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ChemEngSquirrel (Chemical) (OP)
17 Mar 11 19:41
I'm trying to determine if the sizing case for an existing PRV is correct.

The PRV is located on the production header of an offshore platform upstream of the HP separator.  A number of subsea wells supply the production flowline.  The subsea line is rated for CITHP (closed in tubing head pressure).  Topside there is a HP/LP interface and is therefore protected by the PRV i'm considering.  In addition to the PRV, two high pressure trips provide further protection by closing an ESDV at the HP/LP interface as well as the sub sea well wing valves.  Note: the two high pressure trips are not considered as HIPPS (i.e. not maintained as SIL 3 protection, ESDV are not fast acting).  There is also a further high pressure trip at the HP sep and this will also isolate ESDVs and wells if activated.

PRV sizing case is blocked outlet based on steady state flow from all producing wells (at maximum expected steady state production).  

However, if a well (initially at CITHP) was inadvertently opened to 100% choke on start up (i.e. operator error), the production rate from that well would exceed both relief valve and flare system capacity.  My question is:

Should relief valve sizing consider a transient peak production flow caused by choke valve moving 100% open with a well at CITHP.    

My initial thoughts are that this would require four failures so may be double jeopardy.  Failures:
1 & 2.  2 high pressure trip failures on production header,
3. operator error in opening choke to 100%
4. Either HP sep high pressure trip failure or blockage upstream of HP sep.

However, for points 1 & 2, it seems possible that the rapid pressurisation caused by the high flow achievable following well opening would prevent the high pressure trip isolating ESDVs / wells quickly enough.  This is because production header trips are only 10% below design pressure; HP sep trip is set 25% below design and is therefore more likely to provide sufficient margin to allow for ESDVs to close before design pressure is breached.  Therefore, i'm not convinced that credit can be given to the two high pressure trip on the production header.

Any advice would be greatly appreciated.
MikeClay (Chemical)
13 May 11 13:33
Restarting is the big problem, especially since your valves cannot close fast enough to prevent the problem.  Items 1, 2 and 4 are not barriers since they cannot stop the problem once the choke is opened.  This leaves a simple operator error causing the problem.  Double jeopardy is not an issue.  The problem needs to be reframed.

What interlocks are in place that prevent restarting?  The systems that I've worked on have one or more shut down valves (SDV) upstream of the choke.  These valves cannot be opened if the flowline pressure exceeds pressure safety high (PSH).  In order to restart, a bypass choke is installed to bleed the flowline pressure to either process equipment or if needed, to a flare. The bypass choke, not the production choke, is one of the sizing cases for the HP Separator PSV.  Once the PSH is cleared, there cannot be a high differenctial pressure (PDSH) across the SDV or else it will not be allowed to open.  In additon, SDV cannot be opened unless the choke is closed.  In all cases, the SDV's, PSH, and PDSH are parts of the instrumented protective system (IPF).  The IPF is evaluated to make sure it meets the proper SIL level.  Then this equipment is routinely tested according to both internal and regulatory specifications to maintain the SIL rating.

Now we are talking about several failures before the SDVs can open:
-  The PSH must fail and
-  The PDSH must fail and
-  The Choke postion indicator must fail.

-  Or the SDV(s) (fail closed) spontaneously opens while the choke is wide open.

TMI?

--Mike--

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