is the difference between theory and practice.

Anyway thanks BenThayer, Jefka and Pleckner.

Regards,

Sachin

 

[sheiko]

So be CONSERVATIVE!

"We don't believe things because they are true, things are true because we believe them."